FULL Report - Alternis Monash PDF

MONASH UNIVERSITY
MALAYSIA
IEM Chemical Engineering

Design Competition 2013/2014
Full Report
________________________________________________
Group Members:
Lee Leong Hwee
Jenny Yap Wee Li
Nisha Thavamoney
Lydia Yap Li-Ya
Fatimah Azizah Riyadi
Supervisor:
Dr. Nagasundara Ramakrishnan
Table of Contents
CHAPTER 1 | FEASIBILITY REPORT .......................................................................................................... 1

1.1 Introduction .................................................................................................................................. 1
1.2 Processing Objectives ................................................................................................................... 2
1.3 Feedstock Specification................................................................................................................. 3
1.4 Product Specifications................................................................................................................... 6
1.5 Scope of Design ............................................................................................................................. 7
1.6 Definition of Terminal Points ...................................................................................................... 11
1.7 Plant Availability and Capacity .................................................................................................... 12
1.8 Feedstock Availability ................................................................................................................. 12
1.9 Site Characteristic Constraints .................................................................................................... 13
1.9.1 Local Climatic Conditions ..................................................................................................... 13
1.9.2 Site Characteristics ............................................................................................................... 13
1.9.3 Utilities and Storage ............................................................................................................. 15
1.10 Feedstock Characteristics ......................................................................................................... 16
1.11 Market of Product and Byproduct ............................................................................................ 17
1.11.1 Market of Product .............................................................................................................. 17
CHAPTER 2 | PROCESS FLOW DIAGRAM AND EQUIPMENT SELECTION ............................................... 20
2.1 Process Flow Diagram (PFD) ....................................................................................................... 20
2.2 Equipment Selection ................................................................................................................... 21
2.2.1 Evaluation of Pre-treatment ................................................................................................ 21
2.2.2 Evaluation of Gasifier ........................................................................................................... 26
2.2.3 Evaluation of Post Treatment .............................................................................................. 32
2.2.4 Evaluation of Shift Converter ............................................................................................... 39
2.2.5 Evaluation of Carbon Dioxide Removal Process .................................................................. 44
2.2.6 Evaluation of Methanation .................................................................................................. 50
2.2.7 Evaluation of Ammonia Synthesis Process .......................................................................... 54
2.2.8 Evaluation of Ammonia Separation and Refrigeration Cycle............................................... 60
CHAPTER 3 | MASS AND ENERGY BALANCES ....................................................................................... 65
3.1 Pre-treatment of Biomass ........................................................................................................... 65
3.1.1 Screw Mill (SR-101) .............................................................................................................. 65
3.1.2 Conveyor Belt Dryer (DE-101) .............................................................................................. 65
Page | i
3.2 Dual Fluidized Bed Gasifier (G-101) ............................................................................................ 67
3.2.1 Mass Balance for Gasifier..................................................................................................... 67
3.2.2 Energy Balance across Dual Fluidized Bed Gasifier .............................................................. 71
3.3 Post-Treatment of Syngas ........................................................................................................... 74
3.3.1 Cyclone ................................................................................................................................. 74
3.3.2 Waste Heat Boiler (WHB-101) ............................................................................................. 76
3.3.3 Tar Removal Process ............................................................................................................ 79
3.4 Autothermal Reformer................................................................................................................ 82
3.4.1 Mass Balance of Conversion Reactor ................................................................................... 83
3.4.2 Energy Balance of Conversion Reactor ................................................................................ 84
3.4.3 Mass Balance of Equilibrium Reactor .................................................................................. 85
3.4.4 Energy Balance of Equilibrium Reactor ................................................................................ 86
3.5 Shift Reaction .............................................................................................................................. 88
3.5.1 Mass Balances across High Temperature Water Gas Shift Reactor (HTWGSR) and Low
Temperature Water Gas Shift Reactor (LTWGSR)......................................................................... 88
3.5.2 Energy Balance across High Temperature Water-Gas Shift Reactor (HTWGSR) and Low
Temperature Water-Gas Shift Reactor (LTWGSR) ........................................................................ 97
3.6 Carbon Dioxide (CO2) Removal ................................................................................................... 99
3.6.1 Mass Balance across Carbon Dioxide Removal Section ....................................................... 99
3.6.2 Energy Balance across Carbon Dioxide Removal Section .................................................. 105
3.7 Methanator ............................................................................................................................... 108
3.7.1 Overview of the process and block diagram ...................................................................... 108
3.7.2 Simulating Software and Fluid package ............................................................................. 109
3.7.3 Assumptions ....................................................................................................................... 109
3.7.4 Basis ................................................................................................................................... 109
3.7.5 Steps for conducting mass balance over the entire system .............................................. 109
3.7.6 Energy Balance ................................................................................................................... 114
3.7.7 Comparison: ....................................................................................................................... 116
3.8 Mass and Energy balance: Ammonia Synthesis Section ........................................................... 118
3.8.1 Mass Balance around the ammonia synthesis rector ........................................................ 118
3.8.2 Energy Balance for Ammonia Synthesis Reactor (R-601) .................................................. 123
3.9 Mass and Energy balance: Refrigeration and Separation Section ............................................ 126
3.9.1 Flash calculations across S-701 .......................................................................................... 126
3.9.2 Flash calculations across S-702 .............................................................................................. 130
Page | ii
3.9.3 Refrigeration Loop Mass and Energy Balance ....................................................................... 131
3.9.3.1 Heat Exchanger HX-702 .................................................................................................. 131
3.10 Energy Balance for Common Equipment ................................................................................ 135
3.10.1 Energy Balance across Heat Exchanger ........................................................................... 135
3.10.2 Energy Balance across Compressor ................................................................................. 138
3.10.3 Energy Balance across Centrifugal Pump......................................................................... 141
3.10.4 Mass and Energy Balance across Fired Heater ................................................................ 144
CHAPTER 4| DEMONSTRATION OF SUSTAINABILITY CONCEPT .......................................................... 148
4.1 Environmental Evaluation: LCA Methodology .......................................................................... 148
4.1.1 Goal Definition ................................................................................................................... 148
4.1.2 Inventory Analysis .............................................................................................................. 152
4.1.3 Impact Assessment ............................................................................................................ 154
4.1.4 Interpretation..................................................................................................................... 157
4.2 Process Integration: Heat integration ....................................................................................... 158
4.2.1 Introduction ....................................................................................................................... 158
4.2.2 Heat integration Approach ................................................................................................ 158
4.2.3Aspen Energy analyzer for the Heat integration................................................................. 159
CHAPTER 5 | DETAILED PROCESS AND EQUIPMENT DESIGN ............................................................. 163
5.1 Detail and Mechanical Design: Autothermal Reformer (R-201) ............................................... 163
5.1.1 Definition of Design and Specification ............................................................................... 163
5.1.2 Basis of Performance ......................................................................................................... 165
5.1.3 Sizing of Autothermal Reformer ........................................................................................ 165
5.1.4 Catalytic Bed Specification ................................................................................................. 168
5.1.5 Burner ................................................................................................................................ 169
5.1.6 Mechanical Design ............................................................................................................. 169
5.1.7 Stress Analysis of Autothermal Reformer .......................................................................... 172
5.1.8 Mechanical Design Feasibility Testing of Inner Shell (Refractory Lining) and Outer Shell
(Stainless Steel) of Autothermal Reformer ................................................................................. 172
5.1.9 Mechanical Design of Vessel Support - Skirt .................................................................... 173
5.1.10 Pipe selection and pipe sizing .......................................................................................... 174
5.1.11 Drawing ............................................................................................................................ 175
5.1.12 Datasheet of Autothermal Reformer ............................................................................... 176
5.2 Detailed Process and Mechanical Design of Low Temperature Water-Gas Shift Reactor ....... 179
5.2.1 Definition of Design and Specification for Low Temperature Water-Gas Shift Reactor
(LTWGSR) .................................................................................................................................... 179
Page | iii
5.3 Detailed Process and Mechanical Design: Carbon Dioxide Absorption Column ...................... 191
5.3.4 Mechanical Drawing and Data Sheet ................................................................................. 210
5.4 Detailed Process and Mechanical Design: Methanator ............................................................ 212
5.4.4 General Arrangement Drawing .......................................................................................... 223
5.5 Detailed Process and Mechanical Design: Waste Heat Boiler (WHB-101, WHB-102, WHB-103)
........................................................................................................................................................ 227
5.5.1 Definition of Design and Specifications ............................................................................. 227
5.5.4 Specific Data Sheet and mechanical design drawing ......................................................... 240
5.6 Detailed Process and Mechanical Design: Synthesis Reactor ................................................... 246
5.6.1 Definition of Design and Specifications ............................................................................. 246
5.6.4 Detailed Mechanical Design............................................................................................... 248
5.6.5 Analysis of stresses ............................................................................................................ 251
5.6.6 Sizing of pipe for the inlet and outlet ................................................................................ 253
5.6.7 Specific Data Sheet and mechanical design drawing ......................................................... 254
5.7 Detailed Process and Mechanical Design: Design of Vapour-Liquid Separator (S-702) ........... 257
CHAPTER 6 | PIPING AND INSTRUMENTATION DIAGRAM (P&ID) ..................................................... 269
6.1 Piping & Instrumentation Diagram for Post-Treatment of Syngas Section .............................. 269
6.1.1 P&ID Flow Sheet................................................................................................................. 269
6.1.2 Brief Description of Flow Sheet ......................................................................................... 270
6.2 P&ID (Autothermal Reformer, Syngas and Air Compression) .................................................. 272
6.2.1 P&ID Flow Sheet................................................................................................................. 272
6.2.2 Brief Description of Flow Sheet ......................................................................................... 273
Page | iv
6.3 P&ID (Water-Gas Shift Reactors) .............................................................................................. 275
6.3.1 P&ID Diagram Flow Sheet with Legend ............................................................................. 275
6.3.2 Piping and Instrumentation Diagram (P&ID) Explanation ................................................. 276
6.4 Piping and Instrumentation Diagram (P&ID): Carbon Dioxide Removal Section...................... 278
6.4.1 P&ID Flow Sheet for Carbon Dioxide Removal Section ..................................................... 278
6.4.2 Brief Description of P&ID Flow Sheet for Carbon Dioxide Removal Section ..................... 279
6.5 Piping and Instrumentation Diagram (P&ID): Methanation Section ........................................ 283
6.5.1 P&ID Flow Sheet for Methanation Section ........................................................................ 283
6.5.2 Brief Description of P&ID Flow Sheet for Methanation Section ........................................ 284
6.6 Piping & Instrumentation Diagram of Ammonia Synthesis Reactor Section ............................ 286
6.6.1 P&ID Flow Sheet................................................................................................................. 286
6.6.2 Brief Description of P&ID Flow Sheet ................................................................................ 287
CHAPTER 7 | PROPER DEFINITION OF BASIS, CRITERIA AND LIMITS OF DESIGN ............................... 289
7.1 Definition of Design Basis ......................................................................................................... 289
7.1.1 Functional Goals................................................................................................................. 289
7.1.2 Budgeting ........................................................................................................................... 291
7.1.3 Reliability and Durability .................................................................................................... 291
7.1.4 Flexibility ............................................................................................................................ 295
7.1.5 Maintainability ................................................................................................................... 295
7.1.6 Environmental Evaluation .................................................................................................. 296
7.1.7 Safety ................................................................................................................................. 310
7.1.8 Plant Layout ....................................................................................................................... 353
7.2 Design Limitation ...................................................................................................................... 365
CHAPTER 8 | ECONOMIC PERFORMANCE .......................................................................................... 369
8.1 Introduction ............................................................................................................................. 369
8.2 Market Evaluation of Anhydrous Fertilizer Grade Ammonia .............................................. 369
8.2.1 Current Global Market Size and Demand of Anhydrous Fertilizer Grade Ammonia .. 369
8.2.2 Selling Price Estimation and Forecasting ........................................................................... 371
8.2.3 Main Cost Drivers ............................................................................................................... 371
8.2.4 Product Quality Requirement ............................................................................................ 372
8.2.5 Means of Supply................................................................................................................. 372
8.3 Capital Cost Estimation ........................................................................................................... 372
8.3.1 Key Assumptions and Parameters .................................................................................. 372
8.3.2 Inside Battery Limits (IBL) Investment ............................................................................... 373
Page | v
8.3.3 Outside Battery Limit (OBL) ............................................................................................... 374
8.3.4 Engineering Costs and Contingency Charges ..................................................................... 375
8.3.5 Total Fixed Capital Cost ...................................................................................................... 375
8.3.6 Start-Up Capital .................................................................................................................. 375
8.4 Operating Cost Estimation ........................................................................................................ 376
8.5 Working Capital Estimation ................................................................................................... 378
8.6 Project Profitability Assessment ............................................................................................ 380
8.6.1 Cash Flow Estimation....................................................................................................... 380
8.6.2 Net Present Value (NPV) and Payback time ................................................................... 380
8.6.3 Internal Rate of Return .................................................................................................... 383
8.7 Sensitivity Analysis ................................................................................................................. 383
8.7.1 Product Selling Price ........................................................................................................ 383
8.7.2 OPT Feedstock Purchase Price ........................................................................................ 384
8.7.3 Fixed Capital Cost ............................................................................................................. 386
8.8 Critical overview on Economic Evaluation ........................................................................... 387
CHAPTER 9 | PROJECT VIABILITY ........................................................................................................ 388
9.1 Introduction .............................................................................................................................. 388
9.2 Technical Viability ..................................................................................................................... 388
9.3 Economic Viability ..................................................................................................................... 390
9.4 Environmental Viability and Sustainability ............................................................................... 393
9.5 Strategic aspects affecting the future viability and sustainability of the project ..................... 395
9.5.1 Future growth and demand of fertilizer grade ammonia .................................................. 395
9.5.2 Future trends in technology............................................................................................... 396
9.6 Future Recommendations ........................................................................................................ 396
Reference ............................................................................................................................................ 398
Page | vi
IEM CHEMICAL ENGINEERING DESIGN COMPETITION 2013/2014 1
Full Report
CHAPTER 1 | FEASIBILITY REPORT

1.1 Introduction
Anhydrous ammonia is a chemical base liquefied gas and used for multiple purposes,
including fertilizer production. This product has a sharp odour and is both naturally occurring
in the environment and industrially-manufactured. With 82% Nitrogen of its content, it is the
most cost-effective concentrated nitrogen fertilizer manufacturer upgrade into other nitrogen
fertilizers such as urea and UAN solution. At present, Asia is stated to be the largest ammonia
producing region in the world mainly because of its large and continuously growing
population which requires the utilization of fertilizer to increase the food production to meet
the demand (Rafiqul et al., 2005).
The conventional method of synthesizing ammonia from hydrogen and nitrogen

which is practiced largely in the industry is the Haber-Borsch process. This process is an
energy and resource intensive although the majority of the feedstock costs and energy are
associated with the production of hydrogen (Bartels, 2008). Currently, hydrogen production
is mostly done by using fossil fuels, such as natural gas and coal. However, this method is
neither economically nor environmentally friendly, as both these fuels have a limited supply
and releases considerable amount of greenhouses gases during the production of hydrogen
(Gilbert et al., 2009). Due to both environmental and economic reasons, it is vital enough to
pursue the production of ammonia using alternative renewable and environment friendly
sources in which at the same time do not affect the rate of production of ammonia or the
capacity in the negative direction.
Biomass seems to have been receiving a lot of attention lately not only because it
provides an effective option for the provision of energy services from a technical point of
view but is also based on resources that can be utilized on a sustainable basis all around the
globe (McKendry, 2002). In fact, biomass has been a major source of energy in the world
until before industrialization when fossil fuels become dominant. For example, countries with
extreme conditions found in many poor regions of the world such as Ethiopia and Tanzania
derive more than 90% of their energy from biomass (Kelly-Yong et al., 2007). The
conversion of biomass by gasification into hydrogen rich syngas greatly increases the
potential usefulness of biomass as a renewable resource in ammonia production.
The objective of this project is to design, investigate and propose economic and
technical potential for the production of ammonia using palm biomass as the feedstock.
Full Report
Malaysia being one of the agriculturally rich countries and largest producer of the palm oil in
the world, the vast availability of biomass is undeniable(I. et al., 2005). The ammonia
producing is designed to utilize the Oil Palm Trunk (OPT). The oil palm tree, which bears
fruit at the age of approximately two to three years, has an economic life of approximately
25-30 years, upon which the tree is felled for replanting which contributes to the OPT
feedstock to be gasified into hydrogen-rich syngas which will need to undergo few
purification and filtration steps to remove other components of the syngas such as Carbon
Dioxide, Carbon Monoxide, Aerosols, Tar and sulfurous compound. The hydrogen gas will
then be reacted with the nitrogen gas obtained from the air separation unit to be synthesized
into ammonia.
1.2 Processing Objectives

The main objective of this design team project is to produce a detailed design of
Alternis BioAmmonia plant that is proposed to be built in Langkap, Perak. This plant will be
designed to produce 30 kilo tonnes of anhydrous fertilizer-grade NH3 for a period of 300 days
per annum for 25 years as per to fulfill the local demand of the fertilizer-grade NH3 in Perak.
Syngas production which is an essential part in ammonia production will utilise the
woody wastes from palm oil industry which is plenty in Malaysia. In specific, Alternis
BioAmmonia plant will use the oil palm trunk (OPT) for the syngas production out of few
other wastes produced from palm oil industry such as fronds, empty fruit bunches, palm
pressed fibers, and the shells.
The processing objectives of Alternis BioAmmonia plant are to optimize sources of

oil palm trunk waste produced by the nearby plantation around Langkap area including Benta
Plantation Sdn. Bhd., United Plantation Sdn. Bhd., Southern Perak Plantation Sdn. Bhd., and
FELDA Besout oil palm plantations. In addition, since the greenhouse gases (GHG) emission
is the main concern for any processing plant, Alternis BioAmmonia plant aims to apply
process that reduces the GHG emissions into the atmosphere. The conventional method of
producing ammonia is very carbon intensive and it is approximated that 1.5 kg CO2/kg NH3
is emitted which correspond to 0.93% of global GHG emission (Gilbert and Thornley,
2010).Therefore, by implementing syngas production from biomass gasification rather than
from natural gases, from the amount of GHG emission as mentioned before, this plant aims to
Full Report
have a carbon footprint of less than 0.8kg CO2/kg NH3 from the ammonia production
processes.
Furthermore, Alternis BioAmmonia aims to maximize the ammonia recovery through

intensive evaluation on the process design as to produce high efficiency processes equipment.
In addition, the whole processes will be made as sustainable as possible where any byproduct
produced from each part of the plant that has potential to be used in other application will be
utilised instead of being disposed as a waste. As for example, by using waste heat boiler,
more economical processes could be achieved in the design of Alternis BioAmmonia plant.
Besides, if it has some potential to be used by other industries, this byproduct could be sold
that extra profits could be generated by this plant.
1.3 Feedstock Specification

As mentioned earlier, Alternis BioAmmonia concentrates on the production of
anhydrous fertilizer grade NH3 with a very minimum impact on the environment as well as
utilizing the waste produced from the very prominent palm oil plantation sector in Perak. The
conventional method of producing NH3 uses the Natural Gas which based on the LCA study
proves to be having higher GHG emission of 1.5 kg CO2/kg NH3. However utilization of the
palm feedstock will eventually solve this issue as the GHG Emission from the process
synthesis will be eventually covered by the CO2 being absorbed by oil palm plantation which
is the source of our feedstock and hence leads to a negative CO2 emission.
Oil palm trunk (OPT) feedstock for the Alternis BioAmmonia plant is obtained from
oil palm plantations nearby the site location, Langkap, Perak. Potential supplier of the OPT
feedstock includes Benta Plantation Sdn. Bhd., United Plantation Sdn. Bhd., Southern Perak
Plantation Sdn. Bhd., and FELDA Besout oil palm plantations.
Full Report
OPT is one of the important sources of biomass in Malaysia. The proximate, ultimate
and compositional analysis of oil palm trunk is listed in Table 1.1.1 below (Goh et al., 2010,
Deris et al., 2006)&(Nipattummakul et al., 2012). In order to maintain oil palm productivity
and harvest the oil palm economically, oil palm tree with age 25 years or above will be felled
and replant with new one. In Malaysia, average of 64 million to 80 million old palm trees will
be felled annually, equivalent to 450,000 to 560,000 hectare of oil palm plantation area
(Kosugi et al., 2010). This generates approximately 15.2 million tonnes of OPT annually
(Jung et al., 2011). For the state Perak itself, 148 kilo tonnes of OPT will be generated
annually (Singh, 2013).
Currently, most of the felled OPT are not utilized, the normal practice would be
discarding and burning the trunks at the plantation site which contributes to air pollution.
Only a small percentage of felled OPT are used as feedstock in plywood, pulp and paper
industries because the structure of OTP is not as strong as lumber and it contain high amount
of moisture (Murata et al., 2013). As for the ammonia synthesis, OPT will be a better choice
as the sulphur content in the feedstock is relatively in a trace amount compared to other
components and the post treatment of the syngas can be simplified by removing the
desulphurization process. Therefore, felled OPT has a large potential to serve as a biomass
resources for our production.
Full Report
Table 1.1.1 Proximate, Ultimate and Compositional Analysis of Oil Palm Trunk
Percentage
Analysis Parameter
(% at dry basis)
Proximate Analysis
1. Volatile matter 76.84
2. Fixed Carbon 11.42
3. Ash 5.85
4. Moisture Content 5.89
Ultimate Analysis
1. Carbon 40.64
2. Hydrogen 5.09
3. Nitrogen 2.15
4. Oxygen 52.12
5. Sulphur -
Compositional Analysis
1. Lignin 17.1
2. α-Cellulose 41.2
3. Hemicellulose 34.4
4. Extractives 2.8
5. Ash 3.4
Full Report
1.4 Product Specifications

Anhydrous ammonia is one of the major sources for commercial fertilizers. Fertilizer grade
anhydrous ammonia has the highest analysis of Nitrogen, N. It contains approximately 82
percent of N. It is stored as a liquid under pressure and has a density of 0.62 kg/l at 16oC.
Anhydrous ammonia has a clear appearance and pungent odour. Other specifications of
anhydrous ammonia are listed in Table 1.2 below (CFIndustries, 2013).
Table 1.4.1: Anhydrous Ammonia product specification

Property Min Max
Ammonia wt % 99.5 99.8
Water wt % 0.2 0.5
BL run down temperature (oC) -30 -25
BL run down pressure (kPa g) 200
Boiling point (oC) -33.4
Melting point (oC) -77.7
Flammable limits (by volume in air) 16% to 25%
Autoignition temperature (oC) 651
Solubility in water (per 100g water) 51.0 g at 20oC

Full Report
1.5 Scope of Design

The main purpose of design of this chemical plant is to determine the economic, social,
environmental and technical feasibility of ammonia production in Perak Malaysia. The scope
of this design consists of a few Steps.
 Site selection in which the location of industrial site is determined by considering

factors such as availability of raw material, utility supplies, transportation, etc.
 Technology evaluation which includes assessing and comparing alternate routes for
chemical process and thus selecting the most economical, environmental, efficient and
safe process.
 The process flow sheet is hence developed which comprises of main equipment and
other necessary drawings.
 A series of mass and energy balance is performed for each equipment item on basis of
relevant assumptions.
 Detailed design of equipment is provided along with specification sheet for each item.
 Piping and instrumentation diagram is developed through analyzing and decision

making of control interlocks, maintenance and safety.
 A safety evaluation is conducted describing layers of protection and thus a HAZOP

study is preformed producing a safe and logical layout for the plant
 Environmental evaluation
 Plant layout
 Economic and feasibility study is performed determining capital and operating cost
and thus assessing the profitability of this project.
Finally the viability of the project is discussed. The proposed boundary for manufacture of
ammonia has been shown in the figure 1.5.1.
Full Report
The scope of this design includes feasibility study for chemical process and investigates both
the technical and economic feasibility of the proposed plant. The feasibility study includes
obtaining information about the alternative process routes, and to provide an assessment of
the suitability and sustainability of the project.
The technical part of the feasibility study considers the alternative processes, and the
equipment that constitutes the chemical plant in each part of the plant. At this stage it is
necessary to identify any items of equipment that pose unusual design, or which are very
expensive or hazardous. The feasibility study should determine whether it is economically
and environmentally acceptable to design and build a chemical plant for a particular
manufacturing process (S.Ray and W.Johnston, 1989). Any external factors that may
influence the operation of the plant should be noted, e.g. discharge levels, stability of raw
materials supply, etc.
The environmental aspects of the project must be considered and evaluated. This involves
treatment of unwanted chemicals (by-products) and reducing the concentrations of liquid
discharges and gaseous emissions during normal operation and also when handling a major
chemical accident, with any subsequent reaction products, containment and clean up.
Hydrogen being an important part of ammonia production is mostly produced using fossil
fuels, such as natural gas and coal. However, both of these fuels have a limited supply, and
they release greenhouse gasses during the production stage of hydrogen. Therefore, for both
environmental and economic reasons, alternative energy sources such as biomass feedstock
must be pursued for the purposes of producing hydrogen in an ammonia economy. Most of
the alternative technologies are still more costly than fossil fuel energy sources, but the
relative cost of alternative fuels is decreasing through technological improvements and
increasing fossil fuel costs requires us to look into the future (Anon, 2000).
By implementing the manufacture of ammonia from biomass there can be a reduction in

global warming potential. Hydrogen rich feedstock from biomass gasification will reduce the
impact of CO2 emissions by minimizing natural gas input. Switching to a H2 rich feedstock
from biomass gasification will reduce the impact of CO2 emissions by minimizing natural gas
input. Therefore, the resultant GWP will be lower from ammonia synthesis, thus increasing
the environmental viability of the process (C.W.Ritz et al., 2004). Sustainability aims at high
material yield by the minimization of by product and waste. The same is valid for energy, for
which considerable saving may be achieved by the heat recovery and steam generation.
Full Report
The proposed project would have minimal effect on the health of either the environment or
local residents during construction and operation, through the implementation of mitigation
measures. The site chosen does not include residential areas within a radius of 5 km, hence
not affecting the lives of people. Therefore the environmental integrity of the site will not be
reduced as a result of the proposed project.
The economic evaluation of ammonia plant must be conducted at feasibility study stage in
order to determine the viability of the plant by assessing if the plant can sustain its own
expenses. This estimation is conducted by considering the fixed capital and the operating
expenditure, interests, tax and insurance and finally assessing the profitability, payback
period and return on investment of ammonia plant (Bartels and Pate, 2008). Excess steam and
electricity will be supplied to nearby industrial sites and by products are to be sold, which
adds to profit. By implementing the most effective and efficient technology as well as
proposing heat, water and energy integration hence achieving an optimized plant there could
be a major reduction in the operating costs of the plant. Hence increasing the profit margin of
the chemical plant and therefore obtaining economic viability.
The proposed project would provide social and economic benefits to the community through
local employment opportunities and by creating export opportunities. The two year
construction phase is expected to require a construction workforce and this provides long
term employment to those personnel providing services such as maintenance, transport and
support services.
IEM CHEMICAL ENGINEERING DESIGN COMPETITION 2013/2014
Full Report 10
Figure 1.5.1: Process Block Diagram indicating Inlet and Outlet

Full Report
1.6 Definition of Terminal Points

Table 1.6.1: Terminal Points of process boundary
Stream Description Flow Method of
direction transport
Input
Palm oil Trunks are chipped and dried to the desired size and IN Truck
trunk moisture content and are used as biomass feedstock for
the process.
Air Air compressor sends the pre filter air to the combustion IN Pipe line
chamber of the gasifier, Auto-thermal Reformer,
Catalyst Catalyst required for water gas shift and ammonia IN Manual
synthesis. handling
aMDEA Solvent to recover CO2 from synthetic gas IN Manual
Handling
Utility
Steam Steam Steam Steam
Electricity Electricity Electricit Electricity
y
Cooling water Cooling water Cooling Cooling
water water
Refrigerant Refrigerant R-717 Refrigera Refrigerant
R-717 nt R-717 R-717
Product and by products
Ammonia Produced and stored in refrigerated form, supplied as OUT Pipe
(main fertilizer.
product)
Ash and char Produced as a result of gasification and syngas cleaning OUT Pipe
(by product) can be utilized by mixing it with ammonia to use as
fertilizer.
Carbon Generated during the process and is sent for storage OUT Pipe
dioxide (by before transporting to glycol plant.
product)
Flue gas Flue gas released from the combustion chamber of the OUT Pipe
gasifier as well as from the combustion reaction of the
auto-thermal reformer
Purge gases Produced upon ammonia synthesis. OUT Pipe
Effluent
Wastewater Produced during the cleaning of bio syngas and can be OUT Pipe
further treated by transferring it to waste water
treatment.
Full Report
1.7 Plant Availability and Capacity

The life span of the Alternis BioAmmonia Plant that to be designed is targeted to be for
25 years. This ammonia producing plant will remain in operation 24 hours per day
continuously for 365 days per annum with planned shutdowns for maintenance resulting in an
approximate average of 300 operating days per annum. Short-term plant shutdown will be
scheduled on regular basis mainly for the general maintenance, cleaning and equipment
substitution. At the same time, the long-term shutdown of the plant will be conducted for 4
weeks for every 6 months operating period in order to inspect, repair, and replace the plant
units whose reliability may fall off. The timing of the shutdown is set to coincide largely with
the feedstock availability. Contemplating all these factors, the overall operational availability
of the plant is 82%.
The Alternis BioAmmonia plant has the capacity to produce 30 kilo tonnes of
anhydrous fertilizer grade per year based on the plant availability as explained above.
This capacity will require a minimum OPT feedstock of approximately 65 kilo tonnes on
annual basis.
1.8 Feedstock Availability

As mentioned above, the plant will be producing 30 kilo tonnes of ammonia per year and
based on the stoichiometric equation the minimum amount hydrogen gas supply that need to
be supplied from the gasification of the OPT will be approximately 5.3 kilo tonnes of
hydrogen gas per annum. Based on a lab scale experiment conducted by Nipattummakul and
Ahmed in the year 2011 showed 35 g of the dry oil palm trunk (OPT) eventually results with
3 g hydrogen gas yield though gasification process (Nipattummakul et al., 2012). Therefore,
it is estimated that annual input of 65 kilo tonnes of OPT will allow to meet the targeted plant
capacity.
According to the Malaysian Palm Oil Board (MPOB), there are approximately
379,946 hectares of oil palm plantation with 89% matured plantation in the state of Perak
(Division, 2012). Benta Plantation Sdn. Bhd., United Plantation Sdn. Bhd., Southern Perak
Plantation Sdn. Bhd., and FELDA Besout oil palm plantations are few of the plantations that
can be named to be located closer to the plant site selected. Nearly 13% of the total area of
the oil palm plantation in Perak will be replanted every year which contributes the old oil
palm trunks that had been felled off during this process (I. et al., 2005). It is estimated that the
Full Report
average of 3 tonnes of dry OPT is obtained per hectare of oil palm plantation (Singh, 2013).
This accounts for the OPT feedstock availability of 148 kilo tonnes per year. This provides
high levels of confidence in the life of the plant as the feedstock availability coincides with
the minimum feedstock requirement on annual basis.
1.9 Site Characteristic Constraints

1.9.1 Local Climatic Conditions
Perak, Malaysia has tropical rainforest climate that is hot, humid and does not
pronounce any dry season, summer or winter throughout the year. The average precipitation
of all months has at least 60mm and the daily temperatures are fall around the range of 21°C
to 32°C (Richmond and Karlin, 2010). The humidity levels of Perak hover around 70% to
90% and the rainfall is usually high and distributed throughout the year with June and July as
the driest month. The site location, Hilir Perak, Perak has a total rainfall of 1500 mm to 1700
mm (Chandrawathani et al., 2013).
1.9.2 Site Characteristics

The major constraint in finding a viable plant site is the availability of raw material,
oil palm trunk (OPT). The accessibility of OPT is very limited as OPT is obtained during
replantation and it is done based on the maturity of the oil palm trunk. A normal oil palm tree
is usually passed their economic age on an average after 25 years and lead to replantation.
Therefore, a location with adequate amount of OPT supplier, convenient transportation
network and large flat land will be the main consideration in site selection. Based on
research, oil palm plantations are mainly distributed in the southern part of Perak such as,
Langkap, Sungkai, Hutan Melintang as well as Sitiawan.
Considering all possible constraints, Langkap is chosen as the ideal plant site as it
satisfied the constraints mentioned. The location of the proposed plant site is shown in Figure
1. The oil palm plantations that are closer to the plant site are Benta Plantation Sdn. Bhd.,
United Plantation Sdn. Bhd., Southern Perak Plantation Sdn. Bhd., and FELDA Besout oil
palm plantations. Furthermore, Lebuh Raya Utara Selatan (PLUS) is 32 km away from the
plant site. Therefore, the issues of accessing raw material in Perak state and transporting raw
materials and product are no longer a concern. Besides, that land cost is another constraint
that will affect the capital cost and return of investment of the company. So, land with
reasonable price complement with market value and strategic location will be the best option.
Full Report
In this case, the selected site does not subject to heavy flood also it is sufficient to occupy the
whole plant and reserved for future expansion. Apart from that, the river near to the plant site
awarded bonus mark to this plant as Ammonia plant is one of the industries that consuming
enormous amount of water. As a result, the site is generally minimize the cost, distance and
time for raw material transportation as well as reduce the utilities cost.
Moreover, the proposed plant site is at an appropriate location which is located 12 km

away from the Teluk Intan where most of the residential area is so the operation of the plant
would not affect the neighbourhood. Not only that, the expenses of the plant will be reduced
by hiring the labour force from the nearby area. Other factors that satisfy the plant selection
requirement are electricity supply, infrastructure and fire protection are available in the town,
Teluk Intan.
Figure 1.9.2.1: Site Location of Alternis BioAmmonia at Langkap, Perak (Google Earth, 2013)
Full Report
Figure 1.9.2.2: Oil Palm Plantations around Proposed Plant Site (Google Earth, 2013)
1.9.3 Utilities and Storage

The table below summarizes the condition specification and cost of supply of the
utilities required for the daily operation of Alternis BioAmmonia Plant, which includes
electricity supply from Tenaga National Berhad and water supply obtained from river located
5km away from the plant while natural gas, compressed oxygen and saturated steam are
purchased from nearby plants.
Table 1.10.1 Condition Specification and Cost of Supply for Utilities

Utilities Condition Specification Cost of Supply
Electricity Distribution voltages: 33kV, 11kV, 22 RM 0.288/kWh
kV, 6.6 kV and 400/230 volts
Supply Frequency: 50Hz ±1%
Earthing System: 3 phase configuration
Natural Gas Lower Heating Value: 34.6 MJ/m3 RM 600/t
Pressure: 30 bar
Cooling Water Nil RM 1.61/m3
Oxygen Condition: Dry RM 25/t
Pressure: 30 bar
Saturated Steam Pressure: 30 bar RM 100/t
Hot Water Temperature: 90 RM 17.5/t
Full Report
The storage of biomass feedstock is often necessary due to its availability versus the
need to maintain the continuous production of the anhydrous ammonia product. Biological
activity during storage can cause variable physical and chemical changes in feedstock
properties. Therefore, to maintain the feedstock quality, the biomass is to be stored in an
enclosed structure with gravel or crushed rock floor.
Liquefied anhydrous ammonia produced is usually stored in 3 different types of tanks

which are fully refrigerated or semi-refrigerated tanks at atmospheric pressure or high
pressure non-refrigerated tanks (Lele, 2008). The ammonia product in Alternis BioAmmonia
plant is stored using the fully refrigerated storage tank at atmospheric pressure and at -33
taking safety and cost into considerations. The ammonia storage terminal is located away
from the main production plants due to various safety aspects and considerations for the
condition of the storage tanks and connected operations.
The production of ammonia typically releases 1.5 – 3.0 tCO2/t of ammonia (ETSAP,
2010) depending on various aspects such as type of feedstock and the overall production
process. The carbon dioxide released during the production of ammonia is captured and
stored. The CO2 produced as a side product can be later sold to other industries.
1.10 Feedstock Characteristics

The feedstock oil palm trunk (OPT) contains 40.64% carbon which can be utilized to
produce hydrogen through gasification process(Deris et al., 2006). Hydrogen is an essential
component for the production of anhydrous ammonia. In order to meet the product
requirement, operating specifications and maximize the performance of the plant, the
feedstock needs to be processed or treated before entering the gasifier. OPT contains 45%
moisture which is too high for the operation within the gasifier. Drying process needs to be
carried out to reduce the moisture content to below 10%. Besides that, size of OPT also needs
to be reduce to 50 – 100 mm to increase the surface area to feed ratio.
Full Report
1.11 Market of Product and Byproduct

1.11.1 Market of Product
1.11.1.1 Current Demand

Ammonia is one of the most highly produced chemicals in the manufacturing industry
and its consumption is driven primarily by the production of downstream fertilizer products,
such as urea, ammonium nitrates, ammonium phosphates and ammonium sulphates. About
80% of ammonia produced is used in the agricultural industry. It also forms a main reactant
in Ostwald process which is the manufacturing of nitric acid.
Global demand for ammonia is the highest in Asia, with China and India accounting
for the majority of global demand. In the developed regions such as North America and
Europe, the demand has largely stabilized where as large populations and growing economies
in countries such as China and have substantial consumption potential, which is reflected in
the high growth of ammonia downstream segments such as urea, ammonium nitrate,
ammonium sulphate and phosphate. The Asia-Pacific region accounted for a 58.7% share of
global demand for ammonia in 2011, with China and India accounting for the majority
(PotashCorp, 2011). As a result, ammonia demand from the Asia-Pacific region will continue
to drive global demand in future. Global demand for ammonia stood at 96,437,749 tons in
2000 and is expected to reach 160,093,693 tons in 2020 (PotashCorp, 2011).
Agricultural has played a pivot role in the development of Malaysia as well as in the
development of national economy. Malaysian Government has committed to promote and
maintain agriculture as the third engine of growth of the national economy, thus the usage of
fertilizer under the agriculture is trended upward. In this case, due to the projected increases
in the expansion cultivated areas and fertilizer is the highest in variable costs in crop
production budget, the availability of fertilizer must be emphasized to sustain the growth of
crops. However, the majority of fertilizers used in Malaysia are mainly imported from
countries such as Indonesia, China and Thailand(Sabri, 2009). Therefore, pragmatic solution
is proposed to improve the efficiency in the fertilizer industry and minimize the fertilizer
price. Alternis BioAmmonia was committed to design an anhydrous ammonia plant by
utilizing oil palm trunk biomass as the main feedstock. The designed plant capacity is 30kT
per year which is targeted on the local demand in Langkap, Perak especially the palm oil
plantation nearby. The proposed production plant was accounted for future expansion to as
Perak has large planted area of oil plant plantation.
Full Report
Figure1.11.1.1.1: Global ammonia consumption 2011 Figure 1.11.1.1.2: World consumption of ammonia 2010
(Potash Corp, 2012) (Potash Corp, 2012)
1.11.1.2 Future Trend

Figure 1.11.1.1.1 shows increasing ammonia consumption in various countries and Figure
1.11.1.1.2 represents a predicted global increase in ammonia worldwide till the year
2015.World consumption of ammonia increased by 2.3% annually from 2005–2010, although
it slowed down during the year 2008-2009. Growth is forecast at 2.7% annually from 2010–
2015 and the current ammonia demand is now at 160 million tons as shown in Figure 3 and a
similar trend is expected to continue over the forecast period.
1.11.1.3 Selling Price

The selling price of anhydrous ammonia increased up to RM2600/ton in recent years. The
figure below shows varying selling price trends over the years.
Figure 2.2.3: Ammonia price trends from 1988-2012
1.11.2 Market of Byproduct

Carbon dioxide has been known to be one of the major greenhouse gas emitted into
the atmosphere. Kyoto agreement stated that CO2 emissions are needed to be reduced by
15% (Svendsen, 1998). Initiatives made to reduce the CO2 emissions was by focusing on the
design and to operate CO2 markets (Veal and Mouzas, 2012). The market and demand for
carbon dioxide continue to increase in many developing and developed countries in the
application for food chilling and freezing, pH reduction, agriculture applications and
Full Report
utilization of CO2 in algae and other forms of corps in biodiesel production, for example
(Rushing, 2010). In the past couple of years, the European Union Allowance (EUA) price
(the current reference price in the carbon market) is between €15-€20 (RM65-RM87) per
tonne CO2. Market analysts expect the prices to increase up to €25 (RM108) globally
somewhat during Phase III (2013) of the European Union Emission Trading Scheme (EU
ETS) (E&Y, 2012). The Carbon Finance at World Bank describes a grew in carbon market
by a total of 11% year of year (yoy) in 2011, where the demand for carbon dioxide in the
industry is expected to continue to rise in both developing and developed countries due to its
vast applications in the industry (Bank, 2012).
Full Report
CHAPTER 2 | PROCESS FLOW DIAGRAM AND EQUIPMENT SELECTION

2.1 Process Flow Diagram (PFD)
Full Report
2.2 Equipment Selection
2.2.1 Evaluation of Pre-treatment
2.2.1.1 Technology Evaluation

The waste oil palm trunks obtained during the replantation of oil palm trees is used as the
feedstock for the production of anhydrous ammonia in this proposed design of plant. The
feedstock will be mainly used for the production of syngas as a fuel in a gasifier in which the
hydrogen component of the syngas will be extracted and will be introduced into the ammonia
synthesis reaction. In order to improve the performance of the biomass feedstock as a fuel in
the gasification process, it is highly necessary to pretreat/process them according to the type
of the gasifier chosen as different type of gasifier has different fuel requirements that need to
be fulfilled. Besides, the degree to which the specific pretreatment process is required also
depends on the gasifier plant. The two most important and relevant feedstock properties that
need to be considered is the size and the moisture content as well as the ash fusion
characteristics(Bronson et al., 2012).
In this design, the gasifier that will be installed is the Fast Internal Circulating dual-
Fluidized Bed, by which the gasifier chamber is based on a Bubbling Fluidized Bed. This
type of gasifier is proved to be more tolerant towards feedstock size and fluctuation in feed
quantity and moisture compared to the other type of gasifier (Chiang et al., 2012). The
maximum size of the feedstock particles that can be accepted by the BFB gasifier is 50 to 150
mm accompanied by the optimal moisture content of 10-15%(E4Tech, 2009). Since gasifier
does not have a specific chemical properties requirement of the feedstock, the pretreatment
process is focused on physical pretreatments such as Sizing and Drying.
2.2.1.2 Sizing
Since the feedstock is delivered directly from the plantation area to the plant, the oil palm
trunks with bole length of 7 m to 13 m, with a diameter of 45 cm to 65 cm, measured at
breast height need to be chipped or shredded into 50 to 100 mm sized fibers(BFPIC, 2009).
Smaller fuel particle size will eventually increase the surface area to feed-rate ratio and thus
resulting in higher rate of gasification process(Bronson et al., 2012). Currently, there are
various kind of size reduction equipment that are available in the market and they are
normally classified according to the method they are employed to process the waste. Four
different type of size reduction equipment which was considered are Hammer mill, Screw
Full Report
Mill and Ball Mill. Hammer mill consists of rotating sets of swinging steel hammers that can
be either fixed or flexible ‘flap’ hammers. The maximum particle size output is often varied
by using different number of hammers(Laurence and Ashenafi, 2012a). Screw mill
meanwhile involves the action of 2 high-level screws that will draw the feedstock into the
mill and force it down to a lower spinning roller. Geometry of the cutter can be varied
according to the required particle size output(Banks et al., 2010). The ball mill also known as
cascade mill consists of a slow running rotary drum with a diameter of 4 - 7 m, where 17% of
the volume is filled with steel balls that will crush the feedstock input due to the relative
motion between the steel balls and the input(Banks et al., 2010). Table 2.2.1.2.1 lists down
the advantages and disadvantages between the mills (Banks et al., 2010, Knoef, 2010).
Table 2.2.1.2.1: Comparison between different types of millers for feed size reduction (Banks et al., 2010, Knoef, 2010)
Hammer Mill Screw Mill Ball Mill
 High degree of
 High throughput shredding
 Low wear and tear
rates  Low dust
Advantages  Low noise emission
 High degree of emissions
 Low dust emissions
shredding achieved  Small space
requirement
 Low throughput  High energy
 High wear and tear
rates demand
Disadvantages  Noise emissions
 Labor requirements  Low throughput
 Dust emissions
and maintenance rates
Diameter of
particle 80-100 mm 50-80 mm 20-40 mm
output
Waste with wide
Brittle, high density Waste with wide range
range of brittleness,
Suitable for: waste easily split or of brittleness, density
density and physical
broken and physical durability
durability
Higher than Screw
Cost Lower than Ball Mill Excessively High
Mill
Power  High energy
 High energy demand  Low energy demand
Consumption demand
Full Report
2.2.1.3 Drying
Feedstock OPTs that to be delivered to the plant is consisting of moisture of 45% as
mentioned in the previous sections. For a thermal conversion of biomass via gasification, it is
not efficient to utilize a feedstock with 30% moisture content as most of the energy supplied
to the process will be used to evaporate the water content(Bronson et al., 2012).
Consequently, the higher content of steam will affect the composition of the syngas which
may result in low hydrogen percentage. Studies have shown that using feedstock with higher
moisture content results in production of more tar in the syngas due to the large temperature
drop during the process(Roos, 2008). Therefore, removal of moisture via drying from the
feedstock to a level of 10% is significantly important and there is few drying equipment that
can be implemented. The 3 types of biomass dryers that were considered are Rotary Dryers,
Conveyor Dryers and Flash Dryers.
Rotary dryers are the most popular choice in the industry which consists of a
peripheral flights fitted slightly inclined rotating cylinder to lift, distribute and transport the
material during the drying process(Worley, 2011). Hot air or gas will be streamed to come in
contact with the feedstock in the rotating drum to promote the evaporation of the moisture.
Flash Dryer meanwhile, is capable of drying the biomass rapidly as in a matter of seconds
due to the easy removal of moisture as the required diffusion to the surface occurs readily(Li
et al., 2010).For belt dryer on the other hand, the feedstock is spread on a moving perforated
conveyor to dry the material in a continuous process(Li et al., 2010).
Full Report
Table 2.2.1.3.1: Comparisons of properties of Dryer Types (Roos, 2008, Worley, 2011)
Dryer Type Rotary Conveyor/ Belt Flash
Fines may need to be
Feedstock Less sensitive to Requires small
screened out first and
Requirement particle size particle size
added back
Temperature (oC) 200-600 150-280 30-200
Moisture
Discharge 10-45 10-45 15-25
(10-45%)
Capacity 3-45 4.4-16 No limits
Comparable to rotary
dryer, but may require
Capital & Higher than
less ancillary equipment
Operating Cost rotary dryers
for treatment of emissions
reducing overall cost
Operation and Subject to
Maintenance Low Greater that Rotary Dryer corrosion and
Requirements erosion
More VOC
emissions
Environmental Lower emissions of VOCs
compared to No emissions
Emissions and particulates
lower temperature
dryers
Less opportunity High opportunity for heat Heat Recovery is
Energy Efficiency
to recover waste recovery due to lower Difficult and high
& Heat Recovery
heat temperature blower cost
Larger than comparably-
sized rotary dryer. Multi- Smaller footprint
Footprint - pass conveyors save space than rotary and
and can have comparable conveyor dryers
footprint to rotary dryer
Greater than
Fire Hazard lower temperature Low Medium
dryers
Full Report
2.2.1.4 Process Selection

Based on the overview of the advantages and disadvantages listed above in Table 2.2.1.2.1,
the option of screw mill seems to be a perfect choice for the design. Particle output with
diameter of 50 to 80 mm falls between the range of diameter of feedstock particle required
for the gasifier and thus the energy and cost intensive screening process can be omitted from
the pretreatment process as the maximum particle size is 80mm which is still able to be fed
into the gasifier. Economic wise, screw mill has more advantages over the other two mills
despite producing smaller diameter particle, it is still a low energy consuming device and the
maintenance cost is also reduced as it has a lower tendency to wear and tear. The purchase
cost is also relatively low compared to the other mills. Considering safety and environmental
issues, screw mill is the best option as the dust emission is relatively low and less noise is
emitted as well.
Based on the comparison done between the 3 types of dryers in Table 2.2.1.3.1, the
Conveyor dryer, also known as belt dryer will be a wiser choice to implant in the
pretreatment process. This is mainly because it operates under low temperature compared to
the rotary dryers by which happened to remove the same amount of moisture, and thus
reducing the power usage as well the operating cost. Furthermore, lower operating
temperature eventually will reduce the fire hazards and also enables it to utilize the heat from
waste heat recovered from exhaust of process heating in other facilities such as the flue gas
from combustor chamber of the gasifier, the flue gas obtained from the methanation process,
as well as the ammonia synthesis reactor(Li et al., 2010). Since the flue or exhaust gas
leaving these facilities is warm, no additional energy need to be supplied for heating the
recycled exhaust gas. When the exhaust gas is passed through, the heat exchanger it will
transfer heat to the inlet air into the dryer resulting in moisture removal(Roos, 2008). Studies
also shows that the emission of VOCs from the belt dryer is relatively lower compared to that
of the rotary and flash dryer due to the fact that lower temperature operation is applicable. In
the economical aspect, the belt dryer does not require expenses for the treatment of the
emission compared to the rotary and flash dryer. However, the operation and maintenance
cost for belt dryer is comparatively higher than the others as it is not a single-pass dryer and
often multi-pass conveyor is required(Worley, 2011).
Full Report
2.2.2 Evaluation of Gasifier

The current thermo-conversion technology for production of syngas from the biomass
includes pyrolysis, gasification, and the conventional combustion of biomass. Among these
technologies, biomass gasification has garnered much attention from the researchers as well
as the industry mainly due to its higher efficiency. Besides, combustion of biomass is a
process that will lead to hazardous emissions making the process to be less sustainable,
meanwhile flash pyrolysis is still under development and the efficiency is not proven under
large scale(Maniatis, 2001).
Gasification is the partial oxidation of the carbonaceous fuel or the biomass feedstock
at high temperature ranging from 800 to 1000oC in which results in the production of
syngas(Kaushal and Tyagi, 2012). The syngas mainly consists of a mixture of primarily
hydrogen, carbon monoxide, carbon dioxide, and methane.In the gasifier unit, the biomass
fed will be degraded thermally in 2 process which is drying followed by the devolatilisation
at temperature ranging from 100 to 500oC(Göransson et al., 2011). The devolatilisation
process is endothermic and it is the most decisive step as it produces 75-90% volatile material
in the form of gaseous and liquid hydrocarbons. The kinetic of this stage highly depends on
the temperature, particle size, feed residence time, biomass composition and heating
rate(Kaushal and Tyagi, 2012). Therefore, it is very vital to ensure that appropriate gasifier
technology is chosen based on the type of the biomass and property of the syngas required.
The operating conditions of the gasifier should be given high consideration as well as the
pretreatment of the feedstock. The thermal degradation then followed by oxidation reaction
of the char produced that will generate combustible gas rich in carbon monoxide and
hydrogen. The oxidizing agents that are commonly used in the industrial application of
gasifier are air, steam, oxygen, mixture of oxygen and steam.
Gasifiers could be classified on the basis of few categories such as the gasifying agents, the
operating pressure, operating temperature, the fluid dynamics and in terms of the heat supply.
Table 2.2.2.1.1 on the following page shows the classification of biomass fired
gasifiers(Kaushal and Tyagi, 2012).
Full Report
Table 2.2.2.1.1: Classification of Gasifiers according to few common categories

Gasifying
Pressure Design Heat Supply
agents
Air Atmospheric Updraft Autothermal
Oxygen Pressurized Downdraft Allothermal
Steam Fluidized Bed
Circulating fluidized
Carbon Dioxide
Bed
Entrained Bed
1) Gasifying Agents
Biomass can be gasified using different gasifying agents,depending highly on the
desired product gas composition and energy consideration as well as the availability of the
agent for the plant. Using air as gasifying media poses a risk of producing syngas with
inferior quality since nitrogen composition of the syngas will be very high and thus
eventually reducing the hydrogen content(Foscolo, 1997). Therefore, air as a gasifying agent
will not be a good choice for this design as we acquire hydrogen rich syngas to be used in the
ammonia synthesis. Despite producing syngas with superior quality, using oxygen as
gasifying agent will impose additional cost for oxygen production(Chen and He, 2011).
Steam gasification seems to be a perfect choice for the design of this plant as it will produce
syngas relatively rich in hydrogen content and nitrogen free. Besides that, the presence of
steam will allow the product gas to be catalytically upgraded resulting in lower production of
tar and char(Inayat et al., 2010).
2) Pressure
Gasifiers could operate under atmospheric pressure or in a pressurized condition.
Each case has its own advantages and disadvantages. Pressurized gasifier will produce syngas
in smaller volume that will be sent for syngas cleaning whereby will reduce the cost and
energy requires(Göransson et al., 2011). Besides, most of the downstream facility for syngas
cleaning operates at high pressure and thus eliminates the cost and energy to compress the
syngas produced. However, the capital and operational cost for the pressurized gasification
will be higher and at the same time, the biomass may be difficult to be fed into the gasifier
under high pressure(Göransson et al., 2011). Under atmospheric pressure, such problems will
not be faced and the pressure balance within the gasifier can be maintained easily. Therefore,
the gasifier that to be used in the plant is to beset to be at atmospheric pressure.
Full Report
3) Heat Supply
Gasifiers can be classified in terms of heat supply technology by 2 means,
autothermal and allothermal this classification actually relies on the type of gasifying agents
that to be chose. Autothermal processes generate heat that will be utilized to sustain the
reactor at the optimum reaction temperature (exothermic) meanwhile allothermal gasifiers
requires heat to be generated outside the gasifier and transferred inside(Kaushal and Tyagi,
2012). Air gasification is highly exothermic reaction and thus falls under autothermal process.
On the other hand, steam gasification is highly endothermic and thus eventually falls under
allothermal process. For the design of the plant, allothermal process was chosen as it will
result in higher hydrogen content and the heat is to the gasifier is to be supplied by circulating
the hot bed between the gasification and combustion zone. Figure below depicts the transfer
of heat and mass within the gasifier:
Figure 2.2.2.1.1: Allothermal means of heat supplky through the circulation of bed in DFBG (Schmid et al., 2012)
4) Design
The design of the gasifier that to be implemented in the plant will eventually depend
on the type of feedstock as well as the factors mentioned above, gasifying agent, pressure and
in terms of heat supply. Generally, gasifiers can be classified into 2 major designs, fluidized
and fixed bed gasifiers by which the fluidized bed gasifier can be further divided into
circulating and bubbling bed gasifier and fixed bed gasifier can be divided into updraft and
downdraft gasifier. Comparison between fixed and fluidized bed gasifiers was done to reduce
the number of technologies that need to be considered. Table 2.2.2.1.2 on the following page
shows the comparison and it can be concluded that fluidized bed gasifiers will be a better
option for the plant(Warnecke, 2000).
Full Report
Table 2.2.2.1.2: Comparison Between Fixed and Fluidized Bed Gasifier (Warnecke, 2000)
Reactor Type
Criteria
Fixed Bed Gasifier Fluidized Bed Gasifier
Simple and Robust
Complexity Less complex technology
Construction
Temperature
Bad temperature distribution Good temperature distribution
Distribution
Heat Exchange Poor heat exchange Very good heat exchange
Conflicting temperature
Possible ash agglomeration requirements exists for low-
Ash
and clinker formation on grate reactivity feedstock with low-
softening ash melting point
Gas-Solid(Biomass) Good gas-solid contact and
Channeling is possible
Mixing mixing
Residence time for solids hours to days seconds to minutes
Residence time for gas seconds seconds
Pressure drop Low High
Very limited scale-up potential
Scale up potential Very good scale-up potential
caused by low maximum size
Startup/shut down Long period to heat up Easily started and stopped
Requirement of High ash content feedstock is Tolerates wide variations in
pretreatment possible fuel quality
Updraft: Product gas contains
tar, oil, phenols Amount of tar and phenols in
Quality of syngas
Downdraft: Amount of tar and product gas is low
phenols in product gas is low
Full Report
As per mentioned earlier, using steam as the gasifying agent requires heat source from
outside of the gasifier chamber as steam gasification is an endothermic equation(Kitzler et al.,
2012). Using circulating fluidized bed (CFB) gasifier or the bubbling fluidized bed (BFB)
gasifier alone will not support the gasification thermal requirement as there is no source of
heat unless heat is generated through combustion of other auxiliary fuel is steam is the
gasifying media. Majority of industrial application of CFB and BFB gasifiers are either air
blown or oxygen blown as these reaction will results in exothermic reactions, however the
syngas will eventually contains lesser hydrogen.
Therefore, a Fast Internal Circulating Fluidized Bed (FICFB) which is a type of steam
blown Dual Fluidized Bed Gasifier (DFBG) seems to be a better option for this plant. This
gasifier consists of 2 chambers of reactors where the first reactor is the bubbling fluidized bed
blown with steam to gasify the OPT biomass that is being fed in to produce syngas at the
temperature range of 800 to 900oC(Kirnbauer and Hofbauer, 2011). The bed material
circulates with the resultant char from steam gasification into the second reactor, combustor
consisting of circulating fluidized bed that is blown with air to oxidize or burn out the char in
order to generate necessary heat for the gasification. Basically, the bed material acts as a
heating carrier or medium that is circulating between the two chambers transferring heat from
the combustion to gasification area, without mixing the combustion and gasification product
gases(Göransson et al., 2011). The product gases include flue gas and syngas respectively.
The diagram below shows the flows of the streams within the gasifier:
Figure 2.2.2.1.2: Gasifier and Combustor Chambers
There are several constraints related to this dual fluidized bed gasifier that need to be
taken into consideration and control measures have to be implemented to avoid future
problems. Firstly, the thermal energy that is being required by the gasification process as well
Full Report
as the heat loss that are being encountered by the gasifier has to be balanced by the heat being
produced in the combustor through combustion of residual char(Göransson et al., 2011). It is
important to ensure that the gasifier is at an elevated temperature in order to favor the
pyrolysis, endothermic steam gasification, Boudouard reaction as well as the methane
reforming reaction so that higher yield of hydrogen component in the syngas is
maintained(Kaushal and Tyagi, 2012). The temperature balance of the DFBG is highly
dependent on the char combustion and the circulation of the bed material. When the
temperature of the gasifier hits a lower range, the conversion of biomass into syngas will
reduce and simultaneously will result in higher yield of char. This means that more fuel is
being circulated into the combustor resulting in more heat being generated and transferred to
the gasifier through the bed eventually restoring the required temperature of the
gasifier(Göransson et al., 2011). The system is an auto stabilizing system which is an
advantage over other type of gasifiers.
Since the circulation of the bed material plays an important role in maintaining the
temperature balance, the gasifier gas distributor plates should be designed as such there is no
back flow of the bed material through the nozzles(Göransson et al., 2011).
Furthermore,efficiency of the heat transfer between the combustor and gasifier also depends
on the type bed material used as the heat carrier. It should have a very good agglomeration
behavior in order to be able to circulate the char produced to the combustor
efficiently.Olivine sand as the bed material is suitable to be used as it possess a very good
agglomeration behavior and additionally, it acts as an catalyst to enhance the tar cracking as
well as promote water gas shift and steam reforming reactions leading to higher yield of
hydrogen(Schmid et al., 2012).
Full Report
2.2.3 Evaluation of Post Treatment

Post-treatment of syngas is the gas cleaning process that is carried out after the gasification
process. During the gasification process, the gasifier not only generating useful product but
also a lot of by-product such as, dust, ash, tar, NOx, sulfur as well as SO2(Salam et al., 2010).
These by-products will leave the gasifier together with the syngas. This created a major
problem when utilizing the syngas produced in a plant. Therefore, cleaning of syngas
produced from gasification is essential in order to avoid any fouling and blockage in
equipment and pipelines as well as inhibition of the ammonia synthesis in the downstream.
Gas cleaning not only important for equipment protection, but also necessary to reduce the
emission to a required limit established by Environmental Protection Agency. The gas
cleaning process can be categorized into primary and secondary method. Primary methods
including the appropriate selection of operational parameters, bed material and modification
that has made to the gasifier design. Primary method will not resolve the problem of tar
formation and other unwanted waste but it is effectively reduce the need of downstream gas
treatments. However, secondary method is the conventional cleaning process applied to the
hot gas leaving the gasifier.
The main components that need to be treated before entering the subsequent process
are particulates, ash, dust, tar, CH4, C2H4, C3H6 and C2H6. However, the remaining
components will be removed in CO2 removal and Methanator. Secondary method of syngas
gas cleaning can be divided into two major processes which are mechanical cleaning of dust,
ash and particulates as well as catalytic steam reforming of tar and hydrocarbon.
Full Report
2.2.3.2 Gas Cleaning: Particulates

The particulates, dust and ash come out from gasifier will be removed using physical method
such as cyclones and gravity settling chamber. In order to increase the efficiency of the
particulates removal, the syngas will be cooled down and enters the particulates removal
equipment.
1) Gravity Settling Chamber
Gravity Settling Chamber uses the gravity force to separate the solid particles from
the syngas. The syngas will be entering the chamber and the velocity will be reduced across
the chamber. Large particles will be settling down and fall into the dust collection
hoppers(India, 2013).
2) Cyclone
Cyclone uses the centrifugal force to separate the particulate from the gas stream. In
the cyclone, the large particulate will moved toward the wall of cyclone separator and fall to
the hopper at the bottom(Swanson, 2009).
The advantages and disadvantages of the physical removal of particulates are shown in Table
2.2.3.2.1 below:
Table3.2.3.2.1 Advantages and disadvantages of cyclone and gravity settling tank
Type of Reverse-flow cyclones (Tangential inlet
Gravity Settling Tank
Equipment and vertical reverse flow cyclone)
 Require small area
 Simple construction and operation
 Simple construction and operation
 Negligible maintenance problem
 Little maintenance problem
 No limitation for temperature,
 No limitation for temperature,
pressure and moisture content
pressure and moisture content
Advantages limitations
limitations.
 Low capital investment, operation
 Medium capital investment,
and maintenance costs
operation and maintenance costs
 Dust will be collected and dispose
 Dust will be collected and dispose
(India, 2013)
(Vasarevicius, 2011)
 Require large area
 High efficiency for fine particulates
 Not economical for large gas
(> 15 microns)
capacity
Disadvantages  Not suitable for sticky and
 High efficiency for fine
flammable dusts
particulates (.>60 microns)
(Vasarevicius, 2011)
(India, 2013)
Full Report
2.2.3.3 Gas Conditioning: Removal of Fine Particles and Tar

Tar removal is one of the important steps in gas cleaning process. Tar must be removed
before syngas enter the next process. Based on literature studies, the condensation point of
the tar is at 200°C to 400°C.When the temperature of the syngas is cooled to the condensation
point, the tar will tend to condense and then polymerize to form sticky or solid condensed
deposit which will clog the pipelines and equipment. This results the process plant required
more maintenance and subsequently increase the overall investment cost. To eliminate all
these problems, tar removal procedure should be applied. Tar removal procedure can be
classified into chemical or physical treatment. For chemical treatment, the tar can be treated
by tar cracking either thermally or catalytically. However for mechanical method, the tar can
be removed by using cyclone, filters, rotating particle separator, electrostatic filter and
scrubber(Vreugdenhil and Zwart, 2009). The tar removal methods that are taken into
consideration are thermal cracking, venturi scrubber and metal filter.
1) Thermal Cracking
Thermal cracking is carried out under high temperature. The tar is decomposed
through the pyrolysis process, where the syngas will be heated up to very high temperature.
Under high temperature, tar will be cracked. The process cracks the tar by breaking the
molecular bond and reducing the molecular weight(Salam et al., 2010).
2) Wet Scrubber
The mechanical method used is wet Scrubber (Water Loop Tar Removal). The tar in
the syngas will be removed by entering a water loop which consist equipment like wet
scrubber, mist eliminator and oil/water separator. Water will be used as the recirculation
liquid to scrub the syngas. Wet scrubber is able to remove water content in the syngas to
minimum.
3) Metallic Filter
Metallic filter is a hot gas cleaning process that operated at a temperature range of
250-700°C and at pressure of 10-25 bar(g)(Grasa et al., 2004).The reason of operating at high
temperature is to prevent the condensation reactions, which will then causes fouling and filter
blockage problems. For stainless steel filter, the applicable temperature is up to
420°C(Heidenreich, 2013).
Full Report
Table 2.2.3.3.1: Comparison of different fine particles and tar removal methods
Thermal Metallic Filter (Stainless

Method Wet Scrubber
Cracking Steel)
 High efficiency
 Protection of
downstream heat
 Efficient in exchangers or catalyst
 Very efficient
removing tar unit operations, thus,
in removing tar
Advantages  Can remove water increase the efficiency
 Remove tar in a
content in the of the reactions
large scale
syngas  Elimination of fine
particulates (PM 10
and PM 25.)
(Heidenreich, 2013)
 Efficient waste
water treatment is
required to treat
the water to a  High investment due
 Operating at
dischargeable limit to the increase in
high
temperature  The syngas need to demands of the
be cooled down to materials
Disadvantages  Require high
a very low  Higher gas volume
electricity cost
temperature, thus will required higher
 High
reducing the net temperature.
investment
efficiency of the (Heidenreich, 2013)
process
(Laurence and
Ashenafi, 2012b)
Economic and  High investment
Environmental  High cost due  High investment due
cost due the
Considerations to the high to the increase in
equipment used in
energy demands of the
the loop
consumption materials
 Produce a lot of
for
waste water
Full Report
2.2.3.4 Gas Conditioning: Secondary Reformer

The gasification of oil palm trunk for the production of hydrogen gas is vital in the anhydrous
ammonia production plant. The hydrogen gas supply should be sufficient for ammonia
synthesis in order to produce desired amount of product. Thus, no extra cost will be required
for acquiring the feedstock outside the plant. Besides getting a sufficient amount of hydrogen
gas, the purity of the hydrogen gas also needs to be taken into consideration. All the tar must
be removed and the amount of methane must be reduced to the minimum amount. In order to
meet all these criteria, reformer is equipped in the plant to enrich the hydrogen production by
undergo catalytic steam reforming of methane. In the meantime, hydrocarbon in the syngas
will be totally removed by the reformer and subsequently, it would not affect the downstream
reaction and prevents fouling and clog in the pipelines. There are 2 types of secondary
reformer that were taken into consideration. They are Autothermal Reformer (ATR) and
Steam Methane Reformer (SMR).
Table 2.2.3.4.1: Comparison of different types of secondary reformer

Steam Methane Reformer
Type of Reformer Autothermal Reformer (ATR)
(SMR)
ATR involved two different
reactions in the same vessel. The
first stage is the Partial oxidation
reaction and the second stage is the SMR involved only catalytic
Description
catalytic steam reforming zone. The steam reforming reaction.
methane will be reacted and
produce hydrogen in the second
stage.
Operating Condition (Temperature and Pressure)
The reactions carried out in the reformer are equilibrium reactions. For equilibrium reactions
the production of the desired product is depending on the temperature and pressure. Le
Chatelier’s Principle is applied in these reactions. Therefore, in order to achieve more
production of hydrogen gas, high temperature and low pressure will be used(Padban and
Becher, 2005). Therefore, the reaction will shift to the right and more products will be
formed. However, the temperature and pressure set should not exceed the range.
Temperature,
350 - 700 850-1300
(Feed)
Pressure, bar 20-70 20-30
CH4
65-95 95-100
Conversion, %
Full Report
Table 2.2.3.4.2: Comparison of economic and environmental aspects of different types of secondary reformer
Type of Reactor Autothermal Reactor (ATR) Steam Methane Reformer (SMR)
Heat Supply
Steam reforming reaction in the reformer is highly endothermic reaction. Heat is required to
drive this endothermic reaction.
In ATR, the heat source for
endothermic reaction is obtained
from the partial oxidation reaction Heat source for the endothermic
in the combustion zone. Heat reaction for SMR is obtained by
Source of heat
generated in this stage will be combusting natural gas with steam
contributed to the following in the furnace.
catalytic fixed bed for catalytic
steam reforming reaction.
Energy content in biomass in
lower compared to natural gas.
Energy Content Energy content of natural gas is
More biomass will be required to
of heat source high.
achieve the same amount of heat
as natural gas
Cost of heat generation is cheaper More capital cost is required as
Economic
as no additional feed is required to natural gas is used as the agent to
considerations
supply heat. ignite the combustion.
Emission from the reformer is The reformer leaving carbon
Environmental
lower due to the internal supply of footprint in which natural gas is use
Consideration
heat. for heat generation.
Effect of catalyst on the performance of the reformer

The selection of the catalyst used in the secondary reformer is very important. This is
because the types of catalyst used will affect the reactant conversion as well as the product
yield in the reformer. Nickel catalyst is chosen as the catalyst for the fixed bed in autothermal
reformer. Nickel catalyst is usually used in the secondary reformer due to its high activity
with larger surface area and thus boosts the steam reforming reactions. Besides, nickel
catalyst is able to adsorb a large amount of hydrogen and yet increase the efficiency of the
reactions(Institute, 2007). In addition, Nickel Catalyst is also cheaper. It is reported that the
noble metal based catalysts are less sensitive to coking as compared to nickel. In spite of that,
noble based metal is not considered because it is very expensive and the availability is
limited(Mottos et al.).
Full Report
Effect of steam to methane ratio for ATR

Based on literature studies, the steam to methane ratio need to be adjust to the ratio of
1-5. This should be done to ensure that the carbonization does not occur in the reaction and
the reaction is homogeneous(Khorsand and Dehghan, 2007).

It is very important to remove the fine particle in the syngas to the minimum amount.
Therefore, a reverse-flow cyclone (Tangential inlet and vertical reverse flow cyclone) is
chosen because the particulates’ size coming from the gasifier is estimated to be at the range
of 30 micron-50 micron. Therefore, as comparing with gravity settling chamber, cyclone
which has a higher efficiency in removing the fine particulates which has the particles size of
30 micron was chosen. Besides, the plant capacity of this project is considered small which is
30kt of anhydrous ammonia will be produced per annum. Thus, it is not an economical way
to remove fine particulates using settling chamber. Plus, settling chamber performs better for
larger particulates.
All of the methods shown in Table 2.2.3.3.1 involve high investment. However, wet
scrubber is chosen because it is able to remove fine particulates and tar efficiently. Even
though wet scrubber will be producing a lot of water but considering that the water generated
will be reused in the scrubbing process, wet scrubber is chosen over others suggested method.
Thermal cracking is not taken into consideration because the tar composition in the syngas
stream is very little, yet it is not economical to use this application in this plant. As for
stainless steel filter, it is because it is not worth it to imply this method in the small capacity
plant since the investment cost is high.
Based on the comparison in Table 2.2.3.4.1 and Table 2.2.3.4.2, the chosen
technology is ATR. ATR is more efficient compared to SMR. Taking economic and
environmental issue into considerations, ATR is better because it produces minimum amount
of emission with lower investment compared to SMR. Thus, it leaves lesser carbon footprints.
Moreover, extra cost is required to purchase for the natural gas using for heat generation in
SMR. In fact, ATR will be using the self-generated heat to support the endothermic reaction.
The chosen catalyst used in the reformer is recommended to be Nickel Catalyst. This is
because nickel catalyst is able to adsorb a large amount of hydrogen and yet increase the
efficiency of the reactions.
Full Report
2.2.4 Evaluation of Shift Converter
A shift converter is required in order to further recover hydrogen from the syngas leaving the
reformer. Since the presence of carbon monoxide (CO) and carbon dioxide (CO2) in syngas
exiting the reformer is poisonous to the downstream ammonia synthesis unit; as carbon
oxides are capable of deactivating the ammonia synthesis catalyst; a shift converter has to be
used for the detoxification of syngas(Newsome, 1980). With this unit, CO content in syngas
can be reduced with steam into CO2 and hydrogen (H2). Subsequently, this intermediate
process allows CO2 to be ultimately removed downstream.
2.2.4.1 Technology Evalution

The reaction that governs the CO upgrade to H2 is the water-gas shift (WGS) reaction and it
has gain wide industrial application in the refining process of synthesis gas. The WGS
reaction expressed below is an equilibrium-limited, heterogeneous and exothermic reaction
whereby it is thermodynamically favoured at low temperatures and kinetically favoured at
high temperatures(Smith et al., 2010).
According to Le Chatelier’s principle, the reaction temperature is one of the parameters

affecting the equilibrium reaction whereby the increase in reaction temperature will hinder
the generation of hydrogen. Pressure on the other hand has no effect on the reaction as there
is no change in the volume from reactants to products(Smith et al., 2010). In contrast,
Arrhenius law which explains the temperature dependence of the specific reaction rate
constant in chemical reactions requires the reactants to gain a minimum amount of energy
called activation energy Ea by increasing the reaction temperature so that the forward reaction
of H2 can occur(Lima et al., 2012).
In order to achieve a balance between these two effects, Alternis BioAmmonia Sdn. Bhd. has
decided to utilize a series of High Temperature Water- Gas Shift Reactor (HTWGSR)
followed by a Low Temperature Water-Gas Shift reactor (LTWGSR) with intercooling stage,
so that the task of CO removal could be executed along with a higher purity of H2 in syngas.
Due to the kinetics and thermodynamic of equilibrium constraints, the selection of catalysts
with different rate expressions is crucial as the reaction results are highly dependent on this
parameter. Hence, an iron oxide-based catalyst with a typical reported composition of 74.2%
Full Report
Fe2O3, 10.0% Cr2O3, 0.2% MgO with the rest being volatiles; is used as the catalyst for
HTWGSR while a copper-based catalyst which contains a mixture of ZnO, CuO and Cr2O3/
Al2O3 is used as the catalyst for LTWGSR(Newsome, 1980). Table 2.2.4.1.1 summarizes the
advantages and disadvantages of various catalysts types.
Essentially, the type of reactor chosen for this heterogeneous catalytic process is a multi-tube
fixed bed reactor as it is able to accommodate stack of catalyst pellets that are compact and
immobile within a vertical vessel. The CO shift reaction is generally conducted in an
insulated adiabatic reactor with temperature increasing along the catalyst bed due to the
exothermic process(Callaghan, 2006). Instead, Alternis BioAmmonia has decided to use a
multi-tubular fixed bed reactor with cooling water circulation in order to keep the reactor
isothermal. It is important to maintain a constant temperature because temperature rise along
the catalyst bed is unfavourable as it may affect the equilibrium conversion, the product
selectivity, the deactivation of catalyst and in extreme cases unsafe operation due to runaway
reactions(Jakobsen, 2008, Eigenberger, 1992).On the other hand, in order to limit the
temperature increase per bed, a multi-tube reactor is recommended as it is able to contain
hundreds or thousands of tubes with an inside diameter of only a few centimetres and
maximise heat transfer to the boiler feed water that will ultimately prevent excessive
temperatures and hot spots(Jakobsen, 2008). Furthermore, the regulation of temperature by
steam pressure is flexible and possible in a multi-tubular fixed bed reactor with boiler feed
water circulation. According to the Linde Group, a capacity of up to 4000MTPD is feasible
in this type of reactor depending on the process condition(2013).
Full Report
Table 2.2.4.1.1: Comparison between catalysts involved in WGSR

Type of Catalyst Advantages Disadvantages
Iron-based shift - Commercially available - The energy efficiency of the
catalyst - Industrially proven reaction is highly dependent on
- Cr2O3 in catalyst acts as a the steam to carbon monoxide
stabilizer and prevents high ratio as reaction conducted at
temperature sintering and loss low ratios could lead to side
of surface area reactions producing unwanted
- Operating temperature range by-products (Callaghan, 2006)
of 250-400 ; with higher
temperatures resulting in
decreased activity due to
sintering of catalyst
- Not particularly susceptible
to poisons as small amounts
of sulphur (<50ppm) have an
insignificant effect on
catalyst(J.M., 2009,
Newsome, 1980)
- Suitable to be used in
HTWGSR
Copper-based catalyst - Remains active at - The reaction cannot be carried
temperatures as low as 200 out at temperatures exceeding
- Industrially proven 220-230 as the highly
- Zinc oxide provides some exothermic reduction of CuO
protection of the copper from could cause sintering
sulphur poisoning while - More prone to deactivation
acting partially as a support caused by sintering due to the
for the copper (Callaghan, relatively low melting point of
2006) copper(Lide, 2004)
- Selectively fewer side
reactions occurring at higher
Full Report
operating pressures
- Normal operating span of 2-3
years (Rase, 1977)
Cobalt-based catalyst - High tolerance to sulphur - Low activity at temperatures
- Preserves their catalytic between 200-300
activity for the WGS reaction - Introduction of Co as a catalytic
regardless of the absence or promoter increases the yield of
presence of sulphur (Farrauto by-product methane while
et al., 2003) causing a depletion in hydrogen
- Exhibits higher catalytic yield
activity compared to standard - Loss of surface area by
commercial iron-based sintering and reduced catalytic
catalyst (Farrauto et al., ability when there is a
2003) temperature rise due to
exothermic CO hydrogenation
(Hutchings et al., 1992)
Gold-based catalyst - High academic and industrial - No industrial applications are
recognition found using Au-ceria materials
- High activity at low at low (Mendes et al., 2009)
temperatures and potential - Only comply to very specific
stability in oxidizing experimental conditions
atmosphere(Mendes et al., (catalyst preparation; pre-
2009) treatment; operating conditions)
- The catalytic activity has to be
increased by 10-100 times over
conventional material in order
to compete on a cost basis for
WGS (J.R. and J.P., 2003)
Full Report
Table 2.2.4.1.2: Comparison of Shift Converter Technologies

Technology Advantages Disadvantages
One-stage - Cheaper in terms of number - Unable to achieve high purity
of equipment and catalysts of hydrogen
used - Steam requirements are
- Less energy intensive considerably more than two-
compared to two-stage stage operation
operation
Two-stage - Able to achieve high purity of - Consumes more energy
hydrogen in syngas compared to one-stage
- Lesser steam requirements as operation (J.M., 2009)
compared to one-stage
operation(Callaghan, 2006)
Membrane reactor - Combines catalysis and - Difficulties in the formation
membrane for higher of a uniform metallic thin
selectivity of H2(Mendes et al., film
2009) - A fairly new technology for
water gas-shift reaction in
industries (still under
research)
- Few literature sources
evaluate the cost analysis of
this devise
- Scale up of water gas shift
membrane reactor s is
still under research (Mendes
et al., 2009)
Full Report
2.2.5 Evaluation of Carbon Dioxide Removal Process

Emission of greenhouse gases (GHG), especially carbon dioxide (CO2), into the environment
resulting in global warming has attracted widespread concern and attention. Thus, CO2
removal and capture has become an important step in many processes in most of the
industrial process plants. In ammonia production plants, CO2 is being removed from the
process stream as it is an undesirable component in the syngas. Presence of CO 2 tends to
cause temperature excursions in the process. Moreover, CO2 may poison the iron catalysts
present in the ammonia synthesis reaction in downstream process, which will affect the
reaction rate in the reactor and lead to lower production of ammonia(Derks, 2006). This will
also increase the production cost of ammonia as new active catalysts have to be replenished
frequently to replace the poisoned catalysts.
Membrane separation is a capture concept that uses selective membrane that is semi-
permeable to separate and remove different components from the gas stream. Membrane
separation has various viable working mechanisms, which include Knudsen diffusion,
molecular sieving, solution-diffusion separation, surface diffusion and ionic
transpoer(Mondal et al., 2012). In membrane separation process, the sour gas (syngas) stream
will pass through the membrane that acts as a barrier to separate CO2 out from the stream.
The CO2 rich stream is known as permeate while the purified gas stream is called retentate.
Pressure difference across the membrane is normally the driving force for the flow of gas
through the membrane(Metz et al., 2005). The performance of membrane separation is
decided by two characteristics; permeability of the membrane which is the flux of a gas
species passing through the membrane, and selectivity of the membrane which is the
preference of one gas species to pass through over the other gas species(Olajire, 2010).
Absorption-desorption method is a common technology used in removing and capturing CO2.

Normally, CO2 removing process takes place in two different operating units, namely
absorption column and stripping column. Absorption column is used for removal of CO2 from
the gas stream whereas stripping column is responsible for solvents regeneration. The sour
gas (syngas) stream will enter the absorber column from the bottom and contact with the
solvent stream that flows counter-currently from the top of the column. CO2 in the sour gas
stream will be removed and absorbed by the solvent and leaves the column at the bottom. The
gas stream that has been purified will leave the absorber at the top and continue to the next
Full Report
process. While in the stripping column, the absorbed CO2 will be stripped from the solvent
from the effect of higher temperature or lower pressure. The CO2 will leave at the top and the
regenerated solvent will leave at the bottom and recycle back to the absorption column.
However, membrane separation process can only achieve low degrees of separation(Olajire,
2010). Thus, multiple stages or recycling is necessary in order to achieve higher separation
degree, but this will in turn increase the cost which is not economically viable. Moreover,
membrane may be clogged by impurities in the gas stream which will lead to low separation
efficiency and less CO2 can be recovered at later stages. Impurities in the gas stream may also
pass through the membrane together with CO2 which will results in low purity of product(Yu
et al., 2012). On the other hand, efficiency of CO2 removal of absorption-desorption method
is much higher(Yu et al., 2012). In stripping column, solvent is regenerated and recycled back
to the absorption column to be reused. This can reduce the makeup cost of the solvent and is
more economically viable. Besides that, CO2 stream exit from the stripping column has much
higher CO2 purity than that of membrane separation process(Smith et al., 2012). Therefore,
absorption-desorption process technology is chosen as the CO2 removal technology in the
anhydrous fertilizer-grade ammonia production plant proposed by Alternis BioAmmonia Sdn
Bhd.
The solvent used in absorption-desorption method is either chemical solvent or physical

solvent. Physical solvent is an organic solvent and it reacts with acid gases physically without
chemical reaction. Absorption using physical solvent is highly dependent on temperature and
pressure. Physical solvent absorbs acid gases in proportion to the solubility of CO2 in the
solution and the concentration of the solution. The commonly used physical solvent are
Selexol, Rectisol and Fluor. Selexol is a mixture of dimethylether of polyethylene glycol and
it is widely used for removal of bulk CO2 and H2S. Whereas, Rectisol process uses chilled
methanol in removing the acid gases. Selexol and Rectisol are widely used as the cost of the
solvent is low(Mondal et al., 2012). Furthermore, solvent regeneration for both the processes
can be done easily as there is no chemical process involved. However, Selexol solvent has
high affinity towards hydrocarbon and Rectisol solvent has high affinity towards metallic
trace components like mercury(Mondal et al., 2012). This indicates that the purity of CO2 is
low and extra cost may be required in order to further purify CO2. Both processes require
extra refrigeration cycle to cool the solvent which will result in high operating cost of the
plant(Padurean et al., 2012). As for Fluor, the solvent is propylene carbonate, is used to treat
Full Report
CO2 when partial pressure of CO2 is high. Fluor solvent has high CO2 solubility and loading
capacity. It does not require makeup water as well as heat duty for solvent regeneration.
However, Fluor solvent is very costly and has high circulation rate. Similar to Selexol, Fluor
solvent has high affinity to hydrocarbon which will lower the purity of CO2(Padurean et al., 2012).
Thus, in short, CO2 removal process using physical solvent will lead to low purity of CO2.
This is not favorable as high purity of CO2 is expected in this project, thus physical solvent is
not selected to be used.
Chemical solvent absorbs the sour gas, and heat is needed in order to reverse the reaction to
release the absorbed gas and to regenerate the solvent. The chemical solvents that are mostly
used are various kinds of alkanolamines, such as monoethanolamine (MEA), diethanolamine
(DEA) and methyldiethanolamine (MDEA). MEA is a primary amine, DEA is a secondary
amine and MDEA is a tertiary amine. MEA and DEA is commonly used for H2S and CO2
removal from natural gases and synthetic gases in the industries. MEA solution has high
alkalinity, which results in high efficiency in absorption of acid gases(Aden, 2009). DAE has
low vapour pressure which makes it suitable to be operated at low-pressure condition.
However, although both MEA and DEA have high reactivity with CO2, but CO2 loading
capacity is low. MEA can only be used to treat the gas stream with low concentration of acid
gases as high concentration of acid gases will cause degradation of MEA. As for DEA,
retrieving process of contaminated solution will be difficult because vacuum distillation may
be required(Aden, 2009). MEA solution requires high energy for solvent regeneration as it
has high heat of reaction. In addition, MEA and DEA pose high corrosion rate to the
equipments. For the tertiary amine, MDEA has low heat of reaction which will lead to low
regeneration energy of solvent. MDEA has high CO2 loading and low corrosion rate.
However, the reactivity of MDEA with CO2 is very low which results in low efficiency of the
solvent. Table 2.3.6.1.1 shows the comparison of chemical and physical solvent.
Full Report
In order to compensate the weaknesses of the chemical solvents, a tertiary amine is usually
mixed with a primary or secondary amine. Activated dimethyldiethanolamine (a-MDEA) is a
solvent used widely and in majority of the world’s ammonia plants(Alvis et al., 2012). a-
MDEA is consisted of MDEA activated by piperazine, the most commonly used promoter for
amine solvent solution(Kunjunny et al., 1999). a-MDEA is highly reactive with CO2, which
will enhance the absorption rate of CO2 into the solvent. One of the most significant
advantages of a-MDEA is that it acts as chemical solvent when partial pressure of CO2 is low
and as physical solvent when partial pressure is high(Kunjunny et al., 1999). a-MDEA has
low makeup rate as it has low vapour pressure and rate of solution loss is low. Besides that, a-
MDEA solvent solution has high thermal and chemical stability, which lead to long shelf life
of the solvent(Kunjunny et al., 1999). a-MDEA is also biodegradable and non-toxic, which
will in turn reduce pollution and damage to the environment as the solvent is environmental
friendly. Moreover solution of a-MDEA is non-corrosive, which results in lower operating
and maintenance cost(Alvis et al., 2012). Therefore, a-MDEA is chosen to be used as the
solvent in CO2 removal process in the anhydrous fertilizer-grade ammonia production plant
proposed by Alternis BioAmmonia.
Full Report
Table 2.2.5.1.1: Comparison of Chemical Solvent and Physical Solvent

CO2 Removal Process Advantages Disadvantages
 High Alkalinity  Can only be used to treat acid gas in
 High Acid gas Absorption low concentration
efficiency  High solvent make up rate and
MEA
 High CO2 reactivity regeneration energy
 High corrosion rate
 Low CO2 loading capacity
Chemical
 Low vapour pressure  Complex retrieving process of
Absorption
 High CO2 reactivity contaminated DEA solution
DEA
 High corrosion rate
 Low CO2 loading capacity
 Low regeneration energy  Low reactivity with CO2
MDEA  High CO2 loading capacity
 Low corrosion rate
 Low solvent cost  Low purity of CO2 produced
Selexol
 Simple solvent regeneration
 Low solvent cost  Low purity of CO2 produced
Physical Rectisol
 Simple solvent regeneration
Absorption
 High CO2 solubility  Expensive solvent
Fluor  High CO2 loading capacity  High circulation rate
 No makeup requires  Low purity of CO2 produced
Full Report
Moving on, the type of column used for the absorber is the packed column. Since the flow
rate of amine solvent used for the CO2 removal process is high, a packed column is suitable
to be used as it is effective in handling large liquid rate. Packed column would have shorter
tower height as compared to tray column, and it is mechanically simple (Pilling & Holden,
2009). On top of that, the gas-liquid contact in a packed column is continuous, where the
liquid flows down the column over the packed bed and the vapour flows up the column
counter-currently (Sinnott & Towler, 2009). This would increase the contact area and contact
time between the liquid and vapour, and hence increase the efficiency of the process. Packed
column is also more economically beneficial for handling corrosive system (Sinnott &
Towler, 2009). The amine solvent used in the system is corrosive, and the corrosive behavior
of dissolved CO2, thus packed column is suitable to be used. Furthermore, packed column
could be operated at lower pressure drop as compared to tray column (Pilling & Holden,
2009).
For the packing material used in the packed bed of the absorption column, INTALOX saddle
ceramics, random packing, are chosen. Random packing is chosen over structured packing
for the absorption column in this project due to several advantages of random packing. Firstly,
cost of random packing is significantly lower than the cost of structured. This is economically
beneficial as the capital cost could be reduced. Next, the packings are placed in the packing
bed randomly without specific arrangement. Random arrangement of the packings is able to
improve the liquid distribution, which will results in more contact opportunities between the
liquid and the vaour that flows counter-currently and thus higher process efficiency (Sinnott
& Towler, 2009). Ceramic material is chosen because it is more suitable to be used to handle
the corrosive environment in the absorption column. INTALOX saddle ceramic is shown in
the figure below.
Figure 2.2.5.1.1: INTALOX Saddle Ceramic (Pilling & Holden, 2009)

Full Report
2.2.6 Evaluation of Methanation
In methanation process, hydrogen reacts with carbon monoxide and carbon dioxide to
produce methane and steam. It is important to remove the oxides in ammonia synthesis
process as oxides would decrease the activity of ammonia synthesis catalyst and cause
deposition of ammonium carbonate in the synthesis loop. (UN Industrial Development
Organization, 1998)Carbon oxides removals are also required for the protection of
hydrogenation and ammonia synthesis catalyst against rapid deactivation and also prevent
damages in the reactor. Methanation reactions are the reverse of the reformer reaction. The
chemical equations involved in this process are:
Both methantion reactions are exothermic and methane yield is favoured at lower
temperatures. The forward reactions are also favoured at higher pressures. However, the
space velocity becomes high with increased pressures, and contact time becomes shorter,
decreasing the yield(Matar and Hatch, 2001). The normal operating condition of the process
is 250 -300 while the pressure should be at least 18bar. If the pressure is too low, the
targeted CO conversion will not be achieved. (Heyne et al., 2010)
2.2.6.1.1 Pressure Swing Adsoption (PSA)

PSA can also be used to remove carbon dioxide, methane and small amount of carbon
monoxide. PSA process is based on the selective adsorption of gaseous compounds on a fixed
bed of solid adsorbent in a series of identical adsorption beds. Multiple adsorbers are used to
provide constant feed, product and tail gas flows. Typical PSA process requires 4-10 columns.
Adsorbents are selected for each application based on the type of impurities present in the
feed stream. (Kubek et al., 2009)The common adsorbent is an active carbon or a carbon-
molecular sieve. Each bed undergoes a repetitive cycle of adsorption and regeneration steps.
PSA provides the hydrogen about the same pressure as the feed, but recoveries are typically
lower than from other technologies.(Riegel, 2010) The system operates at room temperature
and pressure of 20-25atm(M.A.Fahim et al., 2010).
Full Report
2.2.6.1.2 Molecular Sieving System

Molecular sieving is a process to remove impurities e.g. H2O, H2S, CO2 etc. Molecular sieves
are normally zeolite-based adsorbents which contains crystalline aluminosilicates and clay.
The surface acts as an adsorption site for polar materials even at very low concentration. The
typical system consists of 2 or more fixed-bed adsorbers and regeneration facilities. It works
by feeding the gas stream in one column while regenerating in another column where both
adsorption beds are used in series. The common available molecular sieves include 3A, 4A,
5A and 13X. The molecular sieve system conducts at a temperature range of 25 -400 and
operating pressure of up to 100barg. (Mokhatab and Poe, 2012)
2.2.6.1.3 Membrane Separation

Membrane can be used to remove acid gases e.g. CO2 by passing the gas through a bed of
membranes. The membrane separation is based on selective permeation. The separation
depends on the rate of gas to dissolve through the membrane surface as each gas has its
specific permeation rate through the membrane. Acid gas like CO2 has higher permeation rate
as compared to gases like N2 and CH4. When the feed stream is passed through the membrane,
the methane-rich gas will be left at the exterior of the membrane fibre while the acid-rich gas
will permeate through the interior of the membrane fibre.(VMEprocess, 2012)
2.2.6.1.4 Cryogenic Purification

In the cryogenic purifier, all the methane and excess nitrogen will be removed from the
synthesis gas as well as part of the argon. The syngas exiting the cryogenic purifier will be
practically be free of all impurities except for a small amount of argon and methane before
entering the compressor for ammonia synthesis. The front-end technology simultaneously
removes impurities from synthesis gas by washing it with excess liquid nitrogen while
adjusting the hydrogen to nitrogen ratio to 3:1. This ratio can be controlled independently.
The cooling is produced by depressurization and no external refrigeration is required. There
are several values and benefits of providing clean, dry, make-up gas to the synthesis loop and
precise hydrogen to nitrogen ratio control. This includes reduced energy consumption,
reduced capital costs, increased flexibility of operation and increased reliability of associated
process equipment.(KBR, 2013) This cryogenic purification after methanation process results
a more efficient process compared to the conventional purification by methanation only due
to the combination of high conversion per pass and reduced purge flow.
Full Report
Table 2.2.6.1.1: Comparisons between different technologies

Technology Advantages Disadvantages
 Simple process  Hydrogen consumption
 Low cost  Production of additional
Methanation
 High conversion inert gas
(Maxwell, 2005) (Maxwell, 2005)
 Ability to remove
impurities to any level
 Large systems and high
 High hydrogen purity
cost due to slow cycle
Pressure Swing (Kubek et al., 2009)
speed
Adsorption (PSA)  Carried out at room
 Complex design
temperature
(Babicki, 2003)
(M.A.Fahim et al., 2010)
 High affinity for polar  Poor chemical resistance

molecules and unsaturated  High energy expenses
organic compounds  Requires recycle gas
Molecular Sieve
 Removes water and acid regeneration stream
gas simultaneously  Lower product recovery
(Armarego and Chai, 2009) (Armarego and Chai, 2009)
 High recovery of products
 Moderate purity
 Good weight and space
Membrane  Removes high level of CO2
efficiency
separation stream
 Less environmental impact
(Shimekit and Mukhtar, 2012)
(Shimekit and Mukhtar, 2012)

Methanation process was selected for gas purification process due to its design simplicity,
low installation cost, high conversion and smaller site requirement as compared to the other
technologies. Although PSA is able to achieve high conversion of impurities removal, it
requires a large site and complex set up. Moreover, the amount of CO and CO2 are present at
a very small amount, so a simpler process will be sufficient. Molecular sieve is not selected
because it is less effective and it is usually used to remove final traces of H2O and CO2 after
Full Report
methanation step, so it is not suitable to be the main process of gas purification in this plant.
Membrane separation was not selected as it usually removes CO2 from gas products that
contain high levels of CO2, which contrast with the situation at this stage of the plant.
In this plant, an adiabatic fixed bed reactor was chosen for catalytic methanation reaction.
Adiabatic fixed bed reactor provides better opportunities for power generation due to heat
release at higher temperature compared to isothermal fluidised bed reactor. However, the
syngas productions do no differ significantly for both types of reactors.(Seemann, 2006) The
catalyst used in the reactor is Nickel Alumina catalyst as it is relatively cheap, very reactive,
and it is the most selective to methane compared to other metals.(Dyer et al., 2013) The CO
conversion for the methanation step is required to be above 99.99% in order to achieve a low
CO content in the methane rich gas. (Heyne, 2013)
Cryogenic Purifier will also be used to remove excess nitrogen and part of argon and methane
in this plant. It will be used to adjust the ratio of hydrogen to nitrogen ratio to 3:1. This
technology was chosen because of its ability to ratio of hydrogen to nitrogen can be
controlled independently and also due to its high product purity.
Full Report
2.2.7 Evaluation of Ammonia Synthesis Process
Ammonia synthesis occurs on a catalyst at pressures above 100 bar and a temperature range
of 350 to 550°C. As it is a highly exothermic reaction, the temperature needs to be controlled
to favour the equilibrium reaction and at the same time avoiding catalyst denaturation when
operated at high temperatures. Several technologies which can be employed for temperature
control converter are discussed in the following sections.
3H2 + N2 2NH3 H0298 = – 92.4 kJ/mole
2.2.7.1.1 Converter Design

An adiabatic design is preferred over an isothermal reactor, as it more common in the
industry and does not involve any temperature regulations. The main types of adiabatic
converters available can be classified as discussed below.
Continuous packed bed

This would be one of the simplest designs. The synthesis gas fed which is fed into the packed
catalyst chamber becomes excessively hot due to the heat released from the exothermic
reaction not being removed. Subsequently a substantial reduction in catalyst activity,
reducing the ammonia yield could result even after a short operation period(Hindrichs, 1962).
The following are examples of multi-bed converters:
Quench converters
As for quench converters, only a portion of the synthesis gas is sent to the first catalyst bed at
about 4000C. The volume of the catalyst used in the first bed is chosen so that the reacted gas
will leave it at about 500 °C. The packed converter is then cooled by the injection of cold gas
(125–200 °C) between the separate catalyst sections before entering the next catalyst bed.
This also reduces the ammonia content in the reacted gas. The cooled synthesis gas then
enters into the second catalyst section where it will react further. The cooling of reaction
gases is alternated with heating as the reaction proceeds in subsequent catalyst beds in the
converter. Although this type of reactors prevents overheating of the catalyst used and also
maintains a decent reactor efficiency, a disadvantage would be that not all of the synthesis
gas flow over the entire catalyst volume(Hindrichs, 1962). Consequently, most of the
Full Report
ammonia formation occurs at higher ammonia concentration and thus greatly reducing the
rate of reaction. Typically, this translates to a larger catalyst volume being required as
compared to indirectly cooled multi-bed reactor. Also, the relative capacities of such
converters are expected to be lower compared to converters of the same length. However, the
total volume of the converter would remain about the same as an extra space is not required
for inter-bed heat exchangers(Appl, 2005). The total effect is that the relative capacity of this
converter is lower than that of other converters of the same overall length.
Multi-bed Converters With Indirect cooling

In this type of converters, the same cooling pattern as described above is now carried out
between the catalyst beds by indirect heat exchange with a cooling medium. The cooling
medium generally used would be a cooler synthesis gas or boiler feed water to produce high
pressure steam. The heat exchanger can be installed together with the packed catalyst beds
inside one pressure shell, or the catalyst beds may be accommodated in separate vessels and
heat exchangers are used separately. Certain technical difficulties are involved with such
design; mainly the vessel has to be safeguarded against leakage from the tubes which contain
the cooling medium. Also, a very long heat exchange surface may be required to heat the
synthesis gas to the reaction temperature (3800C-4500C) using only the reaction gases exiting
the final catalyst bed(Appl, 2005). This converter design however provides an excellent
solution in controlling the reactor temperature profile for maximum ammonia production.
Besides that, with an installation of a secondary heat recovery system, approximately 0.8 ton
of high pressure steam for every ton of ammonia produced may be achieved.
Tubular
In tubular converters, the heat of reaction is removed by having cooling tubes running
through the packed beds. The nitrogen hydrogen mixture then flows either counter currently
or concurrently with the reacted gases, where heat is then transferred to reactor feed gas
heating it to the reaction ignition temperature. It is however difficult to maintain the operating
temperature from increasing to above 550°C in the last catalyst bed. A temperature of 580°C
and higher not only results in the premature denaturation of the catalyst but also reduces the
yield of ammonia due to large deviation from the optimum reaction rate curve(Schl€ogl,
1991). Such converters are only suitable for smaller production capacities and are currently
obsolete.
Full Report
2.2.7.1.2 Comparison between Axial and Radial Flow Converters

In axial flow reactors, the feed goes in at one end of the reactor, passes through the catalyst
bed in the direction along the axis of the reactor and leaves from the other end. As for radial
flows, the feed is distributed along the length of the converter, flowing in radial direction
over the catalyst bed contained in between two perforated screens (known as basket) (Li,
2005). Axial flow converters generally face a problem with increasing capacity, as the
diameter of the reactor has to be increased in order to maintain a similar pressure drop across
the catalyst bed. However, with increasing diameter, the shell of the reactor becomes too
thick, extremely heavy and costly for fabrication. Therefore, due to technical and economic
reasons, the bed diameter, subsequently the pressure vessel diameter should not be increased
above a particular limit. Beyond a certain plant size, a single axial-flow reactor gets
impractically huge, thus requiring several converters in parallel(Li, 2005). On the other hand,
radial flow converters are easy to scale up by increasing the length instead of having
excessive diameters. Besides that, as reacting gases enter along the length of reactor, a bigger
surface are is made available for the flow with smaller catalyst beds resulting in lower
pressure drop in a radial converter compared to an axial flow catalyst bed. With an increase
in pressure drop in axial flow converters, compensation is made for this dilemma by usage of
larger catalyst particles. This adds to its disadvantage as smaller particles (as the one used in
radial) have a higher activity compared to larger particles, essentially attributed to its
diffusion constraint(Appl, 2005).
2.2.7.1.3 Comparison of Catalyst

The ammonia synthesis reaction is too slow that practically no reaction occurs in an
acceptable amount of time. Therefore, catalyst is required to ensure the reaction is rapid
enough for a dynamic equilibrium to be achieved in a short period of time during which the
gases are passing through the reactor. Examples of different types of catalyst used in an
ammonia converter are discussed briefly in the following table.
Full Report
Table 2.2.7.1.3.1: Comparison of catalyst available

Catalyst Advantages Disadvantages
-Widely accepted in the
-Extremely prone to
industry
containing high level of
-Economical as it is cheap
impurities
-Less susceptible to catalyst
Iron based -Low ammonia conversion
poisoning
-Catalyst only reduces at
-Stronger mechanical
high pressure, resulting in
strength
higher compressor duty
-Easiest reduction
-Hydrogen may poison the
catalyst by inhibiting certain
sites for nitrogen
dissociation
-Energy savings are offset
-Higher activity per volume
by higher energy
resulting in lower catalyst
requirements for
volume
refrigeration of ammonia
Ruthenium -Lower operating pressure
-Expensive
-Higher conversion rate per
 104 times more
pass
expensive than
magnetite catalyst
 Price has increased
by a factor of 4 since
2000)
-Ammonia has poisoning

-Catalyst is stable even
effect
when operated at high
Cobalt modified magnetite -Conversion is two times
temperatures
lower than ruthenium
Full Report
2.2.7.1.4 Available Technologies
Table 2.2.7.1.4.1: Comparison of available technologies in the market

Companies/ Reactor Process
Advantages Disadvantages
Licensors Type Conditions
Vertical, Pressure :
-For large scale
multiple 200 to 250 -Uses iron based catalyst
production of 4500 mtpd
Ammonia catalyst beds, bar therefore it is more
-Operates at high pressure
Casale axial-radial Conversion: economical
requiring a higher
flow 20%
compresser duty
Pressure: -Iron magnetite catalyst is

150 to 175 used and thus it cheaper
-Cooled via quenching
Horizontal, bar -Low pressure drop of
therefore, higher catalyst
multiple Temperature: about 2.3 bar
volume is required
catalyst bed, 360 to 5100C -Ease in loading the
radial flow Conversion: catalyst loading/unloading
20% because it is a horizontal
Kellogg
design
KAAP Pressure:
-Not cost effective as the
Design: 90 bar
catalyst used is ruthenium
Four vertical Temperature:
-Low Operating Pressure (to enable operation at low
beds, radial 3700C
pressures)
flow Conversion:
20%
Pressure: -Iron based catalyst is used
150 to 200 and it is cost effective
Vertical,
bar -Intermediate cooling via -Operates at high pressure,
multiple
Topsoe Temperature: inter-bed heat exchangers therefore compressor duty
radial flow
370- 5400C -Can be scaled down to is higher
catalyst bed
Conversion: smaller production
20% capacities
Pressure :
Vertical,
140- 210 bar -Only suitable for large
multiple -Small grain iron catalyst
Uhde Conversion: capacities at about 1000
catalyst beds,
20 % mtpd
radial flow.
Full Report

Based on the evaluation of current technologies available in the market as summarised in the
table, the technology chosen for the converter design is the Topsoe, vertical multi-bed radial
flow converter. This reactor is designed for indirect cooling where the synthesis gas is fed at
the bottom of the reactor, flows as a cooling gas to the top of the reactor. It is then passed into
the tube side of the inter-bed heat exchanger where it gets heated up to the reaction
temperature and then passes through the first catalyst bed from the outer core radially inwards.
This exit gas from the catalyst bed then enters the shell side of the intermediate heat
exchanger before entering the second bed. The final exit gas leaving the reactor is then passed
through a series of heat exchangers for further heat recovery. Based on the selection criteria
of catalyst used, operating pressure, way of cooling, direction of flow and conversion; this
design fulfils the requirements of Alternis BioAmmonia. With a radial flow converter, a
lower pressure drop is achieved without excessive increase in diameters of the vessel making
it less expensive and easy to fabricate. Iron based catalyst used in the packed beds makes this
design more economical compared to the Kellogg design which runs of the extremely costly
ruthenium catalyst. As the production capacity of Alternis BioAmmonia is only 100 tpd, this
design could sufficiently cater for the specified scale.
Full Report
2.2.8 Evaluation of Ammonia Separation and Refrigeration Cycle
After ammonia is synthesized from the reactors, the crude ammonia generally contains traces
of Hydrogen, Nitrogen, Methane and inert gas Argon (Wostbrock, 2001). In the current
industries in the market, large amounts of pure ammonia are needed with preferably less than
3 ppm of impurities (Shields, 2013). Therefore, it is essential to remove these impurities to
obtain a maximum purity producing pure anhydrous Ammonia. Some of these gasses for
example Hydrogen and Nitrogen can be recycled back to the reactor stream to be reused and
hence makes the system more economical.

There are a few technologies currently used in the market to separate pure ammonia out from
crude ammonia. One of the most common methods for pure ammonia to be isolated from
crude ammonia is solely by distillation. Ammonia which is the intermediate-boiling desired
product needs to be separated from low boilers (methane, hydrogen, nitrogen and argon) and
from high boilers (water) (Wostbrock, 2001). Also due to this, a dividing wall column or a
thermally coupled distillation needs to be added. However addition of these requires pre-
vaporization of feed that causes feed to have a mixture of both gas and liquid phase.
Furthermore, with these additions, more energy will be required. Product will be separated
based on their boiling points. Since multi-component mixture is present in the Ammonia
production plant, a side off take will be introduced but the product removed still contains
large amount of impurities and may not be pure to reach the industrial standards. The column
can be configured as packed column or tray column. In provisions of cost wise, it is more
sensible to use trays such as sieve or valve trays. The operating pressure is usually between
10 to 20 bar with a preferred number of theoretical plates ranging from 25-30 (Wostbrock,
2001). The average cost of a 15tray carbon steel distillation column is $12000.
Another separation technology is by using membranes. This involves contacting ammonia

with membrane and allowing permeability of only ammonia molecules. It is ideal to choose a
membrane that is many times more permeable to Ammonia than to Nitrogen and Hydrogen.
Such membranes present are made of cross-linked block copolymers of styrene and
sulfonated styrene. Other material to produce the membrane includes cellulose acetate,
polyvinyl- ammonium thiocyante and polyperfluorosulfonates (Phillip, et al., 2007). These
membranes are 50-150 μm thick and operate by a diffusion-solubility mechanism. However
Full Report
this technology is at a higher cost and more currently more common and efficient for Carbon
Dioxide separation. They are also more advance and available in trapping high concentration
of Hydrogen molecules that is not our intention in this plant.
2.2.8.2 Chosen Technology: Flash Vessel

The third and most appropriate technology for this plant is the flash vessel. Based on the
Joule Thomson effect, when condensates are discharged from a higher pressure stream into a
lower pressure area inside the flash vessel through a throttling device, the condensate
temperature will correspondingly drop to that temperature inside the vessel. This can be
referred to as “auto-refrigeration”. The heat released from this will be use to evaporate some
of the condensate forming flash steam. This steam formed can be trap using steam traps and
to be reused in other units while the purge stream is to be disposed. Purging is important to
avoid any build up in the reactor. The liquid and vapor will then separate under equilibrium.
The vessel is an implication of the Cryogenic system that operates at extremely low
temperature and high pressure to separate the components (Sinnot, 2009). The vessel is also
known as a flash drum if the evaporation occurs within the vessel. The flash vessel
technology is a single stage continuous operation and the vessel is assumed to be adiabatic
with no heat loss or gain. Liquid ammonia is a single component liquid and thus part of it will
“flash” into vapor separating the pure ammonia from the other impurities. The flash
evaporation of a single component liquid is isenthalpic and can also be known as flash
distillation or adiabatic flash. The final products will correspondingly be at a low
temperature and hence no further cooling is required in the case of ammonia for storage.
The flash vessel is one of the simplest unit operations and therefore it will be economically
feasible to install and also maintain it. Comparing to a distillation column of approximately
the same volume, the vessel is slightly more expansive (Owlnet 1997). It can be constructed
using carbon steel or stainless steel material depending on the allocation of cost. The
installation can be supported on the body itself and may not need supporting feet (Spiraxsarco
2013). It is important to ensure that the steam released by the vessel is dry to avoid any
droplets form due to carryover and so vessel must be large enough to ensure this. There will
be pressure gauge and safety valves installed to minimize any risk of unwanted incident from
happening. One of the most common problem occurring is uncontrolled evaporation that
causes boiling liquid expanding vapor explosion, BLEVE (Roberts, 1999).
Full Report
In the vessel, there is no reaction occurring and thus it can be safely stated that there is no
direct harmful emission produced other than the purge gases from the vapor outlet that
consist mostly of inert gas in an ammonia plant. The flash steam produced can be transported
to other units in the plant to be used for example to drive compressors. Any un-reacted
Hydrogen or Nitrogen will be compressed and recycled back to the reactor and this leads to
the minimization of new feedstock required. High purity can be achieved from a flash vessel
ranging from 99% to 99.8% (Google Books, 1968).
2.2.8.3 Refrigeration Cycle and Storage of Ammonia

The term anhydrous in anhydrous ammonia means “the absence of water”. Although it is
classified as non-flammable, it can burn in high concentrations especially in confined spaces
and large spills (Nowatzki, 2007). Even in discrete amount, anhydrous ammonia is corrosive
and hazardous to human’s health. The physical property of ammonia requires it to be stored
in low temperature and thus refrigeration cycle will be incorporated to cool ammonia to a
temperature of -30°c for safe storage purposes. A suitable refrigerant will be selected to be
used in a heat exchanger together with the best compressor cycle. The refrigerant will run
within the loop consisting of a compressor, condenser and a throttling valve (Price, 2007). A
comparison was done and the final choice was Ammonia refrigerant together with a
reciprocating compressor in a vapor compression cycle. Vapor compression cycle is one of
the simplest and most common cycle with relatively high efficiency (Singal, 2010).
Full Report
Table 2.2.8.1.1: Summary of compressors for refrigerant cycle.

Compressors Advantage Disadvantage
 Low capacity
 Less cost
 High maintenance
Reciprocating  Efficient
 High level of noise
Compressor  Simple controls
and vibration
 Able to control speed
 Smaller & lighter

 High Compression
Screw
ratio  Higher cost
Compressor
 Long continuous
operation
 Prone to surging
 Few moving parts
 Designed for specific
Centrifugal  High vapor flow
pressure differential
Compressor capacity
 Difficult to operate at
 Large capacities
low cooling loads
Table 2.2.8.1.2: Summary of Refrigerants

Refrigeration Advantage Disadvantage
 Inexpensive
Ammonia  Leakage easily  Toxic
(Boiling point, -33°c ) detected  Combustible
 Limited pollution  Large amount
(Osha.gov, 2013) with water required
 Oil-free
R-124 (HCFC)
(Boiling point, -12°c )  Alternative for R-12  GWP of 599
 High ambient air  Causes ozone
(Worldwide Refrigerant cooling depletion
Supplier, 2004, )
Full Report
The ammonia liquid product exiting the flash vessel will be send straight to storage at its
liquefied form and low temperature. The product will be stored in an atmospheric storage
tank at -30°c. This form of cryogenic storage is suitable for liquefied gasses with capacities
of 2000MT and higher (Lele, 2008). A single wall tank is sufficed to hold the liquid at low
temperature under normal operating conditions. To minimize heat leakage, insulation is
added to the external surface. Anhydrous ammonia is corrosive to metals such as copper and
zinc and thus most common material used for construction of tank and piping’s is steel.
Full Report
CHAPTER 3 | MASS AND ENERGY BALANCES

In this chapter the mass and energy balance is carried out in order to verify the results
obtained from the simulations performed on the Excel spreadsheets (Solvers) and Pro II
model of the processing plant. Followed by the mass and energy balance, the overall product
yield is determined as well the utility requirements and generations and waste production.
3.1 Pre-treatment of Biomass

The pre-treatment of the Oil Palm Trunk (OPT) biomass before being sent to the gasifier
consists of two process: size reduction through screw mill, SR-101 and the removal of
moisture by conveyor belt dryer, DE-101.
3.1.1 Screw Mill (SR-101)

The mass balance is done for this size reduction process and there are several assumptions
were made and listed below:
 The system operates at steady state and the operating temperature and pressure are
similar to ambient condition of 1 atm and 25 and remain constant.
 The loss of biomass during the process is negligible.
The basis of feed that were used 13186 kg/h of OPT biomass with moisture content of 10%.
The basis was selected based on the required syngas to produces the required production of
hydrogen for the synthesis of ammonia as per the product specification. Since there is no loss
of biomass occurring at this stage, the outlet mass flow rate of the biomass will be equal to
the inlet which is 13186 kg/h.
3.1.2 Conveyor Belt Dryer (DE-101)

The conveyor belt as mentioned in Chapter 2, utilize the heat from waste heat recovered from
exhaust of process heating in other facility such as the flue gas from combustor chamber of
the gasifier. Since the flue gas leaving the gasifier is warm, it will be used to heat to the inlet
air into the dryer resulting in moisture removal of the biomass (Roos, 2008)The assumptions
made during the mass balance for the dryer are as follow:
Full Report
 The system operates at steady state and the operating temperature and pressure are
350 and 1 atm respectively and remains constant throughout the process.
 The moisture discharge percentage is assumed to be at 45% (Roos, 2008)
 There is no loss of dry biomass
 Emission of VOCs from the Conveyor belt dryer is relatively lower and thus assumed
negligible.
The basis of feed that were used 13186 kg/h of OPT biomass with moisture content of 10%.
The feed then leaves the dryer with moisture content of 5.89%. Due to loss of moisture the
overall mas flow rate of biomass leaving the dryer will be 12557 kg/h. The preheated air inlet
to the dryer is at 545 kg/h.
The diagram below summarizes the mass flow across the 2 pretreatment units:
S12
S1 S2 S3
Conveyor
Screw Mill
Belt Dryer
S13
Figure 3.1.2.1: Pre-treatment process of OPT
Table 3.1.2.1: Mass flow rates across the pretreatment process

Mass Flowrate (kg/h) S1 S2 S3 S12 S13
OPT (Dry Biomass) 11817 11817 11817 - -
Moisture Content 1368 1368 740 - 629
Air - - - 545 545
Total 13185 13185 12557 545 1174
Please Refer to Appendix A1.1 for detailed calculation.
Full Report
3.2 Dual Fluidized Bed Gasifier (G-101)

3.2.1 Mass Balance for Gasifier
The steam-blown Fast Internal Circulating Fluidized Bed (FICFB) consists of 2 chambers of
reactors where the first reactor is the bubbling fluidized bed blown with steam to gasify the
OPT biomass that is being fed in to produce syngas at the temperature 850 (Kirnbauer and
Hofbauer, 2011). The bed material circulates with the resultant char from steam gasification
into the second reactor, combustor consisting of circulating fluidized bed that is blown with
air at 350 to oxidize or burn out the char in order to generate necessary heat for the
gasification. The compositions of the biomass OPT based on ultimate pyrolysis on wet basis
as specified in chapter 1 is as follow:
Table 3.2.1.1: Ultimate Pyrolysis under wet basis for OPT Biomass (Nipattummakul et al., 2011)
Component Weight percentage (%wt)
Moisture, 5.89
Carbon, 36.01
Hydrogen, 4.51
Nitrogen, 1.90
Sulphur, 0.00
Oxygen, 46.18
Ash Content 5.51
Total 100.00
There were few assumptions made in order to approach the mass balance for this section and
the assumptions will be mentioned along the description below. The mass balance for the
gasifier was done by referring to a pilot scale set up of the dual fluid bed gasifier with the
biomass of 40%straw / 60%wood blended pellets as feedstock which has the composition
almost similar with that of the oil palm trunk, OPTs with 10% difference as explained by
Schmid et.al in their article on “Variation of feedstock in a DFBG – Influence on Product
Gas, Tar Content, and Composition” (Schmid et al., 2012). It is assumed that the operating
conditions and the mass balance will be similar to the above mentioned and the values were
scaled up according to the inlet feed. Therefore, the steam that to be fed to the gasifier to the
biomass ratio is determined to be 0.325 based on the similar article. The syngas composition
was also determined based on the literature that was referred earlier and the composition is as
follow:
Full Report
Table 3.2.1.2 Syngas Composition

Component Weight percentage (%wt)
Moisture, 36.02
Hydrogen, 2.22
Carbon Monoxide, 21.01
Carbon Dioxide, 26.42
Methane, 4.80
Ethylene, 1.26
Ethane, 1.17
Propane, 0.40
Nitrogen, 1.26
Oxygen, 0.00
Tar 0.76
Particulates 1.44
Ash 1.01
Char Content 2.24
Total 100.00
The tar can be divided into specific components and to be produced based on the content of
carbon in the biomass of feed (Pfeifer et al., 2010). It is estimated that 27.9 g tar/kg C is
produced according to the pilot scale test runs. Hence, for 12557 kg/h of biomass with
36.01% of Carbon, 126.1537 kg/h of tar is produced and Table below shows the composition
and the mass flow rates of tar in the syngas:
Table 3.2.1.3: Composition of Tar and flow rates in the syngas

Component of Tar Tar Composition Tar Flow Rate
(kg/h)
Phenol, 0.017 2.1446
Styrene, 0.027 3.4061
Indene, 0.575 72.5384
Naphthalene, 0.299 37.7200
Acenaphthalene, 0.048 6.0554
Fluorene, 0.015 1.8923
Phenanthrene, 0.021 2.6492
Total 1.000 126.1537
The particulates, ash and char content in the syngas were then obtained based on composition
of the syngas in the pilot scale gasification for the biomass of 40%straw / 60%wood blended
pellets (Schmid et al., 2012). The data is shown below:
Full Report
Table 3.2.1.4: Particulates, ash, and char content in syngas

Content in Syngas (g/m3) Flow Rate in syngas (kg/h)
Particulates 20.00 239.18
Ash / Dust 14.00 167.43
Char Content 3.11 371.93
Total 778.54
For the combustion chamber, the overall mass balance at steady state can be estimated as
inlet flow rate to be equal to the outlet, the flowrate of the syngas excluding the tar,
particulates, ash, dust and char content:
Based on above equation, the flow rate of syngas excluding the tar, particulates, ash, dust and
char content is equal to 15733.34 kg/h.
Then, the resultant char from steam gasification is then circulated into the combustion
chamber which results in the production of flue gas. Besides the resultant char, the ash
collected from the cyclone, the bed filter of flue gas and the tar oil obtained from the post-
treatment section of the syngas is then supplied to the combustion chamber as fuel.
According to the pilot scale gasification test described in the “Variation of feedstock in a
DFBG – Influence on Product Gas, Tar Content, and Composition” (Schmid et al., 2012), the
flue gas to the syngas from gasifier ratio is 2.705. Similar literature source also mentioned
that the ratio of air to the flue gas ratio is 0.8987. The composition of the flue gas from the
combustion of char at the present of excess air is obtained based on Gasification of Different
kinds of Non-Woody Biomass Dual Fluidized Bed Gasifier Conference Paper by Kitzler et.
Al. (Kitzler et al., 2012). Therefore, the flow rate of flue gas produced and air that is
supplied to the combustion chamber are 46364 kg/h and 40455 kg/h. The table below shows
the composition and the flow rate of flue gas.
Full Report
Table 3.2.1.5: Composition and flow rates of components in the flue gas
Component Weight percentage (%wt) Flow Rate (kg/h)
Moisture, 3.88 1800.55
Hydrogen, 28.16 13054.01
Carbon Monoxide, 36.89 17105.26
Carbon Dioxide, 14.56 6752.08
Nitrogen, 3.88 1800.55
Oxygen, 5.83 2700.83
Nitrogen Oxides, 2.91 1350.42
Argon, Ar 0.97 450.14
Particulates 1.94 900.28
Ash 0.97 450.14
Total 100.00 46364.26
The diagram below shows the overall block diagram of the Gasifier section and the following
table shows the mass flow rates.
S4 S5
S3
Combustor
Gasifier Chamber
Chamber
S27a S11
Figure 3.2.1.1: Block Diagram of Gasifier

Full Report
Table 3.2.1.6: Mass flow rates across the Gasifier

Mass Flowrate (kg/h) S3 S27a S5 S4 S11
OPT (Dry Biomass) 11817 - - - -
Air - - - - 40455.48
Moisture, 739.61 4081.03 1800.55 5993.27 -
Hydrogen, - - 13054.01 369.59 -
Carbon Monoxide, - - 17105.26 3496.08 -
Carbon Dioxide, - - 6752.08 4395.07 -

Methane, - - - 799.10 -
Ethylene, - - - 209.76 -
Ethane, - - - 194.78 -
Propane, - - - 65.93 -
Nitrogen, - - 1800.55 209.76 -
Oxygen, - - 2700.83 - -
Nitrogen Oxides, - - 1350.42 - -
Argon, Ar - - 450.14 - -
Tar Content - - - 126.15 -
Particulates - - 900.28 239.18 -
Ash/Dust - - 450.14 167.43 -
Char Content - - - 371.93 -
Total 12557 4081.03 46364.26 16638.03 40455.48
3.2.2 Energy Balance across Dual Fluidized Bed Gasifier

The energy balance is done for the gasifier chamber in order to determine the amount of heat
energy that need to be supplied by the combustion chamber. The major assumption made
during the hand calculations are:
 The system is operating at steady state, therefore, the operating condition remain
constant throughout the process. The operating temperature is at 840 and at 5 bar.
 References for the elemental species that form the reactants and products were
chosen to be at 25 and at 1 atm (the state for which heats of formation are known)
and the non-reactive species are also at 25 and at 1 atm.
 The syngas is assumed to be ideal gas whereby the difference in pressure is
negligible.
 No loss of heat to the surrounding for the combustor region.
 Effects of any pressure changes on the enthalpies are neglected.
Full Report
The inlet-outlet enthalpy table is shown below:

Table 3.2.2.1: Inlet-outlet enthalpy table
Components
OPT (Dry Biomass) 12557 (kg/h) - -
Steam, 226.72 - -
Hydrogen, - - 184.7926
Carbon Monoxide, - - 124.8599
Carbon Dioxide, - - 99.8879
Methane, - - 49.9439
Ethylene, - - 7.4916
Ethane, - - 6.4927
Propane, - - 1.4983
Nitrogen, - - 7.4916
Tar Content
Phenol, - - 0.0228
Styrene, - - 0.0328
Indene, - - 0.6253
Naphthalene, - - 0.2947
Acenaphthalene, - - 0.0398
Fluorene, - - 0.0114
Phenanthrene, - - 0.0149
Particulates - - 239.18 (kg/h)
Ash/Dust - - 119.28 (kg/h)
Char Content - - 371.93 (kg/h)
The each unknown specific enthalpy is then calculated. For the reactant and the product, the
calculation was started with the elemental species to be at 25 and 1 atm and form 1 mol of
the process species at 25 and 1 atm whereby the enthalpy of formation, is obtained
from the Table A-26 from “Thermodynamic: An Engineering Approach”(Cengel and Boles,
2010) and for certain components the enthalpy of formation, were obtained from
Chapter 2 of Perry’s Chemical Engineering Handbook (Poling et al., 2008). The species were
then brought from 25 to its process state, calculating the sensible heat enthalpy using the
appropriate heat capacities obtained from Table A-2 (Cengel and Boles, 2010) and from
Chapter 2 of Perry’s Chemical Engineering Handbook (Poling et al., 2008). The specific
enthalpy that is keyed in the inlet-outlet table is the sum of the enthalpy changes for each step
in the process step. The overall is then calculated for the gasifier as below:
Full Report
The calculated value of is then substituted in the energy balance equation to obtain the Q
required. By which with changes in potential, kinetic and shaft work in negligible, the open
system energy balance gives . The enthalpy inlet-outlet table finaaly appears as:
Table 3.2.2.2: Completed Inlet-outlet enthalpy table

Components
OPT (Dry Biomass) 12557 (kg/h) (kJ/kg) - -
Steam, 226.72 -232 332.9596 -211
Hydrogen, - - 184.7926
Carbon Monoxide, - - 124.8599
Carbon Dioxide, - - 99.8879

Methane, - - 49.9439
Ethylene, - - 7.4916
Ethane, - - 6.4927
Propane, - - 1.4983
Nitrogen, - - 7.4916
Tar Content
Phenol, - - 0.0228
Styrene, - - 0.0328
Indene, - - 0.6253
Naphthalene, - - 0.2947
Acenaphthalene, - - 0.0398
Fluorene, - - 0.0114
Phenanthrene, - - 0.0149
Particulates - - 239.18 (kg/h)

Ash/Dust - - 119.28 (kg/h)
Char Content - - 30.994
Total Ethalpy, -2.70E+08 -1.11E+08
The overall is then calculated for the gasifier:
Therefore, the amount of heat energy that to be supplied by the combustor to the gasifier
is . Please refer to Appendix A2.2 for further detailed calculation.
Full Report
3.3 Post-Treatment of Syngas

3.3.1 Cyclone
The syngas from gasifier at temperature of and , S107 enters the cyclone FG-101
for the dust particulates and the ash to be removed from the syngas. The condition of the
syngas entering the cyclone is given in Table 3.3.1.1 in which the values were obtained from
the HYSYS and the anticipated particle size distribution in the inlet syngas is given in Table
3.3.1.2 below.
Table 3.3.4.1: Syngas Condition entering Cyclone

Total Mass Flow Rate,
Pressure,
Temperature,
Gas Density,
Gas Viscosity,
Gas Volumetric Flowrate,
Particulate Density,
Table 3.3.1.2: Anticipated Particle Sizes in Syngas Inlet
Particle Size ( Cumulative in Range

Min Max
0 2 2 2
2 5 5 3
5 10 15 10
10 20 25 10
20 30 40 15
30 40 55 15
40 50 75 20
50 70 100 25
The efficiency of the cyclone is assumed to be 89%. Therefore, it is assumed that 90% of the
ash, particulates, and char content in the gas will be removed by the cyclone, in which will
results in a very negligible amount of particulates in the syngas. The table below shows the
mass flow rates across the cyclone.
Full Report
S19
S4 Cyclone
S18
Figure 3.3.1.1: Block Diagram of Cyclone
Table 3.3.1.3: Mass flow rates of streams across cyclone

Mass Flowrate (kg/h) S4 S19 S18
OPT (Dry Biomass) - - -
Air - - -
Moisture, 5993.27 5993.27 -
Hydrogen, 369.59 369.59 -
Carbon Monoxide, 3496.08 3496.08 -
Carbon Dioxide, 4395.07 4395.07 -

Methane, 799.10 799.10 -
Ethylene, 209.76 209.76 -
Ethane, 194.78 194.78 -
Propane, 65.93 65.93 -
Nitrogen, 209.76 209.76 -
Oxygen, - - -
Nitrogen Oxides, - - -
Argon, Ar - - -
Tar Content 126.15 126.15 -
Particulates 239.18 28.7017 210.48
(Trace)
Ash/Dust 167.43 20.0912 147.34
(Trace)
Char Content 371.93 44.6312 327.30
(Trace)
Total 16638.03 15952.92 685.11
Full Report
The stream S18, which are the ash, particulates and chars that were filtered from the syngas
are sent to the combustion chamber of the gasifier where it will be used as fuel for the
combustion. The particulates, ash and char content in the treated syngas is very low. It’s
assumed that there is only negligible temperature drop across the cyclone and thus, energy
balance is not required.
3.3.2 Waste Heat Boiler (WHB-101)

The Stream S19 syngas leaving the cyclone enters the Waste Heat Boiler at temperature
840 and 5 bar, WHB-101 where the heat from the syngas will be recovered to produce
superheated steam that may be used for the power generation. The syngas need to be cooled
to temperature of 204 and not below before being sent to the scrubber for tar removal to
prevent the condensation of tar in the syngas while being processed in the waste heat boiler.
Energy balance was done in order to determine the required mass flow rate of high pressure
water at 50 bar to reduce the temperature of the syngas from 840 to . Following are
the assumptions that were considered during the hand calculation:
 The system is operating at steady state, therefore, the operating condition remain
constant throughout the process.
 The ash, particulates, and char content in the syngas at S19 is very low and thus their
effect on the heat transfer process is assumed to be negligible and not taken into
consideration.
 All the heat lost by the syngas is completely transferred to the water to be heated.
There is no heat loss to the surrounding from the waste heat boiler.
 The syngas is assumed to be ideal gas whereby the difference in pressure is
negligible.
 Effects of any pressure changes on the enthalpies are neglected.
 Negligible changes in potential, kinetic and shaft work in the waste heat boiler.
Therefore, the energy balance across the waste heat boiler gives:
The Waste Heat Boiler is designed to be consisting of 3 sections, a super heater, kettle
evaporator, and an economizer operating at a single pressure. Based on the suggested WHB
temperature profile by V.Ganapathy in his article titled “Heat Recovery Steam Generators:
Full Report
Understand the Basics”(Ganapathy, 2001), the range of pinch and approach point is
and respectively for syngas inlet temperature ranging from to
. Therefore, the temperature profile that is used in this design based on the energy and
mass balance is as shown in figure below:
Superheater Kettle Evaporator Economizer
Figure 3.3.2.1: Temperature Profile
The total heat lost from the syngas, S19, or in other words, the duty of the waste heat boiler is
calculated as follow:
Where the change in the enthalpy, for each component of the syngas is calculated by
using the appropriate heat capacity equation and respective constants based on the Appendix
C: Physical Property Data Bank (Sinnott and Towler, 2009) as shown below:
Based on the calculation, the total heat lost from the syngas is obtained as
. The negative sign indicates the loss of heat to the steam.
Next, the enthalpy change in the water stream is calculated. The compressed water, S24 is
supplied to the waste heat boiler at 100 at 50 bar. The superheated steam leaving the waste
heat boiler is set to leave at temperature of 700K or 427 . The calculation was done in
three sections as the heat is transferred to the steam in 3 different sections, super heater and
Full Report
economizer where the sensible heat enthalpy change is considered and the evaporator where
besides sensible heat, latent heat is taken into consideration due to phase change. The
enthalpy changes at all 3 sections were summed up as to obtain the molar flow rate of
the water required as shown below:
Based on the calculation,
Please refer to Appendix A3.2, for detailed calculation.
PRO-II was used to simulate the waste heat boiler as 3 heat exchanger representing super
heater, evaporator and the economizer based on the temperature profile above in order to
check the mass and energy balance calculations conducted by hand calculations. The values
obtained from PRO-II were compared with the hand calculation values to use them as a
means of justification. The start-up, maintenance or shut-down processes were not taken into
consideration during the simulation. The PRSV fluid package is the package used for the
simulation. The table below shows the comparison between both PRO-II and hand calculated
values. The percentage difference is calculated as follow:
Table 3.3.2.1: Comparison of values from Hand Calculation and PRO-II

Parameters Unit Hand PRO-II Percentage
Calculations Difference
(%)
Duty, 5762 0.02
Molar Flow rate of Compressed 433.761 406.681 6.24
water,
Mass Flow rate of Compressed 7807.692 7320.251 6.24
water,
These small percentage differences between the values are due to certain limitations in both
the methods to solve the overall material and energy balances. For instance, the syngas was
assumed to be an ideal gas for the ease of hand calculations and thus the effect of pressure in
the enthalpy calculation was ignored.
Full Report
However, equation of state, EOS method is used to estimate the actual enthalpies in PRO-II.
The pressure drop in both shell and tube side of the 3 sections in the waste heat boiler is
being taken into consideration for the waste heat boiler. However, for hand calculation, the
pressure drop is assumed to be negligible and the effect on the duty and enthalpy of the flow
is also assumed to be very minor and thus negligible. This explains the discrepancies in the
values calculated.
S19 S22
Waste Heat Boiler (WHB-101)
S27 S24
Figure 3.3.2.2: Waste Heat Boiler Block Diagram
Table 3.3.2.2: Streams mass flow rates across waste heat boiler
S19 S22 S24 S27
Temperature, 840 204 100 427
Pressure, 5 5 50 50
Mass Flowrate (kg/h)
Syngas 15952.92 15952.92 - -
- - 7320.251 7320.251
3.3.3 Tar Removal Process
3.3.3.1 Mass and Energy Balance across the Heat Exchanger

The cooled syngas from the waste heat boiler S22 at the temperature of 204 , will then
passes through a heat exchanger, HX-103 that cools the gas to 65 to condense the tars
present in the syngas. The total heat lost from the syngas, S22, or in other words, the duty of
the heat exchanger is calculated as follow:
Where the change in the enthalpy, for each component of the syngas is calculated by
using the appropriate heat capacity equation and respective constants based on the Appendix
C: Physical Property Data Bank (Sinnott and Towler, 2009) as shown below:
Full Report
Based on the calculation, the total heat lost from the syngas is obtained as
. The negative sign indicates the loss of heat to the steam. Assuming no heat loss
to the surrounding and the heat loss by the syngas is equal to heat gained by the water. The
enthalpy change in the water stream, is calculated. The cooling water, S30 is supplied to
the waste heat boiler at 25 at 1 bar and is set to leave at temperature of 50 . The required
flowrate of cooling water,
Based on the calculation,
Please refer to Appendix A10.1, for sample calculation.
3.3.3.2 Mass Balance across the Wet Scrubber and Mist Eliminator
The heat exchanger is followed by a Venturi scrubber to remove particulate and condensed
tars. After the Venturi scrubber, the syngas passes through a mist eliminator that cools the
syngas further to 40 to condense the water droplets in the syngas and remove them. Based
on the article entitled “Wet Scrubber Technology for Controlling Biomass Gasification
Emissions” (Bartocci and Patterson, 2007), the wet scrubber is assumed to fully remove the
tar content of the syngas as well as the particulates and 99% of the water content in syngas is
removed in the mist eliminator. The condensed tar is in the form of oil at this temperature
with the water content removed will be sent to the oil/water separator, to separate the water
from the tar oil and recirculate it back to the venture scrubber. The recycled water will be free
of oils that either float or sink. The collected tar oils is sent back to the gasifier.
Full Report
S30 S35
S29
Wet Scrubber and Mist
S22
Heat Exchanger (HEX-103) Eliminator
S30a
S34
S31
Combustion Oil Water S32

Chamber of Separator Pump
Gasifier
S33
Figure 3.3.3.2.1: Block Diagram for Tar Removal Process (Wet Scrubber)
Table 3.3.3.2.1: Streams mass flow rates across Tar Removal Process (Wet Scrubber)
Mass Flowrate (kg/h) /S29
Steam, 5993.27 5933.34 - 5933.34
Hydrogen, 369.59 369.59 - - -
Carbon Monoxide, 3496.08 3496.08 - - -
Carbon Dioxide, 4395.07 4395.07 - - -
Methane, 799.10 799.10 - - -
Ethylene, 209.76 209.76 - - -
Ethane, 194.78 194.78 - - -
Propane, 65.93 65.93 - - -
Nitrogen, 209.76 209.76 - - -
Tar Content
Phenol, 2.1446 - 2.1446 2.1446 -
Styrene, 3.4061 - 3.4061 3.4061 -
Indene, 72.5384 - 72.5384 72.5384 -
Naphthalene, 37.7200 - 37.7200 37.7200 -
Acenaphthalene, 6.0554 - 6.0554 6.0554 -
Fluorene, 1.8923 - 1.8923 1.8923 -
Phenanthrene, 2.6492 - 2.6492 2.6492 -
Particulates 28.7017 - 28.7017 28.7017 -
Ash/Dust 20.0912 - 20.0912 20.0912 -
Char Content 44.6312 - 44.6312 44.6312 -
Total 15952.92 9800 6152.9179 219.5778 5933.34
Full Report
3.4 Autothermal Reformer
49
Conversion Equilibrium 71
Reactor Reactor
70
Figure 3.4.1 Block Diagram of Autothermal Reformer
Autothermal reformer consists of both conversion reaction and equilibrium reaction in a

reactor. Pro II is used in simulating the reactor. The reactor was simulated as 2 reactors which
were conversion and equilibrium reactor in Pro II. The inlet of the reactor is obtained from
outlet of the preheater which is the hand calculation value from post-treatment. With the
assumption of no generation and loss during the compression and preheating, the inlet into
reformer is the mass flow after post-treatment of syngas.
Autothermal reformer is an adiabatic reformer. In this case, heat generated in the conversion
reactor through combustion will be utilised in the equilibrium reactor for reforming reaction.
Assumptions
 All the heat generated is used to heat up the syngas to desired temperature without
any heat loss to the surrounding
 The system is steady state where there is no mass accumulation and generation in the
system (Syngas passing through the fired heater)
 The syngas is assumed to be ideal gas and its difference in pressure is negligible
 The references for each composition that for the reactants and products were chosen
to be at 25 and 1 atm
 The effect of pressure changes on the enthalpies are neglected
Full Report
 The change in kinetic, potential and shaft work is negligible. Thus, the open system
energy balance gives
 Only carbon monoxide and hydrogen involved in the combustion reaction
 All the heat generated in partial oxidation reaction will be fully utilised in equilibrium
reactor
3.4.1 Mass Balance of Conversion Reactor

In conversion reactor, oxygen is limiting reactant in the reactions. Therefore, in order to
determine the amount of air required for the combustion, Goal Seek in Microsoft Excel is
used as an approach to identify the exact amount of air required. It is done by complementing
the amount of heat required in the equilibrium reactor. Some assumptions are made in order
to fulfil all the condition.
The reactions that involved in the conversion reactors are:
As there are two reactions occur at the same time in the same reactor, the rate of distribution
of oxygen for both reactions are set based on the literature studies. According to Haslam, in
the simultaneous combustion of hydrogen and carbon monoxide, the rate of combustion of
carbon monoxide to rate of combustion of hydrogen is 1:2.86(R.T.Haslam, 1923). Based on
this ratio, the conversion of hydrogen and carbon monoxide in limiting air are 74% and 26%.
After iterations, the mass balance and the operating condition in the conversion reactor are
shown in Table 3.4.1.1. The air inlet tabulated in table is obtained after iterations are done.
Detailed steps on iterations will be discussed in next sections. Brief discussion on the
calculation steps are shown as below:
Step 1: Assuming an amount of air entering the conversion reaction. The conversion for both
reactions between oxygen with carbon monoxide and hydrogen are determined based on the
information from literature studies.
Step 2: Mass balance for conversion reactor is carried out.

Full Report
Table 3.4.1.1 Mass Balance of the Conversion Reactor

Composition Inlet Inlet Reacted/Produce Outlet Outlet
(kmol/h) (kg/h) (kmol/h) (kg/h)
Tin: 736.82 Tout: 1186
Hydrogen 183.33 366.65 -60.21 123.11 246.23
Water 637.22 11469.99 60.21 697.43 12553.82
Carbon 124.81 3494.71 -21.05 103.76 2905.21
Monoxide
Carbon Dioxide 99.87 4394.09 21.05 120.92 5320.44
Methane 49.81 796.96 0 49.81 796.96
Ethylene 7.48 209.37 0 7.48 209.37
Ethane 6.48 194.33 0 6.48 194.331
Propane 1.50 65.78 0 1.50 65.78
Oxygen 40.63 1300.26 -40.63 0 0
Nitrogen 160.35 4489.79 0 160.35 4489.70
Total Flow 1311.46 26781.84 1270.83 26781.84
3.4.2 Energy Balance of Conversion Reactor

Energy balance is carried out based on the combustion reaction in conversion reactor. As
mentioned before, the reactions involved are:
Steps that have taken to complete the energy balance calculations are discussed as follow:
Step 1: The specific enthalpy for each composition will be calculated. The calculations will
begin by taking the elemental species to be at 25 and 1 atm and form 1 mol of process
species at 25 and 1 atm. Thus, the species will bring from the reference condition to its
process state. In this case, the entering temperature will be bring down to the reference
temperature and again bring up to the process state. Thus, sensible heat enthalpy and overall
will be calculated based on inlet and outlet temperature obtained from Pro II. The
constants of integration of heat capacities are shown in Table A4.3.
Step 2: Heat of reaction for both partial oxidation reaction will be calculated.
Step 3: The total heat released from the partial oxidation reactions will be calculated by
summing up the overall and the heat of reaction.
Full Report
Table 3.4.2.1 Energy Balance of Conversion Reactor

Energy Balance
Heat of Reaction: -3309.84 kJ/s
Heat of Reaction: -4044.80 kJ/s
overall 6907.79 kJ/s
Total Heat Released from Partial Oxidation -446.85 kJ/s
* The results is based on the air inlet after iterations
Detailed Calculations are shown in Appendix A4
3.4.3 Mass Balance of Equilibrium Reactor

The outlet of conversion reactor will be entering the equilibrium reactor. Steam reforming
reaction will be occurred here by utilising the heat generated from partial oxidation reactions.
Based on literature studies, for higher hydrocarbon, the reactions are assumed to be
irreversible reaction and will be reacting completely. The calculation steps are similar to mass
balance in conversion reactor. Detailed calculations are shown in Appendix A4. The
breakdowns of inlet and outlet stream for equilibrium reactor are shown in Table 3.4.3.1
Table 3.4.3.1 Mass Balance for Equilibrium Reactor

Composition Inlet Inlet Reacted/Produce Outlet Outlet
(kmol/h) (kg/h) (kmol/h) (kg/h)
Tin: 1186 Tout: 950
Hydrogen 123.11 246.23 207.25 330.36 660.73
Water 697.43 12553.82 -77.22 620.21 11163.78
Carbon 103.76 2905.21 77.22 180.98 5067.50
Monoxide
Carbon 120.92 5320.44 0 120.92 5320.44
Dioxide
Methane 49.81 796.96 -44.83 4.98 79.70
Ethylene 7.48 209.37 -7.48 0 0
Ethane 6.48 194.33 -6.48 0 0
Propane 1.50 65.78 -1.50 0 0
Oxygen 0.00 0.00 0 0 0
Nitrogen 160.35 4489.70 0 160.35 4489.70
1270.83 26781.84 1417.80 26781.84
Full Report
3.4.4 Energy Balance of Equilibrium Reactor

Energy balance is carried out based on the combustion reaction in conversion reactor. As
mentioned before, the reactions involved are:
Steps that have taken to complete the energy balance calculations are discussed as follow:
process state. In this case, the entering temperature will be bring down to the reference
temperature and again bring up to the process state. Thus, sensible heat enthalpy and overall
will be calculated based on inlet and outlet temperature obtained from Pro II. The
constants of integration of heat capacities are shown in Table A4.9.
Step 2: Heat of reaction for equilibrium reactions will be calculated.
Step 3: The total heat released from the partial oxidation reactions will be calculated by
summing up the overall and the heat of reaction.
Table 3.4.4.1 Energy Balance of Conversion Reactor

Energy Balance
Heat of Reaction: 2567.22 kJ/s
Heat of Reaction: 0 kJ/s
overall -3388.86 kJ/s
Total Heat Released from Steam Reforming 446.85 kJ/s
* The results is based on the air inlet after iterations
Full Report
After mass and energy balance for both conversion and equilibrium reactors are prepared, the
iterations can be carried out.
Assuming all the heat generated in partial oxidation reaction will be fully utilised in
equilibrium reactor. Goal Seek in Microsoft Excel was done. The required amount of air will
be iterated by equating the heat released from partial oxidation reactions to heat required in
equilibrium reactions.
From iteration, the amount of air required is 193.49 kmol/h.
Detailed Calculations are shown in Appendix A4
3.4.5 Comparison with Pro II value

Table 3.4.5.1 Comparison of calculated inlet of Autothermal Reformer with Pro II
Composition Inlet (kmol/h) Inlet (kmol/h) Error %
(Hand Calc) (Pro II)
Tin: 736.82 Tin: 736.82 0
Hydrogen 183.33 183.33 0
Water 637.22 637.29 0
Carbon Monoxide 124.81 124.81 0
Carbon Dioxide 99.87 99.87 0
Methane 49.81 49.81 0
Ethylene 7.48 7.48 0
Ethane 6.48 6.48 0
Propane 1.50 1.50 0
Oxygen 40.63 42.42 4.3
Nitrogen 160.35 167.09 4.0
Table 3.4.5.2 Comparison of calculated outlet of Autothermal Reformer with Pro II

Composition Outlet (kmol/h) Outlet (kmol/h) Error %
(Hand Calc) (Pro II)
Tout: 950 Tout: 950 0
Hydrogen 330.36 375.32 12
Water 620.21 581.82 8.3
Carbon Monoxide 180.98 145.45 24
Carbon Dioxide 120.92 159.70 24.2
Methane 4.98 1.73 65
Ethylene 0.00 0 0
Ethane 0.00 0 0
Propane 0.00 0 0
Oxygen 0.00 0 0
Nitrogen 160.35 167.09 4.0
Full Report
3.5 Shift Reaction

3.5.1 Mass Balances across High Temperature Water Gas Shift Reactor (HTWGSR) and
Low Temperature Water Gas Shift Reactor (LTWGSR)
The water gas shift reaction is expressed below as an equilibrium-limited, heterogeneous and
exothermic reaction whereby it is thermodynamically favoured at low temperatures (Smith, et
al., 2010).

is no change in the volume from reactants to products (Smith, et al., 2010). This principle is
expressed in the equilibrium equation below:
whereby
3.5.1.1 Basis for mass balance calculations

The initial inlet syngas molar flow rate values were taken from PRO II simulation in order to
verify the values calculated by PRO II. The fluid package used during the PRO II simulation
of this section of the plant was Peng-Robinson-Stryjek-Vera (PRSV) Equation of State which
was published by Stryjek and Vera in 1986. This fluid package was used as it is the improved
version of the Peng-Robinson Equation of State and can accurately represent the relation
between temperature, pressure and phase composition in binary and multicomponent systems
since little computer sources are used (Proust & Vera, 2009). Based on the preceding points,
the PRSV fluid package was chosen for the simulation of this section of the plant.
The operating conditions that might affect the mass balance calculations are as shown:
Full Report
Table 3.5.1.1.1: Operating conditions for HTWGSR and LTWGSR

HTWGSR
Syngas component (Stream 72) Molar flow rate (kmol/h) from PRO II
CO 148.1051479
H2 372.7175155
H2O (v) 584.362975
CH4 1.753382578
CO2 157.008827
N2 166.7183
C2H4 0.000120011
C3H6 0.000133208
C3H8 0.004445382
Total Inlet Molar Flow Rate of Syngas 1430.670847
Inlet Syngas Temperature 350
Outlet Syngas Temperature 350
LTWGSR
Syngas component (Stream 79) Molar flow rate (kmol/h) from hand
calculation values after mass balance
across HTWGSR
CO 15.82942698
H2 504.9932364
H2O (v) 452.0872541
CH4 1.753382578
CO2 289.2845479
N2 166.7183
C2H4 0.000120011
C3H6 0.000133208
C3H8 0.004445382
Total Inlet Molar Flow Rate of Syngas 1430.670847
Inlet Syngas Temperature 200
Outlet Syngas Temperature 200
The operating temperatures in HTWGSR and LTWGSR are chosen based on the
Ferrochrome catalyst used in the HTWGSR unit which operates at a temperature range of
330 to 500 and the copper based catalyst used in the LTWGSR unit which operates at a
range of 200 to 250 (Smith et al., 2010).
Full Report
3.5.1.2 Assumptions for mass balance calculations

 Reactions in HTWGSR and LTWGSR reached equilibrium
 No pressure drop across HTWGSR and LTWGSR
 HTWGSR and LTWGSR operate at steady state and isothermally
 The extent of reaction and rate of conversion was only governed by the equilibrium
constant
3.5.1.3 Calculation steps for mass balance

Step 1: Determine the value of equilibrium constant Keq for the reactor to reach the
equilibrium
Since the water-gas shift reaction reaches equilibrium at 350 for the HTWGSR and at
200 for the LTWGSR, the equilibrium Keq values for both reactors can be calculated using
the equation shown in section 3.5.1.
The value of equilibrium constant Keq for HTWGSR is given by:
Meanwhile, the value of equilibrium constant Keq for LTWGSR is given by:
Step 2: Determine the mole fractions of the components present in the syngas inlet
The mole fractions of the components present in the syngas inlet can be determined by
substituting the molar flow rate values from Table 3.5.1.1.1 into the equation below.
Full Report
Table 3.5.1.3.1: Mole fraction of components present in the syngas inlet of HTWGSR
Syngas Inlet Component Mole composition
(Stream 72)
N2 0.116531556
H2 0.260519404
Steam 0.408453822
CO 0.103521469
CO2 0.109744899
CH4 0.001225567
C2H4 8.38847 10-8
C3H6 9.31085 10-8
C3H8 3.1072 10-6
Total 1
Table 3.5.1.3.2: Mole fraction of components present in the syngas inlet of LTWGSR
Syngas Inlet Component Mole composition
(Stream 79)
N2 0.116531556
H2 0.352976534
Steam 0.315996692
CO 0.011064339
CO2 0.202202029
CH4 0.001225567
C2H4 8.38847 10-8
C3H6 9.31085 10-8
C3H8 3.1072 10-6
Total 1
Step 3: List down the all molar flow rates in the outlet in terms of extent of reaction
The molar flow rates for each product species were determined firstly by listing out the
expressions for each product species molar flow rate in terms of extents of reaction using the
equation shown below.
Full Report
whereby
HTWGSR
Figure 3.5.1.3.1: Summary of molar flow rates in the product stream in terms of extent of reaction for HTWGSR
Full Report
LTWGSR
Figure 3.5.1.3.2: Summary of molar flow rates in the product stream in terms of extent of reaction for LTWGSR
Step 4: Express the mole fractions of the products in outlet stream in terms of extent of
reaction at equilibrium
The expressions for mole fractions of the product in HTWGR are as shown below:
; ;
; ; ;
; ;
Meanwhile, the expressions for mole fractions of the product in LTWGR are:
The expressions for mole fractions of the product in HTWGR are as shown below:
; ;
; ;
Full Report
; ;
Step 5: Substitute the mass fraction expressions into the equation for equilibrium
constant, Keq
The mass fraction expression shown in step 4 was then substituted into the equilibrium
constant equation below
Step 6: Determine the extent of reaction,
The extent of reaction can then be determined by using the goal seek function in Microsoft
Excel whereby the equilibrium constant in step 5 was set to reach the targeted Keq value in
step 1 by iterating the extent of reaction, .
Therefore, the extent of reaction for HTWGSR was calculated to be 297.1163596 kmol/h
while the extent of reaction for LTWGSR was calculated to be 35.89437492 kmol/h.
Subsequently all molar flow rates in the outlet stream can be automatically calculated by
Excel as shown:
Table 3.5.1.3.3: Molar flow rate of components present in the syngas outlet of LTWGSR and HTWGSR
Syngas Product Component Molar flow rate, kmol/h
Stream 78 Stream 84
N2 166.7183 166.7183
H2 504.9934856 519.1069281
Steam 452.0870049 437.9735624
CO 15.82917778 1.715735271
CO2 289.2847971 303.3982396
CH4 1.753382578 1.753382578
C2H4 0.000120011 0.000120011
C3H6 0.000133208 0.000133208
C3H8 0.004445382 0.004445382
Total 1430.670847 1430.670847
Full Report
Step 7: Changing the molar flow rates of the components present in the syngas outlet to
mass flow rates
Since the molar flow rate of each individual component in the syngas outlet was known, the
mass flow rate of each individual component can then be determined by multiplying the
molar flow rate of each individual component with its corresponding molecular weight
(kg/kmol).
The results obtained from this step is as shown:
Table 3.5.1.3.4: Mass flow ra6e of components present in the syngas outlet of LTWGSR and HTWGSR
Syngas Product Component Mass flow rate, kg/h
Stream 78 Stream 84
N2 4668.1124 4668.1124
H2 1009.986971 1038.213856
Steam 8137.566088 7883.524123
CO 443.2169778 48.04058759
CO2 12728.53107 13349.52254
CH4 28.05412125 28.05412125
C2H4 0.003360317 0.003360317
C3H6 0.003996227 0.003996227
C3H8 0.19559681 0.19559681
Total 27015.67058 27015.67058
Step 8: Calculating the overall percentage conversion of CO
The overall percentage conversion of CO can be calculated using the formula below:
Therefore, the conversion of CO in HTWGSR was found to be 87.8645014% while the

conversion of CO in LTWGSR was found to be 87.46930063%. The total conversion across
the whole water-gas shift reaction section was calculated to be 98.47933715%.
Full Report
3.5.1.4 Verification of the results obtained using PRO II simulation with hand calculation
values
The values obtained via PRO II were compared with the hand calculation values in order to
justify and check the mass balance calculations performed.
Table 3.5.1.4.1: The comparison between the flows obtained using hand calculation and PRO II
Syngas Product Molar flow rate through PRO II simulation (kmol/h)
Component Stream 48 Stream 84
N2 166.7183 166.7183
H2 504.9932364 519.107008
Steam 452.0872541 437.9734822
CO 15.82942698 1.715655221
CO2 289.2845479 303.3983196
CH4 1.753382578 1.753382578
C2H4 0.000120011 0.000120011
C3H6 0.000133208 0.000133208
C3H8 0.004445382 0.004445382
Total 1430.670847 1430.670847
Comparing Table 3.5.1.4.1 and Table 3.5.1.3.3, it can be said that there is very little
difference between the values obtained using hand calculation and PRO II as the percentage
difference calculated is approximately zero.
Full Report
3.5.2 Energy Balance across High Temperature Water-Gas Shift Reactor (HTWGSR) and
Low Temperature Water-Gas Shift Reactor (LTWGSR)
Circulating water is used to keep the reactors isothermal (constant temperature), whereby all
the heat released from the exothermic reaction is used to raise the temperature of the cooling
water.
3.5.2.1 Basis for energy balance calculations

 The cooling water is supplied at 25 and 1atm
 The exit temperature of water is 70 and it is sent to the cooling tower for
recirculation purpose
 The specific heat capacity of water is obtained from Table B.2 of the book entitled
“ Elementary Principles of Chemical Processes” (Felder & Rousseau, 2005)
3.5.2.2 Assumptions for energy balance calculations

 The system is at steady state
 The work, kinetic energy and potential energy are negligible in the system
 The syngas in the system is assumed to be acting as an ideal gas
 The mass of cooling water used is conserved
 All heat released from the exothermic reaction is absorbed by the cooling water
 No pressure drop across the shell side of the reactor
The energy balance for open system is reduced to
3.5.2.3 Calculation steps for energy balance

Step 1: Determine the enthalpy change associated with the extent of reaction
Since the values for HTWGSR and LTWGSR are already known from Section 3.5.1.3 of
Chapter 4, the enthalpy associated with the extent of reaction can be calculated using the
equation shown below.
The total enthalpy from the exothermic reaction in HTWGSR was 1509.783225kW while the
total enthalpy from the exothermic reaction in LTWGSR was 161.0892646kW.
Full Report
Step 2: Determine the molar flow rate of cooling water required

Since the specific heat capacity of cooling water, at 25 and 1atm is given by
75.4kJ/mol K, the molar flow and the enthalpy change associated with the extent of reaction
is already known, the molar flow rate of cooling water can be calculated as shown.
HTWGSR
Figure 3.5.2.3.1: Summary of molar flow rates of the cooling utility supplied to HTWGSR
LTWGSR
Figure 3.5.2.3.2: Summary of molar flow rates of the cooling utility supplied to LTWGSR
Table 3.5.2.3: Summary of molar flow rates of cooling water into and out of the reactor
Stream Molar flow rate (kmol/h)
76 1601.892016
77 1601.892016
82 170.9169917
83 170.9169917
Full Report
3.6 Carbon Dioxide (CO2) Removal
3.6.1 Mass Balance across Carbon Dioxide Removal Section

Carbon dioxide removal section consists of two main operation units, namely the absorption
column (AC-401) and stripping column (SC-401). The mass balance across both absorption
column and stripping column focuses on the application of law of conservation of mass. As
stated by the law of conservation of mass, mass cannot be created or destroyed (Felder &
Rousseau, 2005). The mass that enters a system can either leave the system or accumulate
within the system by conservation of mass. For a continuous system operating at steady state
condition, the mass balance can be represented by the general equation as shown below:
Due to the complexity of the CO2 absorption process using amine solvent, the process is not
able to be simulated. Mass balance in CO2 removal section is carried out based on appropriate
assumptions and literature data.
3.6.1.1 Basis for Mass Balance Calculations

The amine solvent used in this project is piperazine-activated MDEA (aMDEA), is mostly
made up of MDEA (tertiary amine) and water, as well as small amount of piperazine (cyclic
diamine) as promoter. As stated in journal by Yildirim et al. (2012), for tertiary amine, the
maximum CO2 loading is 1mol CO2 / 1mol amine. CO2 loading of aMDEA is not affected by
the piperazine (PZ), since PZ only acts as a promoter. Thus, the theoretical maximum CO 2
loading can be taken as 1mol CO2 / 1mol amine.
3.6.1.2 Assumptions for Mass Balance Calculations across Absorption Column

 The treated syngas stream (sweet gas) contains 100ppm of CO2, as reported in the
journal by Combs and McGuire (2007).
 Only CO2 is being absorbed into the amine solvent
 Lean amine stream recycling back to absorption column does not contain CO2.
Full Report
3.6.1.3 Calculation Steps for Mass Balance across Absorption Column

Step 1: Determine the flow rate of syngas stream entering absorption column(Stream 89)
The syngas stream entering the absorption column is from the stream exiting the separator (S-
401) at upstream. After removing most of the water content, the syngas stream will exit from
the overhead vapour outlet of the separator and enter the absorption column from the bottom.
The results are summarized in the table below:
Table 3.6.1.3.1: Flow Rate of Syngas Stream Entering the Absorption Column
Components Molar Flow Rate (kmol/hr) Mole Fraction
CO2 303.3993 0.3040
N2 166.7217 0.1671
H2 519.1063 0.5202
H2 O 5.2006 0.0052
CO 1.7157 0.0017
CH4 1.7534 0.0018
TOTAL 997.8969 1
Step 2: Determine the total amount of amine solvent used
By using the theoretical maximum CO2 loading of the aMDEA solvent (1mol CO2 / 1mol
amine), the mole of amine (mol of MDEA + mol of PZ) can be easily calculated. Next, taking
a basis of 1 kg/hr of aMDEA solvent, the mole fraction of MDEA, PZ and H2O can be
determined. Lastly, the total molar flow rate of aMDEA solvent used as well as the
component molar flow rate is calculated. The results calculated are summarized in the table
below.
Table 3.6.1.3.2: Summary of Results Calculated

Component Mole Fraction Mole Flow Rate (kmol/hr)
MDEA 0.0973 258.6682
PZ 0.0168 44.7311
H2 O 0.8859 2354.5864
TOTAL 1 2657.9857
Full Report
Step 3: Determine the flow rate of the treated syngas (sweet gas) stream (Steam 107)
As stated in the journal prepared by Combs and McGuire (2007), for CO2 removal process
using amine solvent, the concentration of CO2 in the sweet gas stream is usually less than
100ppm. Thus, it is assumed that the concentration of CO2 in the treated syngas stream is
100ppm. Based on Denn (2012), parts per million (ppm) is expressed as mole fraction for
gases. This means that the amount of CO2 in the treated syngas stream has a mole fraction of
0.0001.
In order to determine the flow rate of treated syngas stream exiting from the top of the
absorption column, the flow rate of other components that is not absorbed (V’) is first
calculated.
Next, the amount of treated syngas leaving the absorption column can be calculated as shown
below:
The amount of CO2 that present in the stream can be calculated by multiplying the flow rate
of treated syngas with the mole fraction, 0.0001.
Besides CO2, the molar flow of other components in the stream is the same as that in dry
syngas stream. The molar flow and mole fraction of all the components in treated syngas
stream is summarized in the table below:
Table 3.6.1.3.3: Flow Rate of Different Components in Treated Syngas Stream

Components Molar Flow Rate (kmol/hr) Mole Fraction
CO2 0.0695 0.0001
N2 166.7217 0.2400
H2 519.1063 0.7474
H2 O 5.2006 0.0075
CO 1.7157 0.0025
CH4 1.7534 0.0025
Full Report
Step 4: Determine the flow rate of the rich amine stream (Stream 90)
CO2 is absorbed by the amine solvent and left the absorption column from the bottom, which
is known as rich amine. As reported by Mohamed & Javed (2010), only trace amount of other
components is observed in the rich amine stream. Thus, it is safe to assume that only CO2 is
absorbed by the amine solvent. The flow rate of rich amine can be easily calculated by
performing overall mass balance across the absorption column.
The total amount of CO2 being absorbed by the amine solvent can be calculated from the
difference between the amount CO2 entering the absorption column and the amount of CO2
present in the treated syngas stream. Then, the mole fraction of CO2 present in the stream can
be calculated.
It is assumed that all the amine solvent entering the absorption column will exit at the rich
amine stream after absorbing CO2. Thus, the molar flow of the solvent in rich solvent stream
is the same as that of amine solvent stream. The molar flow and mole fraction of all the
components in rich amine stream is summarized in the table below:
Table 3.6.1.3.4: Flow Rate of Different Components in the Rich Amine Stream
Component Molar Flow Rate (kmol/hr) Mole Fraction
MDEA 258.6682 258.6682/2961.3141 = 0.0874
PZ 44.7311 44.7311/2961.3141 = 0.0151
Water 2354.5864 2354.5864/2961.3141 = 0.7951
CO2 303.3298 0.1024
Full Report
3.6.1.4 Assumptions for Mass Balance Calculations across Stripping Column

 Concentrated CO2 stream contains CO2 with 0.995 mole fraction
 Amine solvent lost in the stripping column from the concentrated CO2 stream is
replaced by making-up with fresh solvent of the exact amount
 All CO2 is desorbed from the amine solvent and leave the stripping column in
concentrated CO2 stream
3.6.1.5 Calculation Steps for Mass Balance across Stripping Column

Step 1: Determine the flow rate of rich amine stream entering stripping column (Stream
92)
The rich amine stream exiting from the absorption column will pass through the amine-
amine heat exchanger to be heated up before entering the stripping column. Thus, the flow
rate of rich amine stream is the same as that leaving the absorption column.
Step 2: Determine the flow rate of concentrated CO2 stream (Stream 93)
The purity of CO2 that can be achieved by using aMDEA solvent is 99.5% (Kunjunny, et al.,
1999). Thus, this also means that the mole fraction of CO2 in concentrated CO2 stream is
0.995. On top of that, in real industrial CO2 removal process, the amount of CO2 that is not
desorbed is very minimal to of no concern. Thus it is safe to assume that all CO2 is desorbed
from the amine solvent and leave from concentrated CO2 stream.
By performing overall material balance on CO2 across the stripping column, the flow rate of
concentrated CO2 stream can be calculated.
Since it is assumed that all CO2 is desorbed from the amine solvent and leave from
concentrated CO2 stream, the molar flow rate of CO2 in concentrated CO2 stream is the same
as the feed stream. Thus, the total amount of amine solvent lost in the concentrated CO2
stream can be calculated by finding the difference between the total flow rate of concentrated
CO2 stream and the flow rate of CO2 absorbed. The results are summarized in the table below:
Full Report
Table 3.6.1.5.1: Flow Rate of Different Components in Concentrated CO2 Stream

Component Molar Flow Rate (kmol/hr) Mole Fraction
MDEA 0.1483 0.1483/304.8541 = 0.0005
PZ 0.0257 0.0257/304.8541 = 0.0001
Water 1.3503 1.3503/304.8541 = 0.0044
CO2 303.3298 303.3298/304.8541 = 0.9950
Step 3: Determine the flow rate of lean amine stream (Stream 98)
Since it is assumed that all CO2 is desorbed, thus there is only lean amine solvent in this
stream. The total flow rate of lean amine can be calculated by calculating the difference
between the total flow rate of amine solvent in feed stream and the amount of amine solvent
lost in the concentrated CO2 stream.
The component flow rate can be easily calculated by multiplying the total flow rate with the
respective mole fraction. The results are summarized in the table below:
Table 3.6.1.5.2: Flow Rate of Different Components in Lean Amine Stream

Component Mole Fraction Molar Flow Rate (kmol/hr)
MDEA 0.0973 2656.4600×0.0973 = 258.5197
PZ 0.0168 2656.4600×0.0168 = 44.7054
Water 0.8859 2656.4600×0.8859 = 2353.2348
Step 4: Determine the aMDEA make-up rate (Stream 101)
The aMDEA solvent make-up rate is taken to be the same as the aMDEA solvent lost in the
concentrated CO2 stream.
aMDEA solvent make-up = 1.5243 kmol/hr

Full Report
3.6.2 Energy Balance across Carbon Dioxide Removal Section
3.6.2.1 Basis for Energy Balance Calculations

The CP value of component in the inlet and outlet streams of the absorption column is
determined from Table B2 of “Elementary Principles of Chemical Process” (Felder &
Rousseau, 2005), whereas the CP value of MDEA and PZ is obtained from literature source
by Chiu & Li (2008) and Chen et al. (2010).
3.6.2.2 Calculation Steps for Energy Balance across Absorption Column

Step 1: Determine the specific heat capacity (CP) of all the components
The Cp value of each component is shown in the table below:
Table 3.6.2.2.1: CP Values

Component Specific Heat Capacity, CP (kJ/kmol)
Syngas Treated Syngas Lean Amine Rich Amine
CO2 9.414×102 9.437×102 - 9.528×102
N2 7.273×102 7.290×102 - -
H2 7.212×102 7.229×102 - -
H2O 8.432×102 8.453×102 1.885×103 1.908×103
CO 7.277×102 7.295×102 - -
CH4 9.091×102 9.114×102 - -
Ar 1.301×104 1.304×104 - -
MDEA - - 8.125×103 8.224×103
PZ - - 5.160×103 5.225×103
Step 2: Determine the and across the absorption column
Since the absorption column is operating adiabatically, thus the and can be
calculated as shown below:
Full Report
The calculated results are shown in the table below:
Table 3.6.2.2.2: Results Calculated

(kJ/hr) (kJ/hr) Percentage difference (%)
-7.8106×108 -7.8097×108 0.01
3.6.2.3 Assumptions for Energy Balance across Condenser

 Cooling water enters the condenser from the cooling tower at 25°C and 1atm
 Reflux ratio used for the absorption column is 4
 There is no accumulation , loss or generation of mass in the condenser
 The heat released by hot processing fluid is completely transferred to the cold cooling
water
 No heat loss from the condenser to the surrounding
3.6.2.4 Calculation Steps for Energy Balance across Condenser

Step 1: Determine the flow rate of the processing fluid into the condenser
Using the reflux ratio of 4, the total mass flow rate of the processing fluid into the condenser
can be calculated.
Step 2: Determine the , and the percentage difference
Since it is assumed that the heat released by hot processing fluid is completely transferred to
the cold cooling water, the heat absorbed and heat released can be calculated as follow:
Next, the percentage difference between the heat absorbed and the heat released can be
calculated.
Full Report
Table 3.6.2.4.1: Percentage Difference between the Heat Released and Heat Absorbed
(kW) (kW) Percentage Difference (%)
2180.79 2228.50 2.14
3.6.2.5 Assumptions for Energy Balance across Reboiler
 Saturated stream at 3000kPa enters the reboiler and exit as condensate

 There is no accumulation , loss or generation of mass in the condenser
 The heat released by hot saturated steam is completely transferred to the cold
processing fluid
 No heat loss from the condenser to the surrounding
3.6.2.6 Calculation Steps for Energy Balance across Reboiler

Step 1: Determine the flow rate of the processing fluid into the condenser
The total mass flow rate of the processing fluid into the condenser is calculated.
Step 2: Determine the , and the percentage difference
Since it is assumed that the heat released by hot saturated steam is completely transferred to
the cold processing fluid, the heat absorbed and heat released can be calculated as follow:
Next, the percentage difference between the heat absorbed and the heat released can be
calculated.
Table 3.6.2.6.1: Percentage Difference between the Heat Absorbed and Heat Released
(kW) (kW) Percentage Difference (%)
118.25 118.44 0.16
Full Report
3.7 Methanator
3.7.1 Overview of the process and block diagram

The presence of carbon oxides such as carbon monoxide and carbon dioxide can lead to
poisoning of the catalyst downstream of the Methanation in ammonia synthesis reactor; hence
it is of crucial importance to remove these components from the processing stream.
Methanation process converts these compounds to methane and water over nickel catalyst in
an adiabatic packed bed reactor. The reactions occurring in the Methanation section are stated
below:
Methane is considered as an inert in the synthesis reactor and thus does not pose any danger
to the catalyst or cause any risks. The water however does and thus has to be removed from
process stream using a flash separator. The block diagram below shows the process block
diagram. The dotted line represents the overall system boundary for this particular processing
unit.
109
Cooling water
Inlet From Heat Exchanger Methanation

107 108
CO2 Removal HX-501 Reactor (R-501) Outlet entering
115 Ammonia
112 Synthesiser
Condenser (HX- Flash Separator (f-

110 111
502) 501)
113 114
Figure 3.7.1: Simplified Block Diagram representing the equipment’s in the Methanation section, with the dashed line
presenting the system boundary for mass balance
The Table 3.7.5.2 below summarizes the mass balance conducted over this section the
detailed calculations are shown in Appendix A7.
Full Report
3.7.2 Simulating Software and Fluid package

The Methanator reactor was modeled using PRO II equilibrium reactor. Selection of the
appropriate thermodynamic package in PRO II is the key to producing accurate simulations.
The thermodynamic fluid package used for this process was Peng Robinson. Peng Robinson
model provides better representation of vapor pressure of pure components and mixtures.
3.7.3 Assumptions
1. The system is operating at steady state.
2. Adiabatic operation in the Reactor and no heat is lost to the surroundings.
3. Pressure drop was assumed to be negligible.
4. The feed stream is separated into liquid and vapor streams at equilibrium.
5. The reactor effluent mixture entering the flash tank is a vapor-liquid mixture. For this
two phase solution to exist, the flash temperature 27 , lies between the bubble point
and dew point of the mixture.
6. Flash tank operates under adiabatic conditions.
7. The Antoine equation is an applicable correlation.
8. The feed inlet is an ideal mixture; hence Raoult’s law is applicable.
3.7.4 Basis
The mass flow rate of the syngas entering the Methanator reactor is the mass flow of the
syngas exiting the carbon dioxide removal unit in section 3.6, labeled as stream 107 which is
694.5670266 The specification of the inlet feed is specified in Table 3.7.6.1.
3.7.5 Steps for conducting mass balance over the entire system
Step 1: Mass balance involving the Heat exchanger (HX-501)
Heat exchanger 501 was considered to be a single system and the mass balance was taken
over the steady state heat exchanger which has no accumulation, production and
disappearance of mass, it is concluded that the stream entering the heat exchanger and the
stream exiting the heat exchanger have the same flow rate and compositions.
Full Report
Step 2: Mass balance over the reactor
Mass balance was performed over the reactor by considering the system boundary around the
reactor only. The detailed calculations for this process are shown in Appendix A7.
109
Cooling water
107 108
115 Ammonia
112 Synthesiser

110 111
502) 501)
113 114
In the packed bed reactor two independent reactions are happening simultaneously. In order
to calculated the outlet flow rate and composition the extent of reaction method was used.
The extent of reaction of carbon monoxide Methanation and carbon dioxide Methanation
were noted by and respectively
Step 3: Determining the inlet composition
The inlet compositions to the reactor are from the carbon dioxide removal unit the table
below summarizes the feed:
Table 3.7.5.1: Summary of the Feed condition

Inlet
Component
Molar flow Mole fraction Mass Flow rate Mass fraction
CO 1.715665291 0.002470122 36.187 0.00604
CO2 0.069456703 0.0001 133.548 0.0223
H2 519.1062512 0.747381075 1038.910 0.174
H2 O 5.200570606 0.0074875 66.474 0.0111
CH4 1.753393477 0.002524441 27.04 0.00452
N2 166.7216893 0.240036862 4679.36 0.782
Total 694.5670264 1 5981.520 1
Full Report
Step 4: Mole balance for each component
The mole balance for each component present in the system was written in terms of extent of
reaction one and extent of reaction two. Consumption means the extent of reaction has to be
deducted from the inlet value and generation is vice versa. Table A7.2 in appendix
summarizes the calculation of the mass balance for individual components.
Step 5: Composition of streams
The composition of the outlet stream was determined by dividing the individual component
mass balance over the total mass flow rate this expression was written for each component in
terms of extent of reaction .
Step 6: Determining the K equilibrium
The K equilibrium data for both reactions was obtained and plotted and a final equation was
obtained as a function of temperature. The outlet temperature was estimated as an initial
guess (347.006) and hence the for both reaction was obtained. This outlet
temperature was hence obtained through energy balance and the deviation percentage was
calculated to be 13.5% from the initial guess hence the assumption was considered to be
inaccurate.
Step 7: Calculation of extent of reaction
The following relation was used to determine the extent of reaction (M.Felder and
W.Rousseau, 2005).
These two equations were solved simultaneously to give values for extent of reactions. Hence
the Molar flow rate of each compound is determined.
Full Report
Step 8: Conservation of mass
The outlet mass flow rate has to be equal to inlet of mass flow rate, this is the law of
conservation and after each mass balance it has to be checked in order to check if the result
tallies with the law.
The outlet mass flow is summarized in table below
Table 3.7.5.2: Summary of inlet and outlet flows for Methanation unit
Inlet Outlet
Compoun Molecular
Molar Flow
d weight Molar Mass flow Molar flow Mass flow Molar
rate
Composition rate (kg/h) rate (kmol/h) rate (kg/h) Composition
(kmol/h)
1.71566529
CO 28 0.002470122 48.03862815 1.163 10-7 3.2564 10-6 1.67857 10-10
1
0.06945670 5.12826 10- 2.25643 10
CO2 44 0.0001 3.056094932 8 -6 2.75 10-11
3
519.106251 1027.91851
H2 2 0.747381075 1038.212502 513.9592557 0.741803054
2 1
5.20057060 124.492244
18 0.0074875 93.61027091 6.916235781 0.009982279
6 1
1.75339347 55.5049384
CH4 16 0.002524441 28.05429563 3.469058652 0.005006931
7 3
166.721689
N2 28 0.240036862 4668.2073 166.7216893 4668.2073 0.24063125
3
694.567026 5879.17909
Total - 1 5879.179092 691.0662396 1
6 2
The inlet mass flow and the outlet mass flows are equal to 5879.179092 since they tally the
law of conservation of mass is conserved.
Step 9: Mass balance over condenser:
Mass balance over the condenser follows the same concept as the heat exchanger mass in
would be equal to mass out. The cooling water used to absorb the heat of process stream
enters at 25 and exits at 40 .The mass of cooling water requires is stated later in the
energy balance for the condenser.
Full Report
Step 10: Mass balance over separator:
The steps for calculation separator mass balance are discussed in section the mass balance
over the separator is summarized in table below:
109
Cooling water
107 108
115 Ammonia
112 Synthesiser

110 111
502) 501)
113 114
Table 3.7.5.3: Summary of flows and the composition for the Separation unit
(0.04
(0.96 Mole Mole
Inlet flow of the
of Total flowrate in flowrate in
Component total
rate Flow the vapor in the liquid
flow
rate) outlet (115) phase (114)
rate)
CO 0 0 0 - 0 0
CO2 0 0 0 - 0 0
H2 513.6815 1.50713E-07 0.749974412 4976180.51 513.6814874 9.14046E-07
H 2O 7.0551 0.992111152 0.00151569 1.53E-03 1.038144658 6.016974794
CH4 3.5385 7.88931E-10 0.00516622 6548381.699 3.538509653 4.78472E-09
N2 166.7217 6.5473E-06 0.243413435 37177.69601 166.7216554 3.97082E-05
Total 690.9968 1 1 - 27.505 660.125 684.9320174 6.064819231
Full Report
3.7.6 Energy Balance
3.7.6.1 Assumptions
1. System is operating under steady state
2. Shaft work is neglected
3. Kinetic and potential energy have been neglected
4. Shaft work neglected
5. Adiabatic Reactor hence Q= 0
6. Heat capacity constants for the compounds are taken for ideal conditions.
7. In the condenser section only water condenses.
3.7.6.1 Heat exchanger energy Balance

The result of the energy balance is shown in the able below:
Table 3.7.6.1.1: Energy Balance Table for Heat exchanger

(
Compone S107 flow rate
nt A B C D (
3.52930634
CO 30.869 -1.29 10-2 2.79 10-5 -1.27 10-8 7405.58 0.000476574
1
0.20656233
CO2 19.795 7.34 10-2 -5.60 10-5 1.72 10-8 10706.30 1.92935E-05
7
H2 27.143 9.27 10-3 -1.38 10-5 7.65 10-9 7298.54 0.144196181 1052.42094
12.6149645
H2O 32.243 1.92 10-3 1.06 10-5 -3.60 10-9 8732.48 0.001444603
2
5.35503518
CH4 19.251 5.21 10-2 1.20 10-5 -1.13 10-8 10994.75 0.000487054
8
341.035338
N2 31.15 -1.36 10-2 2.68 10-5 -1.17 10-8 7363.93 0.04631158
8
1415.16214
Total
7
(hand calculated value)
(PRO II value)
Percentage deviation = 0.2%
Full Report
3.7.6.2 Energy balance for the reactor

Step 1
The values for each element present in the processing stream by using the following
relation
(A.Cengel and A.Boles, 2007).
Table 3.7.6.1.1 shows the constant values.
Step 2
The energy balance equation for the reactor is presented by the following equation:(M.Felder
and W.Rousseau, 2005).
Using the extent of reaction obtained from mass balance with 1.715665175 ,
The heat of formation of two reactions was calculated by using the heat of formation for each
compound. The heat of formation for carbon monoxide Methanation and carbon dioxide
Methanation are -206 and -165.01 respectively.
Step 3
The variables are substituted inside the reaction and the is obtained in terms of the outlet
temperature .
Step 4
The outlet temperature can be calculated by using goal seek in excel and iterating the outlet
temperature until would be equal to zero. The temperature outlet obtained from PRO II is
347.006 the temperature obtained from the Hand calculation is 300.0177382 .The
percentage deviation from PRO II is calculated to be 13.54105167%.
Full Report
Step 5
There needs to be a comparison between PRO II and hand calculation values.
3.7.6.3 Energy balance for condenser

The processing gas is getting cooled from 83.950 to 25 . Water condenses at 40 hence
there is sensible heat associated with the processing stream as well as latent heat.
The energy required for the condenser was obtained to be .The

obtained from PRO II is 405.9 kJ/s. This implies a percentage deviation 23.2%. Mass flow
rate obtained for the water cooling 42203.26223kg/s. The mass flow rate of cooling water
obtained from PRO II is 23269.631kg/s. Hence the percentage deviation is about
44.86295663%.
3.7.7 Comparison:
The deviation of the mass balance and energy balance from the PRO II values are discussed
below:
For the Methanation reactor molar flow is compared with PRO II values:
Table 3.7.7.1: Flow rate for Methanation unit and percentage deviation from PRO II values
Molar flow
obtained from PRO Percentage
Component Molar Flow
Deviation (%)
II
CO 1.163E-07 0 0
CO2 5.12826E-08 0 0
H2 513.9592557 513.6815387 0.05
H2 O 6.916235781 7.055088887 2.01
CH4 3.469058652 3.538490493 2.00
N2 166.7216893 166.7216893 0
Total 691.0662396 690.996843 0.01
Full Report
There are a few reasons for the deviation of the mass flow rate:
1. The initial error: The feed entering the processing unit is deviated from the PRO II value
and hence there is an initial error prior to Methanation unit. This can lead to deviation of
results from PRO II.
2. The PRO II assumed a conversion of 100 % whereas the process has a conversion of
about 99.9% this may have led to a little bit of discrepancy. The conversion obtained from
hand calculation tallies with the literature sources (Gao et al., 2012).
Table 3.7.7.2: Percentage deviation of separator from PRO II
Molar flowrate of
Molar flowrate in Percentage
Component vapor phase from
the vapor phase Deviation (%)
PRO II
CO 0 0 0
CO2 0 0 0
H2 513.6814874 513.5549713 0.0246
H2 O 1.038144658 0.923225874 11.0696
CH4 3.538509653 3.536483834 0.0573
N2 166.7216554 166.6861164 0.0213
Total 684.9320174 684.7008217 0.0338
The Antoine constants for PRO II differs from the book, this might be one of the reasons for
deviation also the composition found from hand calculation had an initial percentage error
which will change the data obtained.
For the energy balance the specific heat capacity values were taken from (A.Cengel and
A.Boles, 2007) which assumes these heat capacity constant are for gasses under ideal
condition whereas the gasses present in the system are not ideal gas since they operate at a
very high pressure.
Full Report
3.8 Mass and Energy balance: Ammonia Synthesis Section
3.8.1 Mass Balance around the ammonia synthesis rector

*All detailed calculations as described in the key steps below are as attached in the Appendix A8
S-163
Ammonia Synthesis S-125

S-120
Reactor (R-601)
Figure 3.8.1.1: Simplified block diagram for the ammonia synthesis reactor stage with system boundary
(dashed line) for mass balance
Table 3.8.1.1: Mass Flow Rates and break down of the respective streams for the ammonia synthesis
reactor (CVR-001)
Flow Rates (kg/hr) S-120 S-163 S-125
Water, H2O 16.62 8.03E-4 16.62
Hydrogen, H2 1027.11 1826.87 2066.56
Nitrogen, N2 4667.21 7562.34 8554.48
Methane, CH4 56.59 418.97 475.75
Ammonia, NH3 0.00 1892.06 6354.10
TOTAL 5767.53 11700.24 17467.50
Table 3.8.1.2: Summary of the total mass flow rate in and out from the defined system boundary
Flow Rates (kg/hr)
IN OUT
S-120 5767.53 S-125 17467.50
S-163 11700.24
Total IN 17467.77 Total OUT 17467.50
3.8.1.1 Assumptions for Mass Balance

 The ammonia synthesis reactor operates in steady state isothermal conditions to ensure
minimum temperature fluctuations in the catalyst bed and product quality.
 The main chemical reaction that the reactant gasses undergo in the reactor is at
equilibrium condition and based on Haber process where (BBC, 2010) :
Full Report
 Based on the ammonia synthesis reactor design S200 adopted from (Topsoe, 2013), the
ammonia synthesis reactor is a two catalyst bed reactor.
 The operating temperature of the ammonia synthesis reactor is at an average temperature
of 480oC
 The operating pressure of the ammonia synthesis reactor is at an average pressure of 150
bar
 The pressure drop of the streams across the reactor is low and at 5kPa.

3.8.1.2 Basis for Calculations

The mass flow rate of the syngas inlet entering the ammonia synthesis reactor (R-601) is the
mass flow rate of the purified syngas from the methanator (S-120) and the mass flow rate of
recycle quench gas from the unreacted N2 and H2 gas streams of the downstream flash
separation process (S-163) which are of inlet flow rates of 5767.53 kg/hr and 11700.24 kg/hr
respectively.
With the availability of inlet content characteristics entering the reactor, the HYSYS
simulation software is used as a basis to simulate the ammonia synthesis process.
3.8.1.3 Key Steps

Step 1: Determine the mass fraction of the entering syngas and key-in entering mass
fraction of syngas material inlet
Based on the component mass fraction of syngas entering the ammonia synthesis reactor, all
components were entered in the material stream.
Step 2: Selection of equilibrium reactor as model for ammonia synthesis reactor
Based on the isothermal and equilibrium conditions for operating the ammonia synthesis
reactor, equilibrium reactors were selected as the Haber proce
ss is an equilibrium process with possible backward reactions as conversion is not 100%.

Full Report
Step 3: Selection of appropriate fluid package and key in the chemical stoichiometric
values for the chemical reaction
The fluid package of PRSV is selected for the simulation of ammonia synthesis reactor as it
served as a good representation of the vapor pressure of mixture gasses. Additionally, the
chemical stoichiometric value is key-in to the reaction set based on the Haber process
chemical equation mentioned above.
Step 4: Configuration of two equilibrium reactors to model the two inner catalyst bed of
R-601 reactor
Based on the design of ammonia synthesis reactor adopted which is a two bed catalyst reactor
with an inter-bed heat exchanger and the realization of conversion differences in each catalyst
bed in the reactor, the arrangement of the two equilibrium reactor and an inner-bed heat
exchanger is done.
Step 5: Addition of recycle quench gas stream (S-163) into the inter-bed heat exchanger
After step 3, the recycle quench gas stream (S-163) which consist of the unreacted N2 and H2
gasses recovered after the downstream flash separation process can be recycled into the inter-
bed heat exchanger to further increase the conversion of ammonia product gas and prevent
wastage of the unreacted reactant gas. The addition of recycle gas stream (S-163) can be seen
in Figure 3.8.1.3.1 below and further reference to recycle gas stream (S-163) can refer to
Appendix A8 for downstream processing.
Full Report
Figure 3.8.1.3.1 Addition of recycle quench gas stream from downstream flash separation
Step 6: Summarize all the product streams to check for tallying in and out of the
ammonia synthesis reactor
Based on all the product streams entering and exiting the ammonia synthesis reactor, the
tables of material flow properties were displayed and checking is carried out to compare the
sum of all the mass flow rates entering and exiting the ammonia synthesis reactor (R-601)
(refer to Appendix A8 for full details of stream properties as Table 3.8.1.3.1 below). The
total mass flow rate entering and exiting the hot cyclone is conserved or equal.
Table 3.8.1.3.1: Component mass balance for flows entering and exiting the ammonia synthesis reactor (R-601)
Flow Rates (kg/hr) S-120 S-163 S-125
Water, H2O 16.62 8.03E-4 16.62
Hydrogen, H2 1027.11 1826.87 2066.56
Nitrogen, N2 4667.21 7562.34 8554.48
Methane, CH4 56.59 418.97 475.75
Ammonia, NH3 0.00 1892.06 6354.10
TOTAL 5767.53 11700.24 17467.50
Full Report
Step 7: Calculation of total conversion of catalyst bed and comparison on the

percentage difference of the conversion between calculated conversion and the
industrial S200 ammonia synthesis reactor conversion
Since the limiting reagent gas in this case is the Nitrogen, N2 gasses. The total conversion of
the syngas into the ammonia product gas is calculated based on the mole conversion formula
as shown below:
Based on the total conversion calculated using the mole conversion formula, the PRO-II
conversion value is compared with the industrial S200 ammonia synthesis reactor conversion
value obtained from (Hignett, 1985) and the percentage difference is calculated based on the
percentage difference formula below:
Figure 3.8.1.3.2: Industrial S200 Values obtained from (Hignett, 1985)

Full Report
All the detailed calculations were provided in Appendix A8 and a general summary table is
provided below in Table 3.8.1.3.2.
Table 3.8.1.3.2: Table of comparison for the % difference between PRO-II calculations and the industrial S200 reactor
values
Parameter PRO-II Industrial S-200 % Difference
Temperature 475.80C 450.00C 5.7%
Pressure 15000 kPa 14000 kPa 7.1%
NH3 Concentration (Mole %) 21.44 % 17.10 % 25.38%
Overall Conversion (%) 30.05 % 20.21% 48.69%
Based on the percentage difference calculated between the PRO-II calculated value and the
industrial S200 ammonia synthesis reactor values, the calculated percentage difference by
HYSYS in this simulation are relatively close to the industrial S200 reactor values with all
parameters having less than 30% percentage except for the NH3 concentration value
(48.69%). This could be due to the fluid package chosen to simulate the ammonia synthesis
reactor as it will cause some deviations in mole concentration due to ideal assumptions made
for gas calculations. Hence, the value for PRO-II is justified.
3.8.2 Energy Balance for Ammonia Synthesis Reactor (R-601)

For the energy balance across the ammonia synthesis reactor, the heat of formation method is
used to calculate the amount of heat generated by the chemical reaction, Haber process.
3.8.2.1 Assumptions for Energy Balance

a) There is no work done and there are negligible changes in kinetic and potential
energies. The energy balance thus reduces to H = 0 which is calculated as follows
using the heat of formation method.
b) There is no side reaction of other gasses besides the main reactant gasses of Haber
process (BBC,2010):
c) No moving parts Ws = 0
d) Neglect effects of pressure changes on enthalpies calculated
e) The average temperature of the entering streams combined (S-120 and S-163) is
340oC
Full Report
3.8.2.2 Basis for calculations

The basis of calculations for the energy balance is based on the mass flow rate of the entering
and exiting streams of the ammonia synthesis reactor (R-601) obtained from PRO-II and is
summarized in the enthalpy Table 3.8.2.2.1 below:
Table 3.8.2.2.1: Enthalpy table for calculations of energy balance
Substance nin Hin nout Hout
(kmol/hr) (kJ/mol) (kmol/hr) (kJ/mol)
N2 436.77 H1 305.52 H4
H2 1426.99 H2 1033.28 H5
NH3 111.30 H3 373.77 H6
3.8.2.3 Key Steps

Step 1: Calculations of all enthalpies for the reactant and product gas involved in Haber
process
With reference to Table B.8 and B.2 from (Felder et.al. 2005), all enthalpies were calculated
based on the following for each component
1. For N2 and H2 gasses, the enthalpy values were obtained through interpolation of inlet
temperature and outlet temperature of the ammonia synthesis reactor using Table B.8.
2. For NH3 gasses, the enthalpy value was calculated using the formula given in Table
B.2, which Cp = a + bT + cT2 + dT3
Detailed calculations for all inlet and outlet enthalpies can refer to Appendix A8.
Step 2: Evaluate the heat of enthalpy H
With all the known enthalpies calculated for the inlet and outlet streams, the heat of enthalpy
can be calculated based on the formula:
Full Report
Step 3: Determine the energy generated from the reactor
Based on the assumptions that there is no work done moving part and there are negligible
changes in kinetic and potential energies. The integral energy balance formula will be
reduced to
where
The heat generated by the system is equivalent to the heat of enthalpy generated from the
chemical reaction, which is the Haber process.
Therefore, the heat generated by the ammonia synthesis reactor = -5776184 kJ/hr
= -1604.5 Kw
Full Report
3.9 Mass and Energy balance: Refrigeration and Separation Section
3.9.1 Flash calculations across S-701

Ammonia recovery process takes advantage of ammonia’s relatively lower boiling point as
compared to other components. Thus, the following separators’ calculations deal with
coexisting phases of vapour and liquid and the resultant mass balances according to
equilibrium T, P and phase compositions. An important application of VLE is the flash
calculations. We consider here only the P, T flash, which refers to any calculation of the
quantities and compositions of the vapor and liquid phases making up a two-phase system in
equilibrium at known T, P and overall composition.
S-140
S-128
S-701
S-141
Figure 3.9.1.1: Separator 701 with respective streams
Assumptions:
1. Steady-state system
2. Vapour phase is an ideal gas
3. Liquid phase is an ideal solution
The gaseous system is assumed to conform closely to Raoult’s law as implied by the above
assumptions. Vapour pressures for the pure species are obtained from individual’s Antoine
equations which were obtained from Perry’s Chemical Engineers’ Handbook and are
tabulated below:
Full Report
Table 3.9.1.1: Individual component’s coefficients

Component C1 C2 C3 C4 C5
water (H2O) 73.649 -7258.2 -7.3037 4.17E-06 2
Hydrogen (H2) 12.69 -94.896 1.1125 3.29E-04 2
Methane (CH4) 39.205 -1324.4 -3.4366 3.10E-05 2
Nitrogen (N2) 58.282 -1084.1 -8.3144 4.41E-02 1
Ammonia (NH3) 90.483 -4669.7 -11.607 1.72E-02 1
In which the equation’s format is such that
When temperature is fixed, the pressure varies along with compositions, xi and yi. For a given
temperature, the pressure range is bounded by the bubble and dew pressure Pbubble and Pdew.
Raoult’s Law gives
And because
Raoult’s Law can be altered to give
In which this equation finds application in bubble-point calculations where vapour-phase

composition is unknown
Similarly, Raoult’s law can also be solved for dew –point calculations by summing over the
xi species
Full Report
Firstly, vapour pressure Pisat for individual component at T = 0oC = 273.15 K is solved from
Table 3.9.1.1. Calculations are shown in Appendix A9.
Table 3.9.1.2: Vapour pressure of individual components at T=50oC

Components H2O H2 CH4 N2 NH3
sat o 10 4 5
P (50 C), kPa 0.61 2.04x10 3.57x10 3.65x10 429.31
In bubble-point calculations, since all species are in liquid phase zi = xi
In dew-point calculations, since all species are in liquid phase zi = yi
Since Pdew (757.77 kPa) < P (1.495 x 104 kPa) < Pbubble (1.2 x 1010 kPa), the system is in the
two-phase vapour-liquid region and a flash calculation can be made.
Ki is the equilibrium ratio which measures the tendency of a given chemical species to
partition itself preferentially between liquid and vapor phases. Ki for individual species is
obtained below. Calculations are shown in Appendix A9.
Table 3.9.1.3: Equilibrium ratio of individual components at T=50oC

K 2.39 24.415 0.0287
Rachord-Rice equation is then solved to obtain the vapour split and is given by the equation
Full Report
Calculations are shown in Appendix A9.This is solved using solver in excel to give
The individual liquid fraction of the components is then calculated as shown in Appendix A9.
Table 3.9.1.4: Liquid and vapour mole fraction of individual components

x
y
The resultant molar flow of each component for respective phases and deviations from PRO-
II are tabulated in Table 3.9.1.5.
Table 3.9.1.5: Vapour-liquid flow from hand calculation comparison to PRO-II for S-701
Vapour (kmol/hr) Liquid (kmol/hr)
Components Hand- % Hand- %
PRO-II PRO-II
calculations Deviation calculations Deviation
water 1.4587E-4 2.4439E-5 83.25 0.8716 0.9232 5.92
hydrogen 1033.2248 976.0550 5.53 1.8520E-4 57.2245 3.1E7
methane 27.0499 25.1540 7.00 2.7590 4.5803 66.01
nitrogen 302.5953 288.0266 4.81 3.0211 17.4903 478.94
ammonia 39.3836 30.6637 22.14 335.0095 343.1068 2.42
Deviations may be as a result of invalid assumptions. By right, Raoult’s law is invalid in this
case since gases are not ideal at a very high pressure of 1.495 x 104 kPa and liquid mixture is
not ideal since they are not all of the same chemical nature and are too different in sizes.
Another error may arise from using Antoine’s equation regardless whether respective
components are outside the temperature range.
Full Report
3.9.2 Flash calculations across S-702

Same calculations were carried out as to the ones for S-701 using PRO-II inlet values and
comparison were made to resulting vapour and liquid.
S-144
S-142
S-702
S-143
Figure 3.9.2.1: Separator 702 and respective streams
Table 3.9.2.1: Vapour Pressure and equilibrium ratio of individual components for S-702
Components Pisat (kPa) Ki yi xi
water 0.0338 1.1077x10-4 3.3274x10-7 0.0030
hydrogen 7.1549x107 2.3459x105 0.4934 2.1031x10-6
methane 1.6420x104 53.8347 0.0376 6.9914x10-4
nitrogen 1.3935x105 456.8824 0.1499 3.2815x10-4
ammonia 97.7477 0.3205 0.3192 0.9959
Using Rachord-Rice equation, it was obtained that V/F = 0.274. The resultant flow of
individual components and deviation from PRO-II values are tabulated below.
Table 3.9.2.2: vapour-liquid flow from hand calculation comparison to PRO-II for SEP-602
Vapour (kmol/hr) Liquid (kmol/hr)
Components Hand- % Hand- %
PRO-II PRO-II
calculations Deviation calculations Deviation
water 2.6011E- 0.00
3.8592E-05 32.60 0.9232 0.9232
05
hydrogen 57.2239 56.9101 0.55 0.0006 0.3144 5.23E4
methane 4.3655 4.4746 2.50 0.2149 0.1057 50.81
nitrogen 17.3895 17.3693 0.12 0.1009 0.1211 20.02
ammonia 37.0200 95.1905 157.13 306.0869 247.9163 19.00
Full Report
3.9.3 Refrigeration Loop Mass and Energy Balance

3.9.3.1 Heat Exchanger HX-702
P = 1.495x104 kPa
Figure 3.9.3.1.1: Heat Exchanger 702 and respective inlet flowrates
The energy balance across the heat exchanger yields
Assumptions:
1. Steady-state system
2. Ideal gas system
3. Enthalpies of mixing are neglected
4. Neglect effect of pressure on specific enthalpy
An enthalpy table for the heat exchanger appears as follows:

Reference: All components (1atm, 50oC)
Table 3.9.3.1.1: Enthalpy table for HX-702

Components nin (Kmol/hr) Hin(J/mol) nout (Kmol/hr) Hout (J/mol)
water (H2O) 0.9233 0 0.05907 -1683.00
Hydrogen (H2) 1033.2795 0 1030.2441 -1444.05
Vapour- Methane (CH4) 29.7343 0 29.5205 -1777.44
phase Nitrogen (N2) 305.5170 0 304.9900 -1459.00
Ammonia
373.7706 0 79.8929 -1781.03
(NH3)
water (H2O) - - 0.1942 -46552.26
Hydrogen (H2) - - 3.7054 0.00
Liquid- Methane (CH4) - - 0.2139 2810.52
phase Nitrogen (N2) - - 0.5270 41431.38
Ammonia
- - 293.8777 -23722.00
(NH3)
Full Report
The inlet enthalpies are equal zero because of the choice of reference conditions (Tref = Tinlet
=50oC). A convenient path was made in which the initial and final points of the chosen path
still correspond as that to the true path. The path is shown below:
State 1: State 1:
vapour, 50oC liquid, 0oC
alternate
State 2: State 3:
sat’d vapour, normal Tbp sat’d liq, normal Tbp
Refer to Appendix A9 for detailed calculations in obtaining Hout.
Table 3.9.3.1.2: Comparison table for HEX-702

Hand-calculation PRO-II % deviation
Q (kW) 2523.76 2302.54 8.8
Assumptions:
1. Ideal vapour-compression refrigeration cycle

2. Steady operating conditions exists
3. Kinetic and Potential energy changes are negligible
Figure 3.9.3.1.2: Refrigerant cycle P-h Diagram

Full Report
Evaporation inlet: As the refrigeration cycle was assumed to behave ideally, the ammonia
refrigerant outlet of HEX-702 should be in vapour form.
(Engineering Toolbox, 1988)
(Engineering Toolbox, 1988)
The evaporator HEX-702 duty of 2739.93 kW which was calculated earlier is the rate of heat
removal from refrigerant ammonia. Thus
Table 3.9.3.1.3: Comparison Table Ammonia refrigerant mass flow required

Hand-calculation PRO-II % deviation
NH3 mass flow (kg/s) 1.73 1.56 9.83%
3.9.3.2 Overall Mass and Energy Balance of Refrigerant Loop

134
TV-701
HX-702
133
131 132
HX-703
K-701
Figure 3.9.3.2.1: Refrigerant Loop

131 132 133 134
Vapour Phase 1 1 1 1
Pressure (kPa) 190 800 790 200
Temperature (oC) 27.53 177.35 -25.10 -24.98
Mass flow (kg/s) 1.56 1.56 1.56 1.56
Full Report
3.9.3.3 Synthesis Loop Material Balance

Material balance for the synthesis loop is done through the sequential modular method. In the
sequential modular approach, each unit operation block is solved one at a time in sequence.
The output calculated from each block becomes the feed to the next block, and so on. In this
case, only the material balance will be solved in order to keep calculations simple. The block
structure of the ammonia synthesis loop flowsheet is given by the figure below
Convergence Purge Gas
Reactor Separator
R-601 S-701
Separator
S-702
Product
Figure 3.9.3.3.1: Block structure of sequential modular calculation
The feed stream is assumed to be known. This then goes to a mixing point where the fresh
feed is mixed with the recycle stream. Since the recycle flow rate and composition is
unknown, the sequential modular solution technique is to tear one of the streams in the
recycle loop. A recycle convergence unit is then inserted in the tear stream. To start the
calculations of the material balance in Figure 3.9.3.3.1, values for the component molar flow
rates for the recycle stream (tear stream) must be estimated. PRO-II recycle values are used
as the first estimate, in order to give a close initial guess. This allows the material balance in
the reactor and separator to be solved. In turn, this allows the molar flow rates for the recycle
stream to be calculated. The calculated and estimated values can then be compared to test
whether errors are within a specified tolerance. If the convergence criteria are not met, then
the convergence block needs to update the value of the recycle stream through repeated
substitution. The calculated value then becomes the value for the next iteration. This is
repeated until convergence is achieved.
Full Report
3.10 Energy Balance for Common Equipment
3.10.1 Energy Balance across Heat Exchanger

The duty of the heat exchanger must be determined in this section as it will help in
determining the size and the cost of the heat exchanger. Some of the reactions require high
temperature whereas some of the equipment in this plant cannot tolerate operation at high
temperature. Hence, heat exchanger is used to raise and reduce the temperature of the process
gas by using cooling water or process gas itself.
The duties for compressor HX-102, HX-102, HX-103, HX-201, HX-202, HX-203, HX-204,
HX-205, HX-206, HX-207 HX-301, HX-302, HX-401, HX-402, HX-403, HX-404, HX-501,
HX-502, HX-601, HX-602, HX-701, HX-702 HX-703, HX-704, HX-705 and HX-706 are
determined based on the same general steps shown in this section of the report.
3.10.1.1 Basis for Energy Balance Calculation

The temperature requirements for each stream are different. Duty of each heat exchanger is
different based on its temperature difference and its composition. Cooling water is supplied at
25 and leaves the heat exchanger at 50 unless stated otherwise. For service fluids other
than cooling water, the inlet and outlet temperatures are specified before performing the
calculations.
3.10.1.2 Assumptions for Energy Balance Calculation

 Steady state operation.
 Adiabatic heat exchanger. All heat loss by process gas = heat gained by cooling water
(or other service fluids).
 Pressure drop has negligible effect on the change in enthalpy.
 The Cp of cooling water is not affected by temperature change.

Full Report
3.10.1.3 Calculation Steps for Energy Balance

Step 1: Establish the equation based on change in enthalpy
Using HX-103 as an example, the following equation is used.
Inlet temperature =
Outlet temperature =
Molar Molar specific heat capacity Constants

Syngas
Flow, + + + ( . )
Component
332.9596 32.243 1.92E-03 1.06E-05 -3.60E-09 -4802.111 -1598909.017

184.7926 27.143 9.27E-03 -1.38E-05 7.65E-09 -4050.357 -748475.951
124.8599 30.869 -1.29E-02 2.79E-05 -1.27E-08 -4089.925 -510667.683
99.8879 19.795 7.34E-02 -5.60E-05 1.72E-08 -5772.313 -576584.242
49.9439 19.251 5.21E-02 1.20E-05 -1.13E-08 -5799.098 -289629.560
7.4916 3.806 1.57E-01 -8.35E-05 1.76E-08 -7625.044 -57123.779
6.4927 5.409 1.78E-01 -6.94E-05 8.71E-09 -9310.463 -60450.041
1.4983 -4.224 3.06E-01 -1.59E-04 3.21E-08 -13379.107 -20045.916
7.4916 31.15 -1.36E-02 2.68E-05 -1.17E-08 -4072.703 -30511.060
0.0228 -35.843 5.98E-01 -4.83E-04 1.53E-07 -19140.667 -436.407
0.0328 -28.248 6.16E-01 -4.02E-04 9.94E-08 -22551.387 -739.685
0.6253 -42.944 6.90E-01 -4.34E-04 9.15E-08 -23868.569 -14925.016
0.2947 -68.802 8.50E-01 -6.51E-04 1.98E-07 -25341.524 -7468.147
0.0398 -64.623 8.85E-01 -5.85E-04 1.31E-07 -28783.764 -1145.594
0.0114 -54.491 9.04E-01 -5.39E-04 1.31E-07 -32320.352 -368.452
0.0149 -58.979 1.01E+00 -6.59E-04 1.61E-07 -34965.132 -520.980
-3918001.53
Full Report
Step 2: Determine duty of the process gas
By substituting all the values required and integrating, the duty can be found.
The total heat lost from the syngas, S120, or in other words, the duty of the waste heat boiler
is calculated as follow:
For HX-103,
The negative sign indicates the loss of heat to the steam.
Step 3: Determine the amount of cooling water needed to achieve duty

Cooling water is fed to the heat exchanger at 1 bar and 25 and leaves at 50 . No phase
change occurs the enthalpy change involves only the sensible heat.
The molar and mass flow rate of the water required is calculated as shown below:
OR for Process fluid-Process fluid Heat Exchanger, Step 3 will be done by:
Step 3: Determine the outlet temperature of the heat exchanger

All heat from tube side will be absorbed by shell side
Using goal seek in Microsoft Excel by setting heat duty required, the outlet temperature of
the process fluid will be calculated.
Full Report
3.10.1.4 Summary Duty of Heat Exchanger

Heat Heat
Service Fluid Duty (kJ/h) Service Fluid Duty (kJ/h)
Exchanger Exchanger
HX-101 Air 8514000 HX-403 Cooling Water 982260.00
HX-102 Cooling Water 4211000 HX-404 Cooling Water 8020800.00
HX-103 Cooling Water 3916800 HX-405 Steam 426240.00
HX-201 Cooling Water 1632461.58 HX-501 Cooling Water 1409418.39
HX-402 Cooling Water 6348000.00
*Detailed calculations are shown in Appendix A10.1
3.10.2 Energy Balance across Compressor

The duty of the compressor must be determined in this section as it will help in determining
the size and the cost of the compressor. Some of the reactions in this plant require high
pressure to achieve high conversion. Compressor is used to raise the pressure of the process
gas.
The duties for compressor K-101, K-102, K-201, K-202, K-203, K-204, K-205, K-206, K-
207, K-208, K-601, K-602, K-701, K-702, K-703, K-704, K-705, K-706 and K-707 are
determined based on the same general steps shown in this section of the report.
3.10.2.1 Basis for Energy Balance Calculation

The design fluid properties such as the inlet volumetric flow rate and specific heat capacity
are taken from the Pro II simulation was done. Since the energy balance calculations for
centrifugal compressors are based on the mass flow rate of inlet, the mass flow rate values of
fluid entering the pumps are extracted from Pro II.
Full Report

 The entire system is steady state with no accumulation
 No mass generation within the compressor
 The loss of mass during the process is negligible
There the loss of mass during the process is negligible. This process only involved energy
balance.
3.10.2.3 Calculation Steps for Energy Balance

Step 1: Determine the type of compressors used
There are 2 types of compressor that are generally used in the industry: centrifugal and
reciprocating. Hence, it is required to determine the type of compressor used before
performing calculation because the calculations involved are different. It is determined based
on the volumetric flow rate of the process gas and developed pressure from Figure 10.60 of
“Chemical Engineering Design” book by Sinnot and Towler 2009.
Step 2: Performing Calculation (Centrifugal Compressor)
The Schultz’s method is applied by using this equation:
where
The calculation is carried out as follows
From volumetric flowrate and mass flow was determined from Pro II
m and n values are determined based on the equations below:
where
By extrapolating the polytropic efficiency curve from Sinnott and Towler (2009), the
polytropic efficiency can be determined.
Full Report
After the constant m and n are determined, work done, by the compressor can be
calculated by using the equation mentioned.
Actual mechanical compression work
The electric motor efficiency is estimated from Sinnot and Towler (2009). Hence, the electric
power required for the compressor can be determined.
Step 3: Calculate the temperature rise for polytropic compression
The temperature rise will be determined by using the following equation,
HOWEVER, FOR RECIPROCATING COMPRESSOR, CALCULATION STEPS AS

BELOW:
Step 2: Performing Calculation (Reciprocating Compressor)
where
The calculation is carried out as follows
From volumetric flowrate and mass flow was determined from Pro II
Based on the design pressure, isentropic efficiency is determined from the isentropic
efficiency curve from Sinnott and Towler (2009)
After the constants are defined, work done, by the compressor can be calculated by
using the equation mentioned.
Actual mechanical compression work
The electric motor efficiency is estimated from Sinnot and Towler (2009). Hence, the electric
power required for the compressor can be determined.
Full Report
Step 3: Calculate the temperature rise for polytropic compression
The isentropic temperature will be determined by using the following equation,
The actual temperature will be calculated using the following equation:
3.10.2.4 Summary results of Compressors

Outlet Outlet
Compressor Duty kW Temperatur Compressor Duty kW Temperatur
e, °C e, °C
K-101 1166 150 K-601 1175.99 234.21
K-102 1166 150 K-602 298.72 96.95
K-201 494.59 150 K-701 510.09 177.40
K-202 450.35 150 K-702 22.73 50
K-203 449.47 150 K-703 37.39 150
K-204 224.41 161.60 K-704 30.02 135.45
K-205 165.95 150.92 K-705 29.67 160.06
K-206 166.15 150.94 K-706 27.26 182.18
K-207 162.10 150.28 K-707 31.02 22
K-208 166.58 150.97
*Detailed calculations are shown in Appendix A10.2
3.10.3 Energy Balance across Centrifugal Pump

The duty of the pump is determined in this section as it will help in the selection of a suitable
pump in which the pump characteristics matches that of the system curve.
The two types of pumps used in the process of Alternis BioAmmonia Chemical Ammonia
Production Plant are single-stage centrifugal pumps and multi-stage centrifugal pumps. The
single-stage, horizontal, overhung, centrifugal pump is commonly used in chemical process
industry for system with moderate flow rate and pressure heads. A multi-stage centrifugal
pump on the other hand is used to pump high flow rate fluids and overcome the huge static
pressure caused by the difference in height and pressure, the dynamic loss due to friction in
the pipe, the miscellaneous losses and the pressure loss through equipment.
The duties for pumps P-101, P-102, P-301, P-401 and P-402 are calculated based on the
general steps shown in this section of the report.
Full Report
3.10.3.1 Basis for Energy Balance Calculations

The design fluid properties such as the inlet density of fluid, and inlet viscosity of fluid
are taken from the Pro II simulation done earlier on this report in Chapter 2. Since the
energy balance calculations for centrifugal pumps are based on the mass flow rate of inlet, the
mass flow rate values of fluid entering the pumps are extracted from Pro II. This was done so
that differences between the duties calculated by Pro II and by hand can be easily compared.

 Fluid such as cooling water is supplied at 25 and 1atm while boiler feed water is
supplied at 30
 The temperature of fluid entering the pump is the same as the temperature of leaving
the pump as the temperature rise across the pump is negligible due to the large heat
capacity of water
 Commercial steel pipe with ASME standards are used to determine diameter of the
pipes and ultimately the velocities at the suction and discharge of the pump.
 The diameter of pipe attached to the suction side of the pump is equivalent to the
diameter of pipe attached to the discharge end of the pump
 The length of pipe at the suction side is 30m while the length of pipe at the after the
discharge of the pump is 10m. Hence the total pipe length, is 40m.
 The difference in elevation between the boiler feed water level in the tank and the
inlet to the heat exchanger is 1m
It is very important for the waste heat boiler to receive high pressure water for the production
of steam and the water from the storage tank need to be pumped to the waste heat boiler. So,
water at and at atmospheric pressure from the water tank need to be pumped to the
waste heat boiler at pressure of . The required flow-rate of water for the waste heat
boiler was calculated to be . Table A5.1.2.1 below shows the design fluid
properties used and the informations on the lines and fittings at suction and discharge of the
pump is shown in Table A5.1.2.2
Full Report
Table 3.10.3.2.1: Fluid Properties

Storage Tank Waste Heat Boiler
Total Mass Flow Rate,
(
Velocity in Equipment,
Pressure,
Temperature,
Density, (
Viscosity,
Volumetric Flow-rate,
Table 3.10.3.2.: Piping and Fitting information

Suction (1) Discharge (2)
Total Mass Flow Rate, (
Line Inner Diameter,
Elevation,
Pipe Length,
Pipe Fittings (unit)
 elbow
 Gate Valve (fully open)
 Globe Control Valve
 Check Valve
 Y-strainer
 Reducer 1 1
 Expander 0 2
 Orifice 0 1
 Entrance to Pipe from Tank – Flush Sharp Edged 1 0
 Entrance to WHB – Protecting Sharp Edged
0 1
Rounded
There are two components of the pressure head that has to be supplied by the pump in a
piping system, Static Head and Dynamic Head. Both the component head was calculated. The
total Dynamic head was obtained to be 0.01095m and Static head was obtained as 530.71 m
and hence, obtaining the total head to be 530.722 m. Based on the mechanical design in
Appendix A5.1.2, the efficiency of the pump is obtained as 69.9%. To transport the boiler
feed water from the water storage tank to the waste heat boiler, the amount of energy required
per kg of fluid, is shown in equation below:
Full Report
Therefore, the amount of Power input to the pump is:
Please refr to Appendix A10.3 for sample calculation
3.10.3.4: Summary of Results of Pumps

Pump Duty (kW)
P-101 11
P-301 8.95
P-401 34.98
P-402 35.05
3.10.4 Mass and Energy Balance across Fired Heater

Assumptions
 The fuel gas is methane

 20% excess air is used in the fired heater (Wildy, 2000)
 The ideal air to fuel ratio which is 20:1 is used in the fired heater(Whittall, 2013)
 Complete combustion of fuel in the fired heater
 The fuel supply just sufficient heat to heat up the syngas, no extra heat is generated
 First law of Thermodynamic is applied. All the heat generated is used to heat up the
syngas to desired temperature without any heat loss to the surrounding
 The industrial outlet temperature of flue gas which is 800°C
 The changes in kinetic and potential energies are negligible
 The system is steady state where there is no mass accumulation there is no generation
in the system (Syngas passing through the fired heater)
 The syngas is assumed to be ideal gas and its difference in pressure is negligible
 The references for each composition that form the reactants and product were chosen
to be at 25 and 1 atm
 The effect of pressure changes on the enthalpies are neglected
Full Report
 The change in kinetic, potential and shaft work in negligible. Thus, the open system
energy balance gives .
3.10.4.1 Mass and Energy Balance of Syngas Passing through the Fired Heater
In order to ensure that the fuel supply to fired heater is sufficient to generate enough heat in
heating up the syngas to desired temperature, energy balance for heating up the syngas from
to 850 is calculated. In this case, syngas will be heating up. For fired heater, syngas
only involved in the energy balance.
Detail calculations for the energy balance are shown in Appendix A10.4
Table 3.10.4.1.1 Summary of Mass and Energy Balance of Syngas across Fired Heater
Mass Balance
Composition Inlet (kg/h) Outlet (kg/h)
Water Vapour 59.9367 3.327
Carbon Monoxide 3496.0681 124.811
Carbon Dioxide 4395.0604 99.8657
Hydrogen 369.587 183.3269
Nitrogen 209.7688 7.4883
Methane 799.101 49.8103
Oxygen 0 0
Ethane 194.7846 6.4777
Ethylene 209.7688 7.4774
Propane 65.9245 1.495
Total 9800 9800
Energy Balance
Heat Required, kJ/h 12939190
3.10.4.2 Mass Balance of the Combustion of Fuel in the Fired Heater

The mass balance in this section is started by assuming an amount of fuel that will involve in
the combustion reaction. The combustion reaction is assumed to undergo complete
combustion where the fuel supply is sufficient to raise the temperature to a higher
temperature. Iterations will be done by using Microsoft Excel to determine the exact amount
of fuel required for the combustion. However, the amount of air required for combustion is
calculated by using the ideal air fuel ratio from literature studies.
Therefore, both mass and energy balances need to be prepared before the iterations are done.
Brief steps descriptions are as follow:
Step 1: Assume an amount of fuel and calculate the air required based on the ideal air fuel
ratio. Complete the mass balance using the current values.
Full Report
Step 2: Using the flow calculated in step 1 to complete the energy balance. The detailed
calculation of energy balance is descripted in the next section.
Step 3: Iterations to obtain the exact amount of fuel are done by equating the amount of
energy requires to raise the temperature of syngas from 150 to 850 to the energy
released by the exhaust gas from the combustion of fuel. This will be done by Goal Seek in
Microsoft Excel.
Step 4: After the exact amount of fuel is calculated, the require amount of air will be
calculated using the same approach with additional 20% of excess air.
Detail calculations for the mass balance are shown in Appendix A10.4
The mass balance for the combustion of fuel is shown in Table 3.10.4.2.1
Table 3.10.4.2.1 Mass Balance for the Combustion of Fuel

IN Dis Gen OUT
Composition Molar Molar Molar Molar
Mass Flow, Mass Flow,
Flow, Flow, Flow, Flow,
kg/h kg/h
kmol/h kmol/h kmol/h kmol/h
Fuel (CH4) 412.55 25.78 -25.78 0 0 0
Oxygen 2307.10 72.10 -51.57 0 656.88 20.53
Nitrogen 7594.19 271.22 0 0 7594.19 271.22
Carbon
0 0 0 25.78 1134.52 25.78
Dioxide
Water 0 0 0 51.57 928.25 51.57
Total 10313.84 369.10 10313.84 369.10
3.10.4.3 Energy Balance of the Combustion of Fuel in the Fired Heater

Assuming that the inlet temperature of air and fuel are at 25 and the outlet temperature of
the flue gas is assumed to be 800 . The fuel will be undergone complete combustion in the
presence of excess air to produce carbon dioxide and water. The fuel will be combusted in the
presence of excess air up to 1000 and flue gas leaving the fired heater at 800 .
The energy balance for combustion of fuel in fired heater will be done as follow:
Full Report
process state. Thus, sensible heat enthalpy will be calculated. The constants of integration of
heat capacities are shown in Table 4.4.3.3.1.
Step 2: After the sensible heat enthalpy for each compositions are calculated and thus the
overall . The overall will be adding up to the heat of combustion of fuel (Methane).
Step 3: By using goal seek, 412.55 kg/h of fuel is required to supply to the fired heater in
order to heat up the syngas to the desired temperature. Taking the air to fuel ratio of 20:1 and
with 20% of excess air, the amount of air will be supply to the fired heater is 9901.29 kg/h.
Detailed calculations are shown in Appendix A10.4
The constants of integration of heat capacities for each composition involved in the
combustion reaction are shown in Table 3.10.4.3.1.
Table 3.10.4.3.1 Constant of Integration of Heat Capacities for each composition

Composition a b c d
Fuel (CH4) 0.03431 0.00005469 3.661E-09 -1.1E-11
Oxygen 0.0291 0.00001158 -6.076E-09 1.311E-12
Nitrogen 0.029 0.000002199 5.723E-09 -2.871E-12
Carbon
0.03611 0.00004233 -2.887E-08 7.464E-12
Dioxide
Water 0.03346 0.00000688 7.604E-09 -3.593E-12
The summary of energy balance of the combustion of fuel in the fired heater is shown in
Table 3.10.4.3.2
Table 3.10.4.3.2 Heat released from the combustion of fuel

Heat Released from the Combustion
Total sensible heat capacities (from 25°C to 800°C) 9058468.801 kJ/h
Heat of Combustion -21997658.6 kJ/h
Total heat released -12939190 kJ/h
Full Report
CHAPTER 4| DEMONSTRATION OF SUSTAINABILITY CONCEPT

4.1 Environmental Evaluation: LCA Methodology
4.1.1 Goal Definition
4.1.1.1 Goal and Scope

The goal of this life cycle assessment study is to evaluate the life cycle environmental
performance of anhydrous ammonia production from biomass whereby a cradle-to-gate
analysis is performed based on the partial life cycle of the product from the acquisition of
biomass (cradle) to the anhydrous ammonia factory gate before the product is transported to
the consumers.
With this analysis taking an in depth look at each of the unit operations from the raw material
acquisition stage, followed by the material manufacturing stage to the product manufacturing
stage, all flows across each unit operation are identified and the overall effect of total flows
across the system boundary across each stage are evaluated. This is done so that a more in
depth insight into the areas that contribute significantly to emissions are identified as these
areas will deliver the greatest results upon improvement.
Furthermore, the purpose of this life cycle analysis is to assess and compare holistically the
environmental impacts from the manufacture of one kilogram of anhydrous ammonia using
renewable oil palm trunk (OPT) feedstock and non-renewable natural gas feedstock.
Essentially, estimations on the environmental impact of product are necessary in order to
identify the environment benefits of utilizing oil palm trunk (OPT) as feedstock. The LCA
conducted is addressed to both manufacturers, Alternis BioAmmonia Pvt. Ltd. and the
consumer as the stakeholders of Alternis BioAmmonia Pvt. Ltd. not only aim to guarantee
ammonia of high purity for further usage as fertilizer but also come up with a good design in
order to practice a low carbon footprint for good sustainability.
4.1.1.2 Functional Unit

The functional unit chosen for this study is 1 kg of anhydrous ammonia produced for both the
resources. Therefore, the inventory analysis has been performed by determining the raw
material and energy consumed as well as the waste produced from the all the stages of the
production phase per kilogram of ammonia produced. As for the raw material acquisition
stage, the data available from multiple literature resource were per kilo gram of the raw
material. Thus, these were then converted to per kg of ammonia produced by determining the
required raw material per kg of ammonia.
4.1.1.3 System Boundaries

The system for this analysis is defined to be Cradle to Gate in which includes only the
production phase in which is divided into 3 stages:
Full Report
Materials Energy Materials Energy Materials Energy
Natural
Resources Raw Material Material Product
Acquisition Manufacture Manufacture Products
By-Products Emissions By-Products Emissions By-Products Emissions
Figure 4.1.1.3.1: Stages in Production Phase of LCA (Brennan, 2012)
Figure 4.1.1.3.2 and Figure 4.1.1.3.3 shows the production phase life cycle of the products
consisting of the 3 stages above and the LCA boundary is shown as well as the input streams
and the output streams crossing the boundaries for plant with biomass feedstock and natural
gas feedstock respectively. The table 4.1.1.3.1 below shows the inclusion and exclusion of
the system boundaries that were taken into account during the Life Cycle Assessment.
Table 4.1.1.3.1: Inclusion and Exclusion across the LCA boundary
Inclusion Exclusion
Raw material input for the raw material Raw material requirement and Emission from
acquisition stage (Palm Plantation/Natural construction of plant
Gas Mining)
Emission from the raw material acquisition Installation process of the equipment in the plant
stage (Palm Plantation/Natural Gas Mining) and Plant Set-up
Fuel Requirement and emissions for Raw Material requirement and emissions during
transportation of the raw material to the plant start-up, plant shut-down, abnormal
ammonia plant (OPT/Natural Gas) operation and maintenance
Raw material and Electrical Energy Manufacture of Catalysts, refrigerants and other
requirement for Material Manufacture and absorbents (MDEAmine, Piperazine)
Product Manufacture
Emissions from Material Manufacture and Fugitive and abnormal Emissions and spills
Product Manufacture
Raw material requirement and Emission from Losses from storage and transport
electricity consumption during material and
product manufacture
Transportation of Product to Fertilizer
manufacturing
Manufacture of Fertilizer using ammonia
produced
Full Report
Process Flow Diagram of Production of Anhydrous Ammonia using

OPT Biomass Feedstock
Raw Material (Input)
Atmospheric Air-borne Emission

Oil Palm Plantation
CO2
Water-borne Emission
Harvesting OPT
Biomass
Transport of OPT to
Fuel Air-borne Emission
Plant
Shredding
Air-borne Emission
Raw Material
(Input) Electricity Energy
Drying Air-borne Emission
Generation
Gasification Air-borne Emission
Post Treatment of
Syngas
Autothermal
Air-borne Emission
Raw Material Reformer
(Input)
Carbon Doxide
High and Low Storage
Temperature Shift
Reactors
Glycol Plant
Carbon Doxide
Removal
Methanator Water-borne Emission
Ammonia Synthesis
Ammonia
Air-borne Emission
Purification
Storage of Ammonia
Transport to Fertilizer
Manufacturing Plant
Agricultural Fertilizer
Figure 4.1.1.3.2: Process Flow Diagram and LCA Boundary of Production of Anhydrous Ammonia using OPT Biomass as
feedstock
Full Report
Process Flow Diagram of Production of Anhydrous Ammonia using

Natural Gas Feedstock
Raw Material (Input)
Air-borne Emission
Natural Gas Extraction
Natural Gas Processing
Transport of Natural
Fuel Air-borne Emission
Gas to Plant
Primary Reformer
Air-borne Emission
Raw Material High and Low

(Input) Electricity Energy Temperature Shift
Generation Reactors
Carbon Dioxide
Removal
Raw Material Methanator

Air-borne Emission
(Input)
Compressors
Ammonia Synthesis
Ammonia Purification
Storage of Ammonia
Transport to Fertilizer
Manufacturing Plant
Agricultural Fertilizer
Figure 4.1.1.3.3: Process Flow Diagram and LCA Boundary of Production of Anhydrous Ammonia using Natural Gas as
feedstock (NETL, 2010)
Full Report
4.1.1.4 Time Horizon

The time considerations for this LCA report are based on the pollutant life span and the
availability of natural resources over a time interval of 100 years. A time horizon of 100
years is commonly used by regulators as the concentration of carbon dioxide takes the longest
time to decay in the atmosphere. On top of that, the operation life span of Alternis
BioAmmonia Ammonia Production plant is 25 years (Dietze, 2013).
4.1.1.5 Assumptions and Limitations

During the Life Cycle Assessment of the production phase of the anhydrous ammonia plant
from both the resource, few assumptions were made in order to simplify the assessment and
due to the limitation in the data available. Below are the assumptions and limitations that
were considered throughout the study:
 It was assumed that the impacts from the construction of the plant, manufacturing of
the catalysts, absorbents, and refrigerants are relatively small compared to the impacts
from the steady state operation of the plant. It also poses limitation to the study as
their assessment will add disproportionately to the time of the study. However, the
study can be revised in future in order to include the omitted impacts.
 The study also has limitations where it excludes the impacts during abnormal
operations, start-up and shut-down, cleaning and maintenance activities as there is no
available data on the emissions during these periods.
 Fugitive and abnormal emissions as well as spillage were assumed to be likely minor
to negligible due to lack of data and besides, in case of fugitive emission to take place
in the plant, assumed that immediate actions will be taken.
 Electricity requirement at the material acquisition stage is assumed to be negligible
 It is assumed that the generation of electricity in Perak, Malaysia is from Natural Gas
and the emissions consists of CO2, SO2, NOX, and CO (Mahlia, 2001))
 Assumptions regarding means of transport for the transport of palm biomass, OPT
largely refers to truck transport whereby the maximum load of 1 truck is 40 ton. It is
assumed that Diesel Oil is used by the truck as fuel and assumed that 30 Liters of
Diesel oil is required to travel 100 km (Smith, 2008).
 For the oil palm plantation, it is assumed that the oil palm absorbed an average of 29.3
tonnes of CO2 per hectare (Packaging, 2010)
 The transportation of Natural Gas for both the plants is through pipeline and emission
that may occur during the transportation of Natural Gas is assumed to be negligible.
 For the plant, the emissions that may occur over time due to corrosion of equipment,
erosion of equipment material by flowing of water or other chemicals or gases are
assumed to be negligible.
4.1.2 Inventory Analysis

This Life Cycle Inventory Analysis is considered to be the most developed stage of the
assessment as it record and quantifies every data which is entering the system boundary
including the raw materials and the energy requirements and is coming out to the
environment across the boundary such as the desired product, air emissions, waterborne
Full Report
emission, and other solid wastes .These data were obtained from various multiple resources
where some values were estimated from similar operations, published data and commercially
available databases. However, these values were expressed in terms of the functional unit
which is per kilogram of anhydrous ammonia, NH3 for both sources.
4.1.2.1 Extraction of Raw Material and Requirement

This section identifies and quantifies the input raw materials required per kilogram NH3
produced by the Alternis BioAmmonia using OPT and the conventional production using
Natural gas and are tabulated in Table B1.1.1 and Table B1.1.2 respectively. The raw
material inputs includes the raw material required for the raw material acquisition stage,
transportation, material and product manufacture and electrical energy production.
4.1.2.2 Electrical Energy Requirement

The average energy requirements for each process identified in the inventory analysis are
quantified in terms of kilowatt-hour (kWh) of electricity per kilogram of NH3. As mentioned
in the assumption, the electricity required for the raw material acquisition is assumed to be
negligible and thus was not considered. For the production of Anhydrous ammonia using
OPT biomass as the feedstock, 37 kWh of electrical energy is found to be required by the
plant. Meanwhile, a total electrical energy of 1810 kWh is required for the production of
anhydrous ammonia using the Natural Gas as the feedstock (Jinenz-Gonzales et al., 2008).
4.1.2.3 Environmental Emission

Environmental emissions are categorized as atmospheric emissions, waterborne emissions,
and solid wastes and represent discharges into the environment after the effluents pass
through existing emission control devices.
4.1.2.3.1 Atmospheric Emission

These emissions include substances classified by regulatory agencies as pollutants, as well as
selected non-regulated emissions such as carbon dioxide. For each process, atmospheric
emissions associated with the combustion of fuel for process or transportation energy,
electrical energy generation for the consumption of the plant as well as any emissions
released from the process itself, are included in this LCI. Based on the analysis, the most
commonly reported atmospheric emissions are: carbon dioxide, carbon monoxide, and
emissions from the pesticides. However, like all plants, oil palm absorbs CO2 from the
atmosphere and emits O2 to the atmosphere through photosynthesis. Literature sources
suggest that oil palm trees absorbs an average of 29.3 tonnes of CO2 per hectare of the
plantation. The CO2 that was removed at the Carbon Dioxide Removal stage was also taken
into consideration. The atmospheric emission per kilogram of NH3 produced from both
resources is tabulated in Table B1.1.3 and B1.1.4.
4.1.2.3.2 Waterborne Emission

The values reported are the average quantity of pollutants still present in the wastewater
stream after wastewater treatment and represent discharges into receiving waters. The water-
borne emission includes only the process related emissions as the waterborne emissions from
the combustion of fuel for transportation and electrical energy is assumed to be negligible.
Full Report
The waterborne emission per kilogram of NH3 produced from both resources is tabulated in
Table B1.1.5 and B1.1.6.
4.1.3 Impact Assessment

The resources and emissions both air and water borne emissions were characterized,
normalized and weighted in order to compare the contribution to the impact categories
selected. This was done by using the ReCiPe108 LCIA Methodology which was designed by
the PRE Consultants, Radboud University Nijmegan, Leiden University, and RIVM
(Goedkoop et al., 2008). This methodology comprises harmonized category indicators at the
midpoint and the endpoint level, and hence, transforms the long list of Life Cycle Inventory
results, into a limited number of indicator scores.
4.1.3.1 Classification
In the classification, the environmental interventions listed in the inventory tables both the
inputs and outputs across the system boundary are attributed to the selected impact categories
under ReCiPe 2008 Method contains the characterization factors. The impact categories
which are contributed by the process prominently are Global Warming Potential (GWP),
Terrestrial Acidification Potential (AP), Photochemical Acidification Potential (POFP),
Particulate Matter Formation Potential (PMFP), Marine Eutrophication Potential (MEP),
Ozone Depletion Potential (ODP), Freshwater Eutrophication Potential (FEP) and Resource
Depletion which have been divided into Water Depletion (WDP) and Fossil Depletion (FDP).
The classification method enables the emissions to be placed into its suitable category with
some species overlapping.
4.1.3.2 Characterization
The species or burdens have then been weighted by multiplying by the characterization factor
as provided in the ReCiPe 2008 Method Spreadsheet to give a value in terms of one species.
For GWP all species are referenced to CO2, AP is referenced to Sulphur Oxides, POFP is
referenced to Non-Methane VOC, PMFP is referenced to Particulates of less than 10μm,
MEP is referenced to N compounds, ODP is referenced to Trichluorofluoromethane (CFC-11)
and FEP is referenced to Phosphate compounds(Goedkoop et al., 2008). For the breakdown
of Resource Depletion, WDP is referenced to water and FDP is referenced to Crude Oil
(Please Refer to Appendix B1.2 for Detailed Calculation). The weighted burdens within
each category are then added to give scores for the different categories. The Table B1.2.1 to
B1.2.7 in Appendix B1.2 shows the characterization of the species that are considered as
burden to the key impact categories These impact scores of each category for both production
phase of ammonia with respective to their feedstock were then displayed in a figure below to
be compared between the impact categories.
Full Report
Impact Score of Production of NH3 using OPT Biomass

5
Feedstock
Impact Score (kg C eq/kg NH3)
0
WDP FDP GWP AP POFP PMFP MEP ODP FEP
-5
-10
-15
-20
-25
-30
-35
Impact Category
Figure 4.1.3.2.1: Impact Score of Production of NH3 using OPT Biomass as Feedstock
Impact Score of Production of NH3 using Natural Gas

Feedstock
1600
Impact Score (kg C eq/kg NH3)
1400
1200
1000
800
600
400
200
0
Impact Category
Figure 4.1.3.2.2: Impact Score of Production of NH3 using Natural Gas as Feedstock
4.1.3.3Normalization
Normalization relates all of the characterized results to a common unit of activity. This stage
eliminates the different units between the different impact categories so that they are all
expressed as a proportion of an average impact of the life cycle of a product. These
normalized results reveal the extent to which different alternatives of the product of the same
application contribute to the different impacts. The impact scores from the classification
section above have been normalized against World consumption or also known as Global
Value. The table below shows the Global Normalization data based on the ReCiPe108 LCIA
Methodology Spreadsheet.
Full Report
Table 4.1.3.3.1: Global Normalization Data from ReCiPe108 LCIA Methodology Spreadsheet(Goedkoop et al., 2008)
Impact Categories Unit ReCiPe Midpoint (H)
Hierarchist World
Water Depletion Potential (WDP) m3/yr 2.57E+12
Fossil Depletion Potential (FDP) kg Crude Oil eq/yr 7.90E+12
Global Warming Potential (GWP) kg CO2 eq/yr 4.22E+13
Terrestrial Acidification Potential (AP) kg SO2 eq/yr 2.34E+11
Photochemical Oxidant Formation Potential kg NMVOC/yr 3.00E+11
(POFP)
Particulate Matter Formation Potential (PMFP) kg PM10 eq/yr 8.61E+10
Marine Eutrophication Potential (MEP) kg N eq/yr 4.49E+10
Ozone Depletion Potential (ODP) kg CFC-11 eq/yr 2.30E+08
Freshwater Eutrophication Potential (FEP) kg P eq/yr 1.77E+09
By dividing the impact scores obtain above with the normalization data respectively for each
impact categories gives the normalized score for the categories which are tabulated in Table
4.1.3.3.2 and comparison between the 2 process is done in the bar chart in Figure 4.1.3.3.1.
Table 4.1.3.3.2: Normalized Impact Score for the impact Categories of both sources
Impact Categories Normalized Impact Score against Global
Value (yr/kg NH3)
OPT Biomass Natural Gas
Feedstock Feedstock
Water Depletion Potential (WDP) 1.92E-13 4.71E-13
Fossil Depletion Potential (FDP) 7.63e-14 1.88E-10
Global Warming Potential (GWP) -7.12E-13 2.27E-11
Terrestrial Acidification Potential (AP) 7.27E-11 7.86E-12
Photochemical Oxidant Formation Potential (POFP) 1.14E-11 5.85E-12
Particulate Matter Formation Potential (PMFP) 9.41E-12 6.31E-12
Marine Eutrophication Potential (MEP) 8.73E-12 4.60E-13
Ozone Depletion Potential (ODP) 1.05E-16 4.86E-20
Freshwater Eutrophication Potential (FEP) 5.84E-15 9.35E-17
Full Report
Normalised Score for the Impact Categories of Biomass and

Natural Gas feedstock NH3 Production
2.00E-10
OPT
Biomass
Impact Score (yr/kg NH3)
1.50E-10
Natural Gas
1.00E-10
5.00E-11
0.00E+00
-5.00E-11
Impact Category
Figure 4.1.3.3.1: Normalized Impact Score for the Impact Categories of Biomass and Natural Gas feedstock NH 3
Production
4.1.4 Interpretation
The major environmental impact categories that are largely contributed by the production of
Ammonia based on both feedstock are Fossil Depletion, Global Warming Potential,
Terrestrial Acidification, and slightly contributes to the Photochemical Oxidant Formation,
Particulate Matter Formation Potential and Marine Eutrophication. It can be clearly seen that,
production of Ammonia using the Biomass Feedstock largely reduced the impact on the
Fossil Depletion as the Conventional production requires mining of the natural gas. Usage of
the palm biomass waste does not require fossil fuels except for the purpose electricity and
transport fuel requirement. This goes along with the current global issue of mitigating the
natural resource depletion as substitution of biomass as feedstock for ammonia production
will help to reserve the natural resources better. Besides that, the Global Warming Potential
due to the carbon dioxide emission is also largely reduced by the usage of Palm Biomass as
the feedstock as the emission of CO2 from the plant is being compensated by the absorption
of CO2 by the palm tree at the plantation stage. Besides, the ammonia plant designed in a way
that the CO2 removed at the Carbon Dioxide Removal Stage is being sold to the nearby
Glycol Plant to be used for other application. The Marine Eutrophication Potential is found to
be greater for the production ammonia using biomass feedstock mainly due to the usage of
the pesticides during the plantation, harvesting of Palm Biomass. This can be overcome, by
reducing the usage of chemical pesticides and substituting it with organic chemicals. In
conclusion, the production of Ammonia using OPT Palm Biomass is highly sustainable
compared to that of the Natural Gas.
Full Report
4.2 Process Integration: Heat integration

4.2.1 Introduction
Ammonia production plant is one of the most energy intensive processes in chemical
industries. Hence it is of crucial importance to maximize energy savings by the means of heat
recovery and process integration. The recovery of energy will reduce huge amount of energy
waste and therefore has an impact on reduction cost for the energy. Consequently heat
integration will provide the plant with sustainable, economic and environmental benefits.
Thus in order to reduce the cost of the plant and achieve the sustainable viability by
substantial reduction in the usage of energy optimization and heat integration would be a
necessity.
4.2.2 Heat integration Approach

The pinch analysis method was used to integrate the energy efficiency of the plant. The
objective of heat integration and process optimization is to reduce the consumption of energy
by having a minimum capital investment. This goal can be achieved by minimizing the usage
of the utility and maximizing the heat transfer from process to process between the currently
existing hot and cold stream. One of the challenges faced during the heat integration is to
utilize the available heat transfer areas to the extent that is possible. Some streams and
equipment require a fixed mass flow rate and a constant firing duty. For instance, the cooling
utility circulating the HTWGSR and LTWGSR unit was also not included in the heat
integration as it is used to remove the fixed heat of reaction from the exothermic process.
Hence, the cooling water circulating the reactors in the shift reaction process were not
included in heat integration. Since some of the processes require the production of high
pressure process steam, the high pressure process steam was treated as a process stream and
not a utility stream. Therefore, the only utility that was considered was cooling water and
refrigerant inlet to heat exchangers.
There are four major design rules associated with the heat integration process for achieving
high efficiency; these four rules are stated below:
 Number of heaters present has to be minimized

 Number of coolers present has to be minimized
 The heaters present in the system can be replaced by process to process heat
exchanger
 The cooler present in the system can be replaced by process to process heat exchanger
If the steps mentioned above are followed as closely as possible the load on the cooler and
the heater duty can be reduced as more heat can be saved by process to process heat
exchangers. The cooling utilities include cooling water and refrigerant, while the heating
utilities include steam and hot water. There are a few sections within the ammonia plant that
consumes cooling utilities such as the refrigeration cycle and ammonia synthesis. The
sections requiring hot utilities would be mainly reformer and water gas shift reactor. Both of
these sections are combined together through pinch analysis in order to provide single heat
exchange network and utility diagram for the entire plant.
Full Report
Aspen Energy analyzer used is one of the most convenient approach as it provided the heat
exchanger network. After obtaining a few HEN designs from the Energy Analyzer the best
network design was chosen on the criteria of cost and the area of heat exchangers.
4.2.3Aspen Energy analyzer for the Heat integration

 The data for each stream has to be extracted these includes the temperature of hot and
cold stream, mas flow rate and the enthalpy of each stream. These data were collected
from the PRO II. These data were entered in the energy analyzer case under the
processing streams.
 The selection of utility is one the most important steps, the cold and hot utilities were
chosen such that it would be sufficient or the process. For this project cooling water
and refrigerant were chosen as the cooling utility and the steam was chosen as the hot
utility.
 The selection of high pressure steam as utility generated is important in order to
utilize the excess heat generated from this process
 The pinch temperature is then obtained from the target section of analyzer. The delta
T minimum was chosen to be . There are two pinch points for this process and
they are considered to be utility pinch.
 The recommended heat exchanger network designs were identified by Aspen Energy
analyzer.
 The most suitable design is chosen on criteria of economics and area.
The data Extraction table is shown below:
Table 4.2.3.5: Process stream data extraction table

Stream Type m (kg/h) MCp (kW/C) dH (kW) TS (degC) TT (degC) T*S (degC) T*T (degC)
Inlet of HX-201 H 2.70E+04 12.507905 1250.7905 150 50 145 45
Inlet of HX-202 H 2.70E+04 12.59427 1511.3124 150 30 145 25
Inlet of HX-204 H 5.82E+03 1.650813889 184.224227 161.596 50 156.596 45
Inlet of HX-205 H 5.82E+03 1.654852083 167.002708 150.917 50 145.917 45
Inlet of HX-206 H 5.82E+03 1.666966667 168.263615 150.94 50 145.94 45
Inlet of HX-207 H 5.82E+03 1.665351389 2.13831118 51.284 50 46.284 45
Feed to Low Temperature Water Gas Shift Reactor H 27035.92 14.95950029 8975.70017 950 350 945 345
Feed to HX-401 H 48298.77 0.000433056 0.06495833 200 50 195 45
Feed to HX-402 H 7.71E+04 0.000982917 0.0230494 73.45 50 68.45 45
Feed to HX-404 H 110600 0.000253125 0.02002219 119.1 40 114.1 35
Feed to HX-502 H 5.89E+03 6.835672222 402.962878 83.95 25 78.95 20
Feed to HX-601 H 5.78E+03 6.67038 1262.12928 234.214 45 229.214 40
Feed to HX-702 H 17501 16.83012833 4037.41315 289.892 50 284.892 45
Feed to HX-704 H 5.62E+03 6.976605556 69.7660556 800 790 795 785
Full Report
Table 4.2.3.2: Problem Table Algorithm

PTA Hot Streams
T*(degC) 1 2 3 4 5 6 7 8 9 10 11 12 13 14 dT Interval (°C) sum CP (kW/°C) dH (kW) Cummulative H
945 14.9595 0
14.9595 150 14.95950029 2243.925
795 14.9595 6.976606 2243.925043
14.9595 6.976606 10 21.93610585 219.3611
785 14.9595 6.976606 2463.286102
14.9595 440 14.95950029 6582.18
345 14.9595 9045.466229
60.108 0 0
284.892 16.83013 9045.466229
16.83013 55.678 16.83012833 937.0679
229.214 6.67038 16.83013 9982.534115
6.67038 16.83013 34.214 23.50050833 804.0464
195 0.000433 6.67038 16.83013 10786.58051
0.000433 6.67038 16.83013 38.404 23.50094139 902.5302
156.596 1.650814 0.000433 6.67038 16.83013 11689.11066
1.650814 0.000433 6.67038 16.83013 10.656 25.15175528 268.0171
145.94 1.650814 1.666967 0.000433 6.67038 16.83013 11957.12776
1.650814 1.666967 0.000433 6.67038 16.83013 0.023 26.81872194 0.616831
145.917 1.650814 1.6548521 1.666967 0.000433 6.67038 16.83013 11957.74459
1.650814 1.6548521 1.666967 0.000433 6.67038 16.83013 0.917 28.47357403 26.11027
145 12.507905 12.59427 1.650814 1.6548521 1.666967 0.000433 6.67038 16.83013 11983.85486
12.507905 12.59427 1.650814 1.6548521 1.666967 0.000433 6.67038 16.83013 30.9 53.57574903 1655.491
114.1 12.507905 12.59427 1.650814 1.6548521 1.666967 0.000433 0.000253 6.67038 16.83013 13639.34551
12.507905 12.59427 1.650814 1.6548521 1.666967 0.000433 0.000253 6.67038 16.83013 35.15 53.57600215 1883.196
78.95 12.507905 12.59427 1.650814 1.6548521 1.666967 0.000433 0.000253 6.835672 6.67038 16.83013 15522.54198
12.507905 12.59427 1.650814 1.6548521 1.666967 0.000433 0.000253 6.835672 6.67038 16.83013 10.5 60.41167438 634.3226
68.45 12.507905 12.59427 1.650814 1.6548521 1.666967 0.000433 0.000983 0.000253 6.835672 6.67038 16.83013 16156.86456
12.507905 12.59427 1.650814 1.6548521 1.666967 0.000433 0.000983 0.000253 6.835672 6.67038 16.83013 22.166 60.41265729 1339.107
46.284 12.507905 12.59427 1.650814 1.6548521 1.666967 1.665351 0.000433 0.000983 0.000253 6.835672 6.67038 16.83013 17495.97153
12.507905 12.59427 1.650814 1.6548521 1.666967 1.665351 0.000433 0.000983 0.000253 6.835672 6.67038 16.83013 1.284 62.07800868 79.70816
45 12.507905 12.59427 1.650814 1.6548521 1.666967 1.665351 0.000433 0.000983 0.000253 6.835672 6.67038 16.83013 17575.67969
12.59427 0.000253 6.835672 6.67038 5 26.10057535 130.5029
40 12.59427 0.000253 6.835672 6.67038 17706.18257
12.59427 0.000253 6.835672 5 19.43019535 97.15098
35 12.59427 0.000253 6.835672 17803.33354
12.59427 6.835672 10 19.42994222 194.2994
25 12.59427 6.835672 17997.63296
6.835672 5 6.835672222 34.17836
20 6.835672 18031.81133
Full Report
The grand composite curve and pinch temperature obtained from Problem Table Algorithm is
shown in the figure below:
Figure 4.2.3.1: Grand composite curve obtained from Energy Analyzer
Pinch analysis was carried out using all the streams connected to heat exchangers, heater and
coolers. Minimum temperature difference was considered to be 10 . The pinch temperature
obtained was 950 for the hot end and 940 for the cold end from the Aspen energy
analyser. The grand composite curve shown in Figure 4.2.3.1 indicates a “Threshold problem”
in which only heat removal is necessary. Therefore, the generation of high pressure utility
steam is encouraged in order to generate additional revenue to cover the cost of cooling water.
The following table provides information for the Heat exchanger network design obtained
from PRO II. This table below shows the operating costs, capital costs and total heat
exchange area for the proposed heat exchanger network:
Table 4.2.3.2: Heat Exchanger Network design From Aspen Analyzer
Full Report
Figure 4.2.3.2: Optimum Heat Exchanger Network Design
Figure 4.2.3.2 shows the optimum heat exchanger network design which is a future
recommendation to our current process design network.
Full Report
CHAPTER 5 | DETAILED PROCESS AND EQUIPMENT DESIGN
5.1 Detail and Mechanical Design: Autothermal Reformer (R-201)

5.1.1 Definition of Design and Specification
Autothermal Reformer (ATR) has a feature of standalone air fired reformer. ATR involved
the best feature of both partial oxidation reaction (Exothermic reaction) and steam reforming
reaction (Endothermic reaction) in a ceramic refractory lined vessel. ATR has an important
characteristic that no external heat source is required during the steam reforming process.
ATR is widely used in methanol, ammonia as well as oil and gas industry in order to produce
hydrogen-rich gas by undergo steam reforming process in the present of catalyst. In Altenis
BioAmmonia plant, hydrogen is the main feedstock of production of Anhydrous Ammonia.
Therefore, the 9800 kg/h of cleaned syngas at average temperature of 736.82°C and elevated
pressure of 25 bar will be entering ATR. The syngas contained of 796.96 kg/h of methane
which required to reform into hydrogen-rich gas. The industrial scale autothermal reformer
has methane conversion range of 96 to 99 percent conversion. The high efficiency of methane
conversion will be able to supply the ammonia plant with sufficient feedstock.
The mixture of syngas and steam at the temperature of 850°C will be entering the ATR with
the 650°C preheated compressed air at pressure of 25 bar. In the burner section, the burner
provides mixing of the feed and the air. The syngas and air are entering the reactor at high
speed and leaves little room for efficient mixing(Andersson and Nilsson, 2013). The ignition
of the feed and the air are mainly affected by the air flow entering the reformer. Under the
high pressure and temperature operating condition, ignition will occur when adiabatic flame
temperature is achieved(Kondratiev, 2013).
In the partial oxidation reaction, the components that involved in the reactions are carbon
monoxide and hydrogen. The reactions of these components are shown as reaction equations
below:
Based on literature studies, in a simultaneous combustion of hydrogen and carbon monoxide,

the ratio of rate of combustion of carbon monoxide to rate of hydrogen is 1:2.86(R.T.Haslam,
Full Report
1923). By calculation, the conversion of carbon monoxide is 26% and the conversion of
hydrogen is 74%.
The syngas leaving combustion zone at the temperature 1186°C. The heat generated in the
combustion zone will be utilized by the endothermic reaction in the steam reforming zone. At
this stage, the components that are involved in the reactions in the present of Nickel-based
catalyst are methane, ethane, ethylene, propane as well as carbon monoxide. The reactions
are as follow:
At this stage, all the hydrocarbons will be reacted through equilibrium conversion and
hydrogen will be produced. Approximately 97% of the methane entering the reformer is
converted and produce hydrogen gas. Apart from that, ethane, ethane, ethylene and propane
are fully converted. The syngas is leaving the reformer at temperature and pressure of 950°C
and 2.45 bar respectively.
In this case, the high pressure and temperature operating condition affecting the design of
autothermal reformer. Due to its high temperature and pressure, a greater investment is
required to ensure safe design. Therefore, all these issues are taken into consideration in
designing the reformer. In designing the autothermal reformer, the design temperature and
pressure are 1350°C and 2.75MPa. The combustion zone which is the conversion reaction
section is sized based on the adiabatic plug flow reactor where there is no pressure drop
involves in the reactions. Besides, the reactions that are taken into account are the
aforementioned reactions: (1) and (2). The main purpose of having an autothermal
reformer is to produce hydrogen-rich gas by reforming the methane in the syngas. Therefore,
when sizing the catalytic zone, the main reactions that are considered are reactions as shown
in (3), (4) and (5). Polymath software is used to determine the conversion of
Full Report
the reactions, height of the catalytic bed, weight of catalyst, volume of the reactor conversion
as well as the pressure drop in the reactor.
5.1.2 Basis of Performance

As mentioned previously, autothermal reformer consists of two sections which are the partial
oxidation reaction and steam reforming reaction. The operating condition and reactions
involved in each section are different. Therefore, when sizing the reformer, it will be carried
out separately. In this case, the combustion zone will be sized based on the adiabatic plug
flow reactor (Conversion Reactor) where no pressure drop is present in the reactor. On the
other hand, the catalytic or steam reforming zone will be sized based on catalytic packed bed
reactor (Equilibrium Reactor). The conversion of each component in each reactor will be
determined by using Polymath Software. The detailed steps in sizing each reactor are
explained as follow:
5.1.3 Sizing of Autothermal Reformer

Table 5.1.3.1 Summary of Autothermal Reformer Design and Sizing
Summary of Autothermal Reformer Design and Sizing
Shape and orientation Cylindrical and vertical
Conversion Reactor Height 0.50m
Equilibrium Reactor Height 0.64m
Total Height 1.14m
Inner Diameter 1m
Weight of catalyst 500kg
5.1.3.1 Sizing Conversion Reactor (Combustion Zone)

Partial oxidation reactions in adiabatic reactor are irreversible reactions. In this case, limited
air is supplied to the reactor to achieve partial oxidation reactions. The inlet flows into
conversion reactor are obtained from the Pro II. Steps in sizing the conversion reactor are
shown as below:
Step 1: Define assumptions
Assumptions:
 Only hydrogen and carbon monoxide are involved conversion reactor’s reactions
 Hydrocarbon such as methane, ethane, ethylene and propane does not involve in the
partial oxidation. This is because the combustion of hydrogen and carbon monoxide
Full Report
are way faster than the combustion of hydrocarbon. Thus, hydrocarbon would not
involve in the reactions(Nanthagopal et al., 2011).
 Reactor is in cylindrical shape
 Diameter of the reactor is 1 meter
 No work done on the system
 Adiabatic operation
Step 2: Identify the reactions involved
Reactions involved:
Step 3: Determine mole balance, rate law, stoichiometry (gas phase reaction with no pressure
drop).
Step 4: Determine the molar flow for each component involved in the reactions, volumetric
flow and average of initial temperature. Determine all the specific heat capacity of each
component involved in the reactions and determine the in order to determine the final
temperature of the reactions.
Step 5: Include all the equations and unknowns in the polymath software and determine the
conversion of the reactions as well as the volume of the reactor.
Summary of results from polymath software

Conversion of each Component:
Table 5.1.3.1.1 Conversion of Each Composition in the Conversion R eactor

Component Conversion Inlet Outlet Outlet(kmol/h) (Pro II)
(kmol/h) (kmol/h)
Oxygen 1 40.63 0 0
Hydrogen 0.1152 183.33 162.21 120.68
Water -0.0331 637.22 658.31 699.93
Carbon 0.4818 124.81 64.68 102.80
Monoxide
Carbon Dioxide -0.6022 99.87 160.01 121.88
Full Report
The volume of conversion reactor is 0.402 m3. The height of the reactor is calculated using
cylindrical shape volume where the height of the reactor is calculated to be 0.5m.
Detailed calculations are shown in C1.1
5.1.3.2 Sizing Equilibrium Reactor (Steam Reforming /Catalytic Zone)

Sizing the equilibrium reactor based on the adiabatic packed bed reactor. The components
that involved in the reactions in this stage include methane, ethane, ethylene, propane and
carbon monoxide. However, the amount of ethane, ethylene and propane in the syngas are
very little. To simplify the calculation these three components are assumed fully converted in
the first place when entering the catalytic bed since the operating temperature is very high
which is around 1186°C. The weight of the catalyst used as well as the height of the catalyst
in the reactor will be determined using polymath software in order to size the reactor.
Assumptions that taken into consideration during sizing are:
Step1: Define assumptions
Assumptions:
 The composition of ethane, ethylene and propane are too little and it is assumed to be
fully converted in the first place when entering the catalytic bed since the operating
temperature is very high.
 Only methane and carbon monoxide are involved equilibrium reaction.
Step 2: Identify the reactions involved
Reactions involved:
Step 3: Determine design equations, rate law and stoichiometry. In equilibrium reactor,
catalyst will be the main consideration in sizing the reactor. Thus, several parameters such as
beta, alpha, porosity of catalyst, viscosity and diameter of particle need to be determined in
order to fulfill the unknowns in the design equations.
Full Report
Step 4: Determine the molar flow for each component involved in the reactions and initial
temperature. Determine all the specific heat capacity of each component involved in the
reactions in order to determine the final temperature of the reactions.
Step 5: Include all the equations and unknowns in the polymath software and determine the
conversion of the reactions as well as the volume of the reactor.
Summary of results from polymath:
Table 5.1.3.2.1: Conversion of Each Component in the Equilibrium Reactor

Component Conversion Inlet Outlet Outlet(kmol/h) (Pro II)
(kmol/h) (kmol/h)
Methane 0.9918 49.81 0.408 1.75
Water 0.1027 699.93 628.05 584.36
Carbon -0.2620 102.80 129.73 148.11
Monoxide
Carbon Dioxide -0.1844 121.88 144.36 157.01
Hydrogen -1.4143 120.68 291.36 372.72
The height of the catalyst calculated using Polymath is 0.64m
The reactor diameter is assumed to be 1 meter. Therefore, the volume of the reactor is
calculated to be 0.5 m3.
Besides that, the weight of the catalyst obtained from Polymath is 500 kg.
Detailed calculations are shown in Appendix C1.2
5.1.4 Catalytic Bed Specification

Table 5.1.4.1 Catalyst Specifications
Catalyst Type Nickel Based Catalyst Catalyst Volume 0.5m3
with MgAl2O4 as
carrier
Catalyst Shape Cylinder with 7 holes Catalyst Bed Bulk 1000 kg/m3
Density
Catalyst Outer 16mm Catalyst Bed Void 0.4652
Diameter Fraction
Catalyst Inner 3mm Type of Catalyst Support Denstone 99 High
Diameter Alumina
Catalyst Bed Depth 0.63m Size of Catalyst Support 19 mm
(AL-Dhfeery and Jassem, 2012)
As mentioned in Chapter 2, nickel based catalyst with MgAl2O4 as carrier is chosen as it is

widely used in industry and the price of this catalyst reasonable and provide satisfy results.
Full Report
Besides that, due to the reason that autothermal reformer operates at high temperature up to
1300°C(Padban and Bacher, 2005). Denstone 99 High Alumina Catalyst support is used
because it is able to withstand high operating temperature up to 1650°C and it is able to
withstand corrosion and thermal shock(Norpro, 2013).
5.1.5 Burner
The factors that affecting the performance of the burner are the air inlet flow, angle of nozzle
as well as the size of the tube connecting to the burner. Syngas and air inlet are mixing at the
burner. Therefore, one of the ways in optimizing the burner design is improving the mixing
of the burner. Ring type burner designed by Topsoe is used in this reformer. Research has
shown that it is very efficient for reformer like autothermal reformer. The angle of inlet
nozzle of burner by Topsoe is designed with respect to the main process gas stream. This type
of nozzle improves the mixing process and thus enhances the performance of the burner
which indirectly improves the performance of the reactor.
5.1.6 Mechanical Design

The main assumptions and calculations steps for mechanical design of autothermal reformer
are discussed and shown in the following section. However, the detailed calculations and
assumptions related to the autothermal reformer are shown in Appendix C1.3.
5.1.6.1 Material of Construction

Autothermal reformer operating under elevated temperature and pressure. Thus, it is very
crucial in choosing the type of material used. Since the reformer is operating at high
temperature up to 1300°C. Alumina Silicate refractory ceramic lining is chosen as the
construction material for the inner vessel of this reformer. Alumina Silicate Refractory has a
density of 2300kg/m3 and lower thermal conductivity of 1.3W/m.K It is able to withstand the
operating temperature at a range of 1300-1700°C (ToolBox, 2013). Then, the reformer will
be coated with a thin layer stainless steel to protect the refractory lining from creating voids
and or lead to hot spots and lining failure. This is because under high temperature operation,
the hotface will thermally elongate and bend toward the heat. Thus, stainless steel is acting
like a lining to tie back the refractory line to the steel shell(Grigson et al., 2009). Stainless
steel chosen is stainless steel 310S. It has density of 8000 kg/m3 and able to withstand the
temperature up to 1400°C(Steelguru, 2011). Epoxy-based paint will be used to paint on the
outer layer of stainless steel to prevent corrosion. Besides that, the outermost layer will be
insulated using mineral wool which has thermal conductivity of 0.04 W/m.K and density of
130 kg/m3.
Full Report
5.1.6.2 Mechanical Design of Shell of Autothermal Reformer

As mentioned in Section 5.1.6.1, there are two layers of shells and a layer of insulation. The
innermost layer is the Alumina Silicate Refractory Lining followed by a layer of stainless
steel and then insulated by using mineral wool. The properties of each lining are shown in
table below:
Table 5.1.6.2.1 Properties of Material of Construction

Type of Material Alumina Silicate Stainless Steel Mineral
Lining 310S Wool
Density (kg/m3) 2300 8000 130
Thermal Conductivity, k 1.3 14.2 0.04
(W/m.K)
Safe Temperature 1300-1700 1400-1450 -
Range, °C
Tensile Strength (MPa) 55MPa 520 MPa -
Modulus of Elasticity 14000 200 GPa -
5.1.6.2.1 Design Pressure, temperature and maximum allowable stress

To begin with the mechanical design, the design temperature and pressure are determined. It
is calculated by taking 50°C above the operating temperature and 10% above operating
pressure. For a reformer that operates at temperature of 1300°C and 25 bar, the design
temperature and pressure are determined to be 1350°C and 2.75MPa. Since stainless steel is
used for the purpose to tie back the refractory lining to the shell, both maximum allowable
stress for refractory lining and stainless steel are determined. According to Sinnot, the
maximum allowable stress is the lowest for the case below:
1. The specified minimum tensile strength at room temperature divided by 3.5

2. Tensile strength at temperature divided by 3.5
The maximum allowable tensile strength for inner shell is 15.71 MPa and the maximum
allowable tensile strength for the outer shell is 148.57 MPa.
Detailed Calculations showed in Appendix C1.3.1.
5.1.6.2.2 Thickness of Refractory Lining

Based on research, it is good practice to ensure that the high temperature furnace wall such as
reformer line with 115mm thickness of refractory lining. Therefore, in this case, the thickness
of the refractory lining is taken to be 115 mm(Bhatia, 2011).
Full Report
5.1.6.2.3 Thickness of Stainless Steel

The main purpose of having a lining of stainless steel is meant to protect the refractory lining
from creating voids and or lead to hot spots and lining failure. This is because under high
temperature operation, the hotface will thermally elongate and bend toward the heat. Thus,
stainless steel is acting like a lining to tie back the refractory line to the steel shell(Grigson et
al., 2009). . Therefore, a thin layer of stainless steel will be sufficient to coat the outer layer
of refractory lining. Stainless steel 310S is chosen due to its property of able to withstand a
very high temperature up to 1400°C(Siebert et al., 2008). The available stainless steel 310S
sheet available in the market has a thickness of 18mm and 20mm thickness of the corrosion
allowance is reserved(Sinnott and Towler, 2009). Therefore, the thickness of the stainless
steel used is 20mm.
Properties of Stainless Steel 310S
Density: 8000 kg/m3
Thermal Conductivity, k: 14.2 W/m.K
(Steelguru, 2011)
5.1.6.2.4 Thickness of the Insulator

The insulator used in this reactor is mineral wool. Mineral wool is cheaper compare to other
insulator and provides sufficient insulation. The thickness of the insulator is determined by
using the energy balance for heat transfer. First Law of Thermodynamic is taken into
consideration when determining the insulation thickness. It is assumed that no heat loss
during the heat transfer and all the heat from conduction area will be fully transferred and
enter the convection area. In order to use this method, the total height of the reactor is
determined. The total height of the reactor involved conversion and equilibrium reactor,
ellipsoidal head, nozzle diameter, catalyst support as well as the thickness of the refractory
line, stainless steel and insulator. By equating the heat of conduction to heat of convection,
the thickness of the insulation is calculated to be 0.17 m. After the insulator is calculated, the
total height of the reactor is 3.81m. Detailed calculations show in Appendix C1.3.2.
5.1.6.2.5 Ellipsoidal Head

Ellipsoidal head is chosen as the head and bottom of the reactor. The standard ellipsoidal
head that usually manufactured has a major and minor axis ratio of 2:1. It is chosen over
other types of head such as hemispherical head and torispherical head because it is the most
Full Report
economical option as it is only require approximately half of the thickness as compare to

other choice.
5.1.7 Stress Analysis of Autothermal Reformer

Stress analysis is carried out to ensure that the designed reformer is able to sustain when
subjected to other loads in addition pressure. Besides that, through stress analysis, it is able to
determine that if the reformer designed is able to withstand the worst combination of loading.
According to Sinnott the main source of dead weight loads are vessel shell, vessel fittings,
manways, nozzles, internal fittings and external fittings.
For autothermal reformer, there are two sections of reactor as mentioned before. There are the
inner vessel and outer vessel which are the Alumina Silicate refractory lining and the
stainless steel vessel. The approximate weight of vessel is calculated by using the equation
below:
The weight of the inner vessel is calculated to be 42.41 kN and the outer vessel is calculated
to be 29.74 kN respectively.
Besides vessel weight, the dead load that will be taken into consideration is the weight of
catalyst, weight of insulation as well as weight of fitting such as caged ladder and platform.
The weights of each dead load are 4.905kN, 2.029 kN and 8.829 kN. The total dry weights
are summarized as below:
Table 5.1.7.1 Dead Weight of the Vessel
Dead weight of vessel 72.14852447
Weight of Catalyst 4.905
Weight of Fittings 2.02932459
Weight of Insulation 8.829390874
Total Dry Weight (kN) 87.91223993
Detailed calculations are shown in Appendix C1.3.
5.1.8 Mechanical Design Feasibility Testing of Inner Shell (Refractory Lining) and
Outer Shell (Stainless Steel) of Autothermal Reformer
Stress analysis such as hydraulic pressure testing, wind loading and stress analysis are carried
out to determine the feasibility of the reactor. Hydraulic testing is carried out to test the
strength of the reactor, material quality, sealing arrangement and leaks. Water is used as the
medium to do the testing. Based on calculations, the total weight during hydrotesting is
164.968 kN. However, in normal operating condition, gas will be the medium in this case. It
is able to sustain as the weight of the gas is lighter than water.
Full Report
In order to calculate the wind loading of the autothermal reformer, the height of the reactor
and the skirt are taking in account when carry out the wind loading test. Taking the wind
speed of 160 km/h, the wind pressure will be 1280 N/m2(Sinnott and Towler, 2009). From
calculation, the wind load per meter length is 2.584 kN/m. The total height including the skirt
is calculated to be 5.005 m. Therefore, the bending moment at the bottom of the tangent line
is 32.338 kNm.
The analysis of stress can be summarized in the table below:

Table 5.1.8.1 Summary of Stress Analysis of Refractory Lining and Stainless
Steel
Stress Analysis Refractory Lining Stainless Steel
Hydraulic Pressure testing, 77.056 44.490
kN
Pressure testing, kN 164.968 132.402
Wind Loading, kN/m 2.584 2.584
Bending Moment. kNm 32.338 36.521
Analysis of Stress
Longitudinal Stress, , MPa 5.978 42.281
Hoop Stress, , MPa 11.957 84.563
Dead Weight Stress, , 0,218 1.119
MPa
Dead Weight Stress, , 0.409 1.685
MPa
Bending Stress
Iv, mm4 2.879×1011 2.247×1011
Bending Stress, MPa 0.069 0.103
Buckling Test, , MPa 1136.68 247105.223
Maximum Compressive 0.287 1.223
Stress, MPa
Resultant Longitudinal Stress
, MPa 5.829 41.265
, MPa 5.691 41.059
Greatest Difference betweem 6.265 43.504
principal stress at downwind,
MPa
Maximum Allowable Stress, 15.71 148.57
MPa
Detailed calculations are shown in Appendix C1.4 and C1.5.
5.1.9 Mechanical Design of Vessel Support - Skirt

The method used to support the vessel is usually dependent on the size, shape and the weight
of the reactor vessel. The maximum dead load of the reformer in normal operating condition
and hydrotesting are calculated to be 88.1869 kN and 164.627 kN respectively. According to
Full Report
AS 1210-2010 Clause 3.24.3.3 about skirt support, if the product diameter thickness and
temperature at the top of the skirt exceed 16×106. However, based on calculation no
discontinuity stress test is required for the criteria of skirt used in this case.
Therefore, the stress analysis for the skirt is carried out and the table below shows the results
from stress analysis:
Table 5.1.9.1 Stress Analysis of Skirt

Stress Analysis of Skirt
Bending Stress in Skirt, MPa 1.758
Dead Weight Stress in the Skirt, kN 164.627 kN (Under Pressure Testing)
88.187 kN (Dry Weight)
Dead Weight in the Skirt – 3.127
Hydropressure, MPa
Dead Weight in the Skirt – Normal 1.723
Operation, MPa
Maximum Tensile Strength, MPa 0.034
Maximum Resultant Compressive, 4.975
MPa
Testing for Tensile Stress Allowable Tensile Stress = 560MPa
Tensile Stress of Skirt = 3.127 MPa
560MPa>3.127MPa, Passed
Testing for Compressive Stress Allowable Compressive Stress
=154.44MPa
Max. Tensile Stress of Skirt = 0.034 MPa
154.44MPa >0.034 MPa, Passed
5.1.10 Pipe selection and pipe sizing

The pipe selection in autothermal reformer is very dependent on the sustainability of the pipe
to high temperature and pressure. In this case, the inlets and outlet of the reformer are at a
very high temperature. Therefore, stainless steel 310S will be used for these three sections.
However, for the streams that have a lower temperature stream flow, carbon steel will be
used. The calculations of sizing of the pipe are shown in Appendix C1.3.2. The table below
shows the summary for the pipe size of the major streams.
Streams Nominal Pipe size, mm

Syngas Outlet 440
Syngas and Steam Inlet 370
Air Inlet 150
Full Report
5.1.11 Drawing
Autothermal Reformer IEM.vsd
Full Report
5.1.12 Datasheet of Autothermal Reformer
AUTOTHERMAL REFORMER
Mechanical Specification Sheet
Company Alternis BioAmmonia Pvt. Ltd. Description Autothermal Reformer

Equipment No. R-201 Sheet No. 1
Number Required = 1 Functions Reforming Methane to Hydrogen –rich gas
Operating Data and condition
Fluid Allocation Inlet Outlet
Stream 30 33 34
Mass Flow Rate, kg/h 21220.87 5815.10 27035.94
Mass Fraction
Methane 0.0377 0 0.0010
Ethane 0.0092 0 0
Ethylene 0.0099 0 0
Propane 0.0031 0 0
Steam 0.5410 0 0.3875
Carbon Monoxide 0.1647 0 0.1506
Carbon Dioxide 0.2071 0 0.2601
Hydrogen 0.0174 0 0.0280
Nitrogen 0.0099 0.7671 0.1727
Oxygen 0.0000 0.2329 0
Phase Vapour Vapour Vapour
Temperature, °C 5 650 930.90
Pressure, bar 25 25 24.5
Design Temperature, °C 1350
Design Pressure, bar 27.5
Material Construction
Material of Inner Vessel Alumina Silicate Refractory
Material of Outer Vessel Stainless Steel 310S
Material of Insulation Mineral Wool
Material of Base Support Carbon Steel AISI 2055
Type of Coating Epoxy-based paint
Dimension and Design of Autothermal Reformer
Specification Data Specification Data
Reactor Shape Cylindrical Total Reactor Height 5.005
(Including Skirt), m
Reactor Orientation Vertical Type of Top and Bottom Heads Ellipsoidal Head
Inner Vessel Thickness, mm 115 Height of Top Head, m 0.0825
Outer Vessel Thickness, mm 20 Height of Bottom Head, m 0.25
Insulation Thickness, mm 170 Height of Skirt, m 1.5
Inner Vessel Diameter 1.23 Number of Ladder 1
(Include Thickness), m
Outer Vessel Diameter 1.27 Height of Ladder, m 3.6
(Include Thickness), m
Outer Diameter 1.61 Number of Manhole Pathway 1
(Including Insulation), m
Height of Catalytic Zone, m 0.64 Diameter of Manhole Pathway, m 0.5
Height of Combustion Zone, m 0.5 Number of Platform 2
Height of Catalytic Support 0.365 Type of Burner Ring-type Burner
Full Report
Catalytic Bed
3
Catalyst Type Nickel Based Catalyst Catalyst 0.5m
Volume
Catalyst Shape Cylinder with 7 holes Catalyst Bed 1000 kg/m3
Bulk Density
Catalyst Outer Diameter 16mm Catalyst Bed 0.4652
Void Fraction
Catalyst Inner Diameter 3mm Type of Denstone 99 High Alumina
Catalyst
Support
Catalyst Bed Depth 0.63m Size of Catalyst 19 mm
Support
Stress Analysis of Refractory Lining
Maximum Allowable Stress at Design 15.714 Axial Dead Weight Stress (Compressive), 0.218
temperature, MPa MPa
Total Reactor Dead Weight, kN 87.747 Axial Compressive Dead Weight Stress 0.409
(Pressure Testing), MPa
Weight of water (Hydrotesting), kN 76.879 Critical Buckling Stress, MPa 1136.684
Total weight (Pressure Testing), kN 164.627 Maximum Compressive Stress 0.287
Wind Loading, kN/m 2.584 Maximum Bending Stress at upwind 5.830
Condition, MPa
Bending Moment (Tangent Line), kNm 32.338 Maximum Bending Stress at downwind 5.691
Condition, MPa
Hoop Stress due to Internal Pressure, MPa 11.957 Resultant Axial Stress at Upwind 6.127
Condition, MPa
Axial Stress due to Internal Pressure, MPa 5.978 Resultant Axial Stress at Downwind 6.275
Condition, MPa
Stress Analysis of Stainless Steel
Maximum Allowable Stress at Design 148.57 Axial Dead Weight Stress (Compressive), 1.119
temperature, MPa MPa
Total Reactor Dead Weight, kN 87.912 Axial Compressive Dead Weight Stress, 1.686
MPa
Weight of water (Hydrotesting), MPa 44.490 Critical Buckling Stress, MPa 247105.223
Total weight (Pressure Testing), MPa 132.402 Maximum Compressive Stress 1.223
Wind Loading, kN/m 2.584 Maximum Bending Stress at upwind 41.265
Condition, MPa
Bending Moment (Tangent Line), kNm 36.521 Maximum Bending Stress at downwind 41.058
Condition, MPa
Hoop Stress due to Internal Pressure, MPa 84.563 Resultant Axial Stress at Upwind 43.298
Condition, MPa
Axial Stress due to Internal Pressure, MPa 42.281 Resultant Axial Stress at Downwind 41.265
Condition, MPa
Stress Analysis of Skirt
Maximum Allowable tensile , MPa 560 Maximum Resultant Tensile, MPa 0.041
Maximum allowable Compressive, MPa 154.44 Maximum Resultant Compressive, MPa 4.351
Bending Stress in the Skirt, MPa 1.764 Dead Weight Stress in Skirt 2.588
Hydropressure, MPa
Full Report
Total Weight –Pressure Testing, kN 132.402 Dead Weight Stress in Skirt Normal 1.723
Operation, MPa
Total Dry weight, kN 88.167
Nozzle Specification
Nozzle No. Nominal Size, mm
N1 Syngas and Steam Inlet, m 370
N2 Air Inlet, m 150
N3 Syngas Outlet, m 440
Full Report
5.2 Detailed Process and Mechanical Design of Low Temperature Water-Gas

Shift Reactor
5.2.1 Definition of Design and Specification for Low Temperature Water-Gas Shift Reactor
(LTWGSR)
Since the presence of carbon oxides will poison the catalyst of the ammonia reactor
downstream, the hydrogen (H2) supplied to the ammonia reactor needs to be free of carbon
monoxide (CO) and carbon dioxide (CO2). Therefore, the first stage for the removal of CO
comes under the water-gas shift reaction. Further removal of CO will then be done
downstream in the methanator before the hydrogen gas is fed into the ammonia reactor. Thus,
in order to achieve a balance between the effect of Le Chatelier’s principle and the effect of
Arrhenius law which are the constraints to the shift reactor design, Alternis BioAmmonia Sdn.
Bhd. has decided to utilize a series of High Temperature Water- Gas Shift Reactor
(HTWGSR) followed by a Low Temperature Water-Gas Shift reactor (LTWGSR) with
intercooling stage, so that the task of CO removal could be executed along with a higher
purity of H2 in syngas (Smith, et al., 2010; Lima, et al., 2012). However, the focus of this
section of the report will be mainly on the design of Low Temperature Water Gas Shift
Reactor which is responsible for the further increase of the overall conversion of CO to
98.8%.
One top of that, the factors that should be considered in the design of the reactor for safe
operation include the high operating pressure of the reactor vessel, the exothermic reaction
between the carbon monoxide and steam and the optimum temperature at which the catalyst
is working at its best without deactivation. Furthermore, the design of the reactor should
consider the worst possible scenario such as runaway reactions. Also, the overall percentage
conversion target of CO that needs to be met is more than 98. It is also important to maintain
a constant temperature in order to prevent temperature rise along the catalyst bed which is
unfavourable as the equilibrium conversion and ultimately the product selectivity will be
affected (Jakobsen, 2008; Eigenberger, 1992). The design of the reactor should also be able
to resist the effect of monsoonal climate changes, corrosion and wear in order to last for 25
years as which is the operating lifespan of the plant.
The temperature, pressure and flow rates of the inlet and outlet streams to the low
temperature water-gas shift reactor are presented in the table below.
Table 5.2.1: Specification of Inlet and Outlet operating data of LTWGSR

Syngas
Component Inlet Outlet
Units Mass flow rate Mass flow rate
N2 kg/h 4668.1124 4668.1124
H2 kg/h 1009.986971 1038.213856
Steam kg/h 8137.566088 7883.524123
CO kg/h 443.2169778 48.04058759
CO2 kg/h 12728.53107 13349.52254
CH4 kg/h 28.05412125 28.05412125
C2H4 kg/h 0.003360317 0.003360317
Full Report
C3H6 kg/h 0.003996227 0.003996227

C3H8 kg/h 0.19559681 0.19559681
Total kg/h 27015.67 27015.67
Operating Parameters
Pressure kPa (abs) 2020 1758
Temperature 200 200
Property
Vapour fraction - 1 1
Density kg/m3 11.56 11.55
-5
Viscosity Pa s 1.963 10 1.962 10-5
Thermal conductivity W/m K 6.376 10-2 6.558 10-2
Utility
Component
Cooling Water kg/h 3076.51 3076.51
Operating Parameter
Pressure kPa (abs) 101.33 101.33
Temperature 25 70
Property
Density kg/m3 1007 972.5
Viscosity Pa s 8.904 10-4 4.004 10-4
Thermal conductivity W/m K 6.110 10-2 6.623 10-2
Mass heat capacity J/ kg 4043 4056
The syngas entering the reactor is not considered as corrosive and does not contain any lethal
components that might poison the catalyst in the reactor. Thus the design of the reactor is
only bound to the operating conditions and the pressure subjected to the reactor due to its
content.
Treated cooling water from the cooling tower is used to circulate the isothermal reactor in
order to maintain the reactor temperature at 200 . Even so, fouling in the shell side of the
reactor is inevitable in a long run. That is why the design of the shell side of the heat
exchanger has taken a fouling factor of 0.00003 into consideration (Sinnott &
Towler, 2009).
5.2.2.1 Background and Principle of the Mechanism of the LTWGSR

The reaction that governs the CO upgrade to H2 is the water-gas shift (WGS) reaction and it
has gain wide industrial application in the refining process of synthesis gas. The WGS
reaction expressed below is an equilibrium-limited, heterogeneous and exothermic reaction
whereby it is thermodynamically favoured at low temperatures and kinetically favoured at
high temperatures (Smith, et al., 2010).
Full Report

is no change in the volume from reactants to products (Smith, et al., 2010). In contrast,
Arrhenius law which explains the temperature dependence of the specific reaction rate
constant in chemical reactions requires the reactants to gain a minimum amount of energy
called activation energy Ea by increasing the reaction temperature so that the forward reaction
of H2 can occur (Lima, et al., 2012).
5.2.2.2 Justification on the selection of catalyst
Copper-based catalyst which contains a mixture of ZnO, CuO and Cr2O3/ Al2O3 is used as the
catalyst for LTWGSR (Callaghan, 2006). This type of catalyst is able to remain active at
temperatures as low as 200 due to the fact that it is susceptible to thermal sintering at
higher temperatures of more than 300 (Smith, et al., 2010). Furthermore, this type of
catalyst is commercially available with a normal operating span of 2 to 3 years as it is used in
many chemical industries producing syngas. The zinc oxide present in the catalyst also helps
provide additional protection to the copper from sulphur poisoninh while acting partially as a
support for the copper (Callaghan, 2006). Besides that, this trait of the catalyst that has
selectively fewer side reactions when the system is operating at higher operating pressures is
crucial when it comes to maintaining the quality of syngas produced so that the operating
conditions for the downstream processes are met. The bulk density of catalyst is1422kg/m 3
for the commercial copper-based catalyst (Morabiya & Shah, 2012). The equivaent diameter
of catalyst on the other hand is 230 (Smith, et al., 2010).
5.2.2.3 Justification on the selection of packed bed

The water-gas shift reaction is an exothermic process and it is generally conducted in an
insulated adiabatic reactor with temperature increasing along the catalyst bed due to the
exothermic process (Callaghan, 2006). The temperature rise along the catalyst bed is
unfavourable as it may affect the equilibrium conversion, the product selectivity, the
deactivation of catalyst and in extreme cases unsafe operation due to runaway reactions
(Jakobsen, 2008; Eigenberger, 1992). Essentially, the type of reactor chosen for this
heterogeneous catalytic process is an isothermal multi-tube fixed bed with cooling water
circulation in order to keep the reactor isothermal. Hence, the temperature increase per bed is
limited by the design of a multi-tube reactor that is able to contain hundreds or thousands of
tubes with an inside diameter of only a few centimetres and maximize heat transfer to cooling
feed water that will ultimately prevent excessive temperatures and hot spots (Jakobsen, 2008).
On top of that, the multi-tube fixed bed reactor is able to accommodate stack of catalyst
pellets that are compact and immobile within a vertical vessel.
Therefore, the design of the multi-tube isothermal fixed bed reactor consists of a tube side
and a shell side whereby catalysts are packed in tubes while water is fed to the shell side and
used to circulate the tube bundle in order to remove heat from the exothermic reaction.
Full Report
5.2.2.4 Justification on the selection of materials for the construction of pressure vessel and
bed internals such as tube plates and baffles
The specifications of steel for the construction of the reactor are shown in Table 6.2.4 below.
Table 5.2.2.4: Steel specifications and maximum allowable stress for vessel
under ASME BPV code (Sinnott & Towler, 2009)
Material Grade Temperature Minumum Allowable Tensile

Strength (MPa)
Low Alloy Steel (1 ¼ Cr, ½ Mo, Si) A387 220 114.45
Gr 22
The compositions of Grade A387 Low Alloy Steel are (in weight percentage) 0.15% Carbon
(C), 0.30-0.60% Manganese (Mn), 0.035% Phosphorus (P), 0.035% Sulfur (S), 0.50% Silicon
(Si), 2.00-2.50% Chromium (Cr) and 0.90-1.10% Molybdenum (Mo).
This steel is selected instead of carbon-manganese A285 Gr A because it has a higher

minimum allowable tensile strength which is strong enough to support the weight of the
catalysts in the reactor including the weight of cooling water along with internal and external
fittings. Hence, low alloy steel has sufficient tensile strength to withstand the maximum
tensile force exerted on the vessel and is able to eliminate major hazards such as buckling and
rupture under fluctuating vessel load due to varying water flow, process gas flow and workers
during maintenance.
Moreover, the low carbon content of 0.15 weight percent promotes ductility and weldability
of the steel (Delmarlearning, 2006). It is also a cheaper alternative compared to stainless steel
as it is also able to last for an operating life span of 25 years and withstand corrosion in a long
run as it.
5.2.2.5 Justification on the type of end used

The type of end selected is a hemispherical head. This type of head has the strongest shape
compared to other types of shape as it has a stronger resistance to abrupt elevated pressures in
the vessel which is a good safety aspect of the design (Sinnott & Towler, 2009). However,
the cost of this type of head is very expensive due to its larger size compared to other heads,
hence, more material is needed to fabricate this head. In spite of the high costs due to the
amount of material used for the construction of the head, fabrication of a hemispherical head
is rather easy as it both ends are symmetrical to one another.
5.2.2.6 Justification on the placement and type of insulation used

The insulation for the pressure vessel is only placed at the hemispherical ends of the reactor
vessel as the temperature of syngas entering and leaving the reactor at the ends are at 200 .
Insulation for the cylindrical part of the vessel is not necessary as water is used to circulate
the isothermal reactor keeping the wall of the reactor at a mean temperature of 47.5 . The
insulation of the pressure vessel at the hemispherical ends is not only essential for
Full Report
maintaining the inlet and outlet temperatures of syngas by reducing the heat loss to the
atmosphere, but also keeping the maintenance workers safe at a comfortable working
temperature when inspection is being done.
Therefore, the type of insulation recommended for used for this vessel is mineral wool
(mineral fibres with woolly texture) as it is made from made from molten glass, rock or slab.
This is because mineral wool is tough enough to resist wear and tear induced by negligence
and site conditions. Considering the properties of syngas involved in the pressure vessel, the
mineral wool used is inert to the components present in the syngas in case of any minor leak
from the equipment. Since the reactor vessel is located outdoors, the mineral wool able to
withstand adverse weather conditions due to monsoonal changes in Serian Sarawak, hence
avoiding any contamination from any weather conditions. Moreover, mineral wool is able to
absorb noise generated from the reactor, thus making the workplace quieter (Pilkington
Insulation & Willoughby, 2003).
5.2.2.7 Justification on the material and type of support used

The material selected for the construction of support is carbon-manganese steel.
Table 5.2.2.7: The specification of steel for the construction of support

according to ASME BPV Code (Sinnott & Towler, 2009)
Material Grade Temperature Minumum Allowable Tensile

Strength (MPa)
Carbon Steel A285 100-220 101
Gr A
This steel is selected as it is relatively cheaper than other types of steel and is able to provide
enough mechanical strength to elevate and support the total weight of the vessel and contents.
Since the reactor vessel is tall and vertical, supports should be designed to allow easy access
to the vessel and fittings during inspection and maintenance. Hence, straight cylindrical skirt
supports are chosen instead of saddle supports as they are able to distribute the load evenly
around the vessel shell, which prevents a localization of stress experienced by the bracket
supports (Sinnott & Towler, 2009).
5.2.2.8 Justification on the type of painting used

The exterior paint coating of the reactor vessel is done to protect the vessel from atmospheric
corrosion. It is important that the paint can last for a long period of time in order to maximize
the efficiency of the protection and reduce costs for regular maintenance. Hence, the
degradation of paint should be considered during the selection of paint. Degradation of paint
can be due to UV radiation, heat, water and atmospheric pollutants. Thus, the paint selected
should have a high tolerance against these causes of degradation. To mitigate the effect of
UV radiation, additives such as ultraviolet light absorber (UVA) and hindered amine-light
stabilizer (HALS) could be added into the paint to extend its lifespan. Also, a lighter colour
in order to reduce the heat absorbed by the paint. On top of that, the performance of the paint
could be enhanced significantly by using the optimum paint thickness (Nichols, 2012).
Full Report
5.2.2.9 Justification on the nozzles and manhole locations

There are five nozzles in total that are used for the syngas inlet and outlet; cooling water inlet
and outlet and venting. The nozzles for the syngas are located at the centre top and bottom of
the hemispherical ends while the venting nozzle is located near the top nozzle where the
syngas enters. The manhole for the low temperature water gas shift is not required as the top
end of the hemispherical head can be removed during maintenance and also when there is a
need to replace the deactivated catalyst due to the multi-tube design of the reactor. Only the
top hemispherical end is bolted to the vessel allowing the top end of the vessel to be removed.
5.2.2.10 Justification on corrosion allowance

The corrosion allowance is necessary to allow for material lost due to corrosion and erosion
or scaling. The minimum wall thickness calculated in this report using the rules given in the
ASME BPV Code assumes a fully corroded condition (Sinnott & Towler, 2009). As
corrosion is unpredictable and no specific rule can be used to predict the corrosion allowance
needed, engineers normally assume a value for corrosion allowance based on their experience
with the material of construction. For low alloy steel, the corrosion allowance should go by a
‘rule of thumb’ of 2mm as the syngas and cooling water has considerably low corrosion rates.
The minimum corrosion allowance is also required in order to ensure that the thickness is
sufficient for the operating life of the vessel.
5.2.2.11 Basis of design pressure used in mechanical design calculation

The pressure vessel is designed to withstand a pressure of more than 10% of the original
operating pressure. The 10% increment in the normal operating pressure is taken as the
design pressure to allow minor operating upsets that can occur during process operation.
Furthermore, the external pressure or different pressure on opposite sides of the wall that
might cause the vessel to collapse is also accounted for in the design pressure.
5.2.2.12 Basis of design temperature used in mechanical design calculation

The design temperature is 30 more than the normal operating temperature. This is because
the maximum allowable stress of the steel is dependent on the temperature in the vessel.
Similar to design pressure, the design temperature allows minor operating upsets that are
regular during plant operation.
Full Report
5.2.3.1 Mechanical Design Drawing

Full Report
5.2.3.2 Data Sheet for Low Temperature Water-Gas Shift Reactor
5.2.2.2.1 Specification Sheet for Bed Internals, Shell, Insulation and Supports
HIGH TEMPERATURE WATER GAS
SHIFT REACTOR SHEET
Company Alternis BioAmmonia Pvt. Description

Ltd.
Equipment No. R-302 Sheet No.
Function Converts CO to H2
No. Required 1
OPERATING DATA
Reactor Type Isothermal Multi-tubular Reactor
Maximum Diameter Total Height
Process Data Units Value
Operating Temperature 200
Design Temperature 230 (Maximum)
Inlet Outlet
Inlet Pressure kPa (abs) 2379.91 1975.89
Mass Flow of Syngas kg/h 27015.67 27015.67
Mass flow of Cooling kg/h 3076.51 3076.51
Water
Residence time s 11.79323129
Catalyst Loading and Material
Type Copper based catalyst (a mixture of ZnO, CuO and Cr2O3/ Al2O3)
Design Orientation Vertical Pipe Material Low-Alloy A387
Steel Gr 22
Number of catalyst- 69 Nominal Bore 150 Schedule 40
loaded tubes (mm)
Units Value
Length of tube m
Density of pipe material kg/m3 8000
Total mass of tubes kg 16404.62804
3
Bulk density of catalyst kg/m 1442
Voidage in catalyst bed - 0.5
Equivalent sphere m 0.00023
diameter
Mass of catalyst per kg 60.40
tube
Total mass of catalyst kg 4236.79
Mesh size m 0.000177 Tyler 80 Mesh
Equivalent
Shell Construction and Materials
Shell Material Low-Alloy Steel A387 Gr22
Shell Inner Diameter 2.172 m
Shell Length 2.265 m
Full Report
Shell Thickness 0.042 m

Pressure drop across 0.654 Pa
Shell Side
Baffle Construction and Materials
Baffle Material Low Alloy Steel A387 Gr22
Density of Stainless 8000 kg/m3
Steel
Baffle Cut 25 %
Thickness of Baffle 0.015 m
Total Mass of Baffle 1003.19 kg
(including liquid on it)
Number of Baffles used 4 -
Tube Plate Construction and Material
Tube Plate Material Low Alloy Steel A387 Gr22
Density of Stainless 8000 kg/m3
Steel
Thickness of Tube Plate 0.010 m
Mass of Tube Plate 587.17 kg
(including liquid/ gas on
it with tube holes
accounted for)
Number of plates used 2 -
Insulation Material
Thickness of Insulator 0.2 m
Total Volume of 1.377 m3
Insulator for
Hemispherical Heads
Density of Mineral 130 kg/m3
Wool
Mass of Insulator 179.02 kg
Total Weight of 1756.17 N
insulator with fitting
allowance
Skirt Construction and Material
Skirt Material Carbon-Manganese A 285 Gr A
Type of Skirt Straight
Cylindrical
Thickness of Skirt 0.042 m
including corrosion
allowance of 2mm
Skirt Height 2.825 m
, bending stress in the 2.04 107 N/m2
skirt
, dead weight stress 2.00 106 N/m2
on skirt
, tensile stress 1.83 107 N/m2
, tensile stress limit 1.01 108 N/m2
, compressive stress 2.24 107
N/m2
, compressive stress 4.42 108 N/m2
limit
Nozzle Construction and Material based on Pipe Selection
Syngas (nozzle located at reactor heads top and bottom)
Full Report
Pipe Selection Carbon Steel A285 Nominal Size 250 Schedule 40

Gr A (mm)
Syngas Inlet Nozzle 0.303 m
Inner Diameter
Required Inlet velocity 9.4-18 m/s
Syngas Outlet Nozzle 0.303 m
Inner Diameter
Required Outlet 9.4-18 m/s
velocity
Outer Diameter of 0.324 m
Nozzle, O
Thickness of Flange, tf 0.046 m
Thickness of Lap Joint 0.048 m
Diameter of Hub,X 0.321 m
Diameter beginning of 0.273 m
chamber, A
Length through Hub, Y 0.111 m
Bore, B 0.277 m
Cooling Water (nozzle located at reactor sides)
Pipe Selection Carbon Steel A285 Nominal Size 20 Schedule XS
Gr A (mm)
Water Inlet Nozzle 0.021 m
Inner Diameter
Required Water Inlet 2.4-3 m/s
velocity
Water Outlet Nozzle 0.021 m
Inner Diameter
Required Water Outlet 2.4-3 m/s
velocity
Outer Diameter of 0.027 m
Nozzle, O
Thickness of Flange, tf 0.014 m
Thickness of Lap Joint 0.016 m
Diameter of Hub, X 0.048 m
Diameter beginning of 0.027 m
chamber, A
Length through Hub, Y 0.056 m
Bore, B 0.028 m
Remarks:
By Lydia Yap Li-Ya Date 14/1/2014
Checked by Lee Leong Hwee Date 16/1/2014
Full Report
5.2.3.2.2 Specification sheet for analysis of stresses

HIGH TEMPERATURE WATER GAS
SHIFT REACTOR SHEET
Company Alternis BioAmmonia Pvt. Description

Ltd.
Equipment No. R-302 Sheet No.
Function Converts CO to H2
No. Required 1
Design Conditions
Design Pressure 2618 kPa
Design Temperature 230
Maximum Allowable Stress 114.45 MPa
Estimation of Thickness
Thickness of column vessel
, minimum thickness of column vessel required 0.0251 m
, estimated thickness of column vessel required 0.040 m
, allowable corrosion thickness 0.002 m
, total estimated column thickness with 0.042 m
corrosion allowed
Thickness of heads and closure (hemispherical head)
, minimum thickness of vessel ends (top and bottom) 0.0125 m
required
, estimated thickness of vessel ends (top and 0.040 m
bottom) required
, total estimated thickness of vessel ends 0.042 m
with corrosion allowed
Dead Weight Loadings
, weight of vessel 106952.76 N
, weight of baffles and tube plates 15601.45 N
, weight of tubes and catalyst 86109.99 N
, total weight of insulator including allowances 3500.34 N
for fittings
, weight of spherical ends (top and bottom) 216630.54 N
, maximum weight of water in the vessel 134991.98 N
under normal operating condition
, total dead weight loads during normal 563787.06 N
operating condition
Wind Loadings
, bending moment due to wind load 1.34 Nm
, bending moment due to dead load under normal 3.22 Nm
operating condition
, total bending moment due to dead load under 3.23 Nm
normal operating condition
Analysis of Stresses (Sinnott Chemical Engineering Design)
Normal operating condition:
Pressure Stresses
, longitudinal stress 33850.40 kPa
Full Report
, hoop stress 67700.81 kPa

Dead Weight Stress
, dead weight stress (compression) 1929.65 kPa
Bending Stress
, bending stress 20319.18 kPa
Resultant longitudinal stress
, the resultant longitudinal stress up-wind 56099.23 kPa
, the resultant longitudinal stress down-wind 15460.88 kPa
, greatest distance between the principle stresses 52.24 MPa
Elastic Stability
, critical buckling stress 354.56 MPa
Resultant compressive stress
, resultant compressive stress 22.25 MPa
Remarks:
By Lydia Yap Li-Ya Date 14//1/2014
Checked by Lee Leong Hwee Date 16/1/2014
Calculation for the catalysts weight, number of tubes, vessel thickness, vessel height and stress
analysis can be found in Appendix C
Full Report
5.3 Detailed Process and Mechanical Design: Carbon Dioxide Absorption

Column

This section will be describing the design of the amine absorption column. Amine absorption
column is used for removing CO2 from the syngas entering the absorber, which forms a part
of the CO2 removal section. The syngas inlet stream information is shown in the table below.
Table 5.3.1.1: Syngas Feed Stream Information

Syngas Inlet
Temperature (ºC) 50
Pressure (kPa) 2370
Molar Flow Rate (kmol/hr) 997.90
Mass Flow Rate (kg/hr) 19239.28
CO2 0.3040
N2 0.1671
Composition H2 0.5202
(Mole Fraction) H2O 0.0052
CO 0.0017
CH4 0.0018
CO2 removal and capture process is an important step in many processes in most of the
industrial process plants. In ammonia production plants, CO2 is being removed from the
process stream as it is an undesirable component in the syngas. Presence of CO 2 tends to
cause temperature excursions in the process. Moreover, CO2 may poison the iron catalysts
present in the ammonia synthesis reaction in downstream process.
Normally, CO2 removing process takes place in two different operating units, namely
absorption column and stripping column. Absorption column is used for removal of CO2 from
the gas stream whereas stripping column is responsible for solvents regeneration. However,
this section is mainly focused on the absorption column. The sour gas (syngas) stream will
enter the absorber column from the bottom and contact with the solvent stream that flows
counter-currently from the top of the column. CO2 in the sour gas stream will be removed and
absorbed by the solvent and leaves the column at the bottom. The gas stream that has been
Full Report
purified will leave the absorber at the top and continue to the next process. There are some
design objectives that have to be met in designing the CO2 removal operation unit:
1. Removal of CO2 to concentration of less than 100 ppm in treated syngas (sweet gas)
stream.
2. CO2 stream with minimum CO2 purity of 95% before sending to glycol drying plant
Table below summarized the stream information of the treated syngas outlet stream.
Table 5.3.1.2: Treated Syngas Outlet Stream Information

Treated Syngas Outlet
Temperature (ºC) 50.06
Pressure (kPa) 2350
Molar Flow Rate (kmol/hr) 694.57
Mass Flow Rate (kg/hr) 5889.83
CO2 0.0001
N2 0.2400
Composition H2 0.7474
(Mole Fraction) H2O 0.0075
CO 0.0025
CH4 0.0025
There are several factors that should be considered for operational safety purposes in
designing the amine absorption column:
1. High operating pressure of the absorption column

2. Exothermic reaction between the acid gas and the amine solvent solution
3. Forming of corrosive environment in the column
Full Report
These factors are also taken into consideration when selecting the materials of construction.
Selection of construction material is one of the important steps in design and the most
suitable options should be chosen. In this design, carbon steel Grade A285 is chosen as the
construction material for the absorption column. Carbon steel is widely used as construction
material for amine absorption unit in industries (Kohl & Nielsen, 1997). It is inexpensive and
it provides good resistance to corrosion, good strength and workability, ease of fabrication, as
well as good weld-ability (Gandy, 2007). Besides that, carbon steel materials could cover
extensive mechanical properties for column design. Although stainless steel has higher
maximum allowable stress as compared to carbon steel, the maximum allowable stress of
carbon steel is sufficient for this absorption column. Furthermore, one major drawback of
stainless steel is the high material cost. Stainless steel is much more expensive as compared
to carbon steel, and this will increase the capital cost of the plant which is not economical
viable.
In the carbon dioxide (CO2) removal process, the activated MDEA solvent used is slightly
corrosive due to the low corrosive nature of piperazine. Moreover, CO2 tends to form
corrosive environment when dissolved in water. Thus, as a safety precaution and for
operational safety purposes, corrosion allowance of 4mm is added to the absorption column.
The detailed design of absorption column was carried out in accordance to the American
Society of Mechanical Engineers (ASME) standard. The procedures for detailed design of the
absorption column and support based on ASME standard were taken from Sinnott & Towler
(2009). The set of codes covered by ASME standard are listed as follows:
 Minimum thickness of the vessel

 Type of head and end and the minimum thickness
 Maximum allowable stress of the material at given temperature
 Corrosion allowance
 Internal and external stresses
 Vessel support
Full Report

The amine solvent used in the CO2 removal process is piperazine-activated
methyldiethanolamine (aMDEA). The amine solvent solution consists of 40% MDEA, 5%
piperazine and 55% water, in terms of weight percentage. MDEA is a tertiary amine and it is
commonly used in CO2 removal processes in most of the industrial plants. On the other hand,
piperazine is a cyclic diamine and it is the most commonly used promoter in industry
processes that involve CO2 removal. Piperazine acts as an activator in the amine solvent
solution which increases the reaction rate of the solvent with CO2 (Alvis, et al., 2012).
Activated MDEA has high affinity towards CO2 and it has high efficiency in absorbing and
capturing CO2 (Kunjunny, et al., 1999).
Moving on, the type of column used for the absorber is the packed column. Since the flow
rate of amine solvent used for the CO2 removal process is high, a packed column is suitable
to be used as it is effective in handling large liquid rate. Packed column would have shorter
tower height as compared to tray column, and it is mechanically simple (Pilling & Holden,
2009). On top of that, the gas-liquid contact in a packed column is continuous, where the
liquid flows down the column over the packed bed and the vapour flows up the column
counter-currently (Sinnott & Towler, 2009). This would increase the contact area and contact
time between the liquid and vapour, and hence increase the efficiency of the process. Packed
column is also more economically beneficial for handling corrosive system (Sinnott &
Towler, 2009). The amine solvent used in the system is corrosive, and the corrosive behavior
of dissolved CO2, thus packed column is suitable to be used. Furthermore, packed column
could be operated at lower pressure drop as compared to tray column (Pilling & Holden,
2009).
For the packing material used in the packed bed of the absorption column, INTALOX saddle
ceramics, random packing, are chosen. Random packing is chosen over structured packing
for the absorption column in this project due to several advantages of random packing. Firstly,
cost of random packing is significantly lower than the cost of structured. This is economically
beneficial as the capital cost could be reduced. Next, the packings are placed in the packing
bed randomly without specific arrangement. Random arrangement of the packings is able to
improve the liquid distribution, which will results in more contact opportunities between the
liquid and the vaour that flows counter-currently and thus higher process efficiency (Sinnott
& Towler, 2009). Ceramic material is chosen because it is more suitable to be used to handle
Full Report
the corrosive environment in the absorption column. INTALOX saddle ceramic is shown in
the figure below.
Figure 5.3.2.1: INTALOX Saddle Ceramic (Pilling & Holden, 2009)
5.3.3.1 Design Pressure and Temperature

The column must be designed in such a way that it is able to withstand the maximum
pressure that will be exerted on it during the operation. For a column under internal pressure,
the design pressure is normally taken at 5% to 10% above the normal operating pressure for
safety operation purposes (Sinnott & Towler, 2009). In this design, the operating pressure for
the absorber is at 23.70bar. The design pressure is taken 10% above the operating pressure,
which is found to be 26.07bar.
Moving on, strength of metals decrease with increasing temperature, which also indicates
that maximum allowable stress is dependent on the temperature (Sinnott & Towler, 2009).
Under ASME BPV Code, the maximum design temperature corresponds to the evaluated
maximum allowable stress should be taken at the maximum operating temperature. For this
design, the operating temperature of the absorber in this design is 50ºC and the design
temperature is taken as 70ºC.
Full Report
5.3.3.2 Materials of Construction

When selecting the construction material for the absorption column, several factors are
considered, which include the operating conditions, the process and the degree of corrosion.
In this design, carbon steel Grade A285 is chosen as the construction material for the
absorption column. Carbon steel is widely used as construction material for amine absorption
unit in most of the industrial plants (Kohl & Nielsen, 1997). The major advantage of carbon
steel is that it is inexpensive and it provides good resistance to corrosion. In addition to that,
one of the advantageous of carbon steel is the ease of fabrication and good weld-ability
(Gandy, 2007). Carbon steel can be welded, machined and fabricated easily. On top of that,
carbon steel materials could cover extensive mechanical properties for column design.
Carbon steel provides good strength and has high toughness (Gandy, 2007). The maximum
allowable stress of carbon steel is sufficient for the absorption column in this design,
although the maximum allowable stress of stainless steel is higher as compared to that of
carbon steel. The major disadvantage of stainless steel is the cost of the material is high.
Stainless steel is much more expensive than carbon steel, which will results in high capital
cost of the plant and it is not economical viable.
As for the skirt support of the column, carbon steel is chosen as the construction material.
The skirt support is in contact with any corrosive material present in the process stream. Thus,
carbon steel is suitable to be used as the construction material for skirt support.
5.3.3.3 Column Head and Closure

Both ends of the column are closed by heads that have different shapes. The most common
types of head used are listed below (Sinnott & Towler, 2009):
 Flat plates and formed flat heads

 Hemispherical heads
 Ellipsoidal heads
 Torispherical heads
Hemispherical, ellipsoidal and torispherical heads are also referred to as domed heads. They
are commonly used heads for vessels operate at high temperature. Torispherical heads are
suitable to be used for vessels with operating pressure up to 15bar, ellipsoidal heads are
usually proved to be the most suitable heads to be used for vessels with operating pressure of
Full Report
above 15bar, and hemispherical heads are the strongest shape and it is able to resist about
twice the pressure of the torispherical heads (Sinnott & Towler, 2009). However, although
hemispherical heads have the strongest shape, the cost to form the heads is very high. Thus,
ellipsoidal heads are deemed to be the most suitable heads to be used for the absorption
column in this project as the operating pressure of the column is 23.70bar.
5.3.3.4 Column Internals

(i) Mist Eliminator
In any process that involves liquid and gas that come into contact, liquid droplets tend to
entrain in the processing gas. This will cause inefficiency of the process, contamination of
the gas product and damage to the equipments. Thus mist eliminator is installed to
improve the product purity and to prevent the entrainment of the liquid droplets. The use
of a mist eliminator in the amine absorption column minimizes the entrainment of amine
solvent in the treated syngas. This helps in minimizing the contamination of the treated
syngas by the amine solvent, as well as helps in recovering the amine solvent thus
reducing the makeup rate and cost of fresh amine solvent.
In this project, DEMISTER mist eliminator, knitted wire mesh pad type, by Koch-Glitsch
is selected. DEMISTER mist eliminator is easy to install in all process equipment and it
provides high separation efficiency with very low pressure drop (Koch-Glitsch, 2012).
Stainless steel is chosen as the construction material to provide corrosion resistance
against the corrosive environment in the absorption column.
(ii) Liquid Feed Devices
Liquid feed pipe is used to channel the amine solvent solution into the center of the
liquid distributor. Model 119 INTALOX High Performance Liquid Only Feed Pipe by
Koch-Glitsch is chosen to be used in the absorption column. The advantage of this liquid
feed pipe is that the excessive turbulence and horizontal flow velocity in the distributor
can be eliminated (Koch-Glitsch, 2010).
Full Report
(Iii) Liquid Distributor
Liquid distributor is normally used in packed column above each packed bed. Liquid
distributor is important as it helps in ensuring a uniform liquid distribution. A liquid
distributor is installed in the amine absorption column so that the amine solvent solution
is distributed uniformly over the packing. Model 136 INTALOX Channel Distributor
with Drip Tubes by Koch-Glitsch is selected to be used in the absorption column. This
model of distributor is efficient to be used in column with diameter greater than 250mm
(Koch-Glitsch, 2010). This distributor has higher fouling resistance as compare to other
type of distributor. In addition, the center channel provides good structural support and
equalization of liquid between troughs. The position of the orifices in the sidewalls
provides optimum distribution quality, while vapour passage can be found between the
troughs. Besides that, Liquid distributor of this model is flexible as the drip tubes are
removable and replaceable (Koch-Glitsch, 2010).
Liquid redistributors are only needed for column with packed bed height exceeds 8-10
times of the column diameter (Sinnott & Towler, 2009). Since the ratio of packed bed
height to column diameter for the absorption column in this project is 4.3, redistributors
are not necessary.
(iv) Bed Limiter
Bed limiter is included in the packed column with random packing to confine the
upward movement of the packing and to prevent the fluidization of packing at the top of
the packed bed. For random packing, there is always a potential for sufficient vapour
load to cause fluidization of packing at the top of the packed bed. Since fluidization of
packing is difficult to predict, a bed limiter is always recommended for packed column
that uses random packing (Koch-Glitsch, 2010). In this project, Model 805 Random
Packing Bed Limiter, Non-Interfering by Koch-Glitsch is selected.
Full Report
(v) Support Plates
Support plate is needed for every packed bed in the column. Support plate must be able
to support and retain the packed bed in the column under operating conditions (Koch-
Glitsch, 2010). Packed column with random packing normally uses gas-injection type
support that provides different passages for liquid and vapour flow. The Model 804
Random Packing Gas Injection Support Plate is chosen in this project.
5.3.3.5 Column Externals

(i) Manways
The absorption column consists of 3 manways, including 1 manway for loading and 2
manways for cleaning and maintenance. Manways for cleaning and maintenance are
located at the top and bottom of the packed bed while manway for loading is located
beside the packed bed. Manways must be large enough for the access of operators without
much difficulty. Sinnott & Towler (2009) stated that the typical diameter of a manways is
0.6m.
(ii) Ladders
Plain ladder is installed to ease the operators to access the manways for maintenance or
cleaning purposes.
(iii) Platforms
Total of 3 platforms are installed on the absorption column. Based on ASME standard,
platform is typically located at 700-900mm below the nozzle. Platforms are installed to
provide a space for the operators to access the manways for maintenance and cleaning
purposes.
Full Report
5.3.3.6 Height of Column

The total height of the absorption column is determined based on the height of the ellipsoidal
heads, height of packed bed, and the height of all the internal fittings installed in the column.
Height of the packed bed was calculated based on the graph plotted by using the equilibrium
data published by Haji-Sulaiman, et al. (2008), as well as the formulas and methods given in
Froust, et al. (2008) and Sinnot & Towler (2009). As for the internal fittings, the height is
estimated based on the data given in Koch-Glitsch (2010). Extra spacing of 0.5m is allowed
between two internal fittings. Spacing of 1m is added between liquid distributor and bed
limiter as the liquid distributor is normally positioned 100-200mm above the bed limiter
(Koch-Glitsch, 2010). The height of different components that make up the total height of the
column is summarized in Table 5.3.3.6.1. Detailed calculations are shown in Appendix
Section C3.1.2.
Table 5.3.3.6.1: Summary of Height of the Internal Fittings of the Column

Component Height (m)
Ellipsoidal Dome (Top) 0.3110
Mist Eliminator 0.1500
Nozzle (Solvent Inlet) 0.1143
Manway 1.0000
Liquid Distributor 0.4000
Bed Limiter 0.1500
Packed Bed 8.2684
Support Plate 0.3000
Manway 1.0000
Nozzle (Syngas Inlet) 0.2731
Ellipsoidal Dome (Bottom) 0.3110
Spacing 0.4000
TOTAL Height of the Column 12.6778
Full Report
5.3.3.7 Thickness of the Column

(i) Shell Thickness
The minimum thickness for a cylindrical vessel that is required to resist the internal
pressure was calculated by using Equation (13.40) in Sinnott & Towler (2009).
The minimum thickness required was calculated to be 18.28mm initially. However, the
vessel with such thickness did not pass the analysis of stress under hydraulic testing
condition. Thus, optimization was carried out by performing iteration on the shell
thickness in Microsoft Excel until the minimum thickness required that is able to
withstand the high pressure in hydraulic testing condition is achieved. Thus, the
optimized minimum thickness required was determined to be 28mm.
Corrosion allowance of 4mm is added to the absorption column due to the fact that the
activated MDEA solvent used is slightly corrosive due to the low corrosive nature of
piperazine. Moreover, CO2 tends to form corrosive environment when dissolved in water.
According to Sinnott & Towler (2009), corrosion allowance of 4mm should be added to
the vessel thickness as a safety precaution. Thus, the thickness of the cylindrical vessel
was calculated to be 32mm with inclusive of corrosion allowance. Detailed calculations
are shown in Appendix Section C3.1.5.
(ii) Head and End Thickness
For ellipsoidal heads, the minimum thickness required was calculated by using Equation
(13.45) in Sinnott & Towler (2009), in accordance to ASME BPV Code Sec. VIII D.1
Part UG-32.
Full Report
The minimum thickness required was calculated to be 18.28mm. However, in order to

have an absorption column with uniform thickness and for operational safety purposes,
the thickness was taken to be the same as the thickness required for the cylindrical
column, which is 28mm. Similarly, corrosion allowance of 4mm was considered and the
thickness was found to be 32mm. Detailed calculations are shown in Appendix Section
C3.1.5.
5.3.3.8 Dead Weight of the Column

The dead weight of the column is evaluated at 2 different conditions, which are under normal
operation condition and under hydraulic testing condition.
(i) Normal Operation
Under normal operation, weight load of column, weight of packed-bed, weight of column
external fittings and weight of column internal fittings are taken in account in evaluating
the dead weight of the column.
Weight load of the column was determined by using Equation (13.73) in Sinnott &
Towler (2009).
Weight of the packed bed was calculated by using the equation shown below:
Weight of column external fittings, which include plain ladders and platforms, was
calculated based on the guide provided in Sinnott & Towler (2009). The guide provided
to calculate the weight of column external fittings is summarized below:
Weight of the column internal fittings is estimated to be roughly 20% of the summation
of weight of vessel and weight of packed-bed.
Full Report
Total dead weight of the column under normal operation is summarized and tabulated in
Table 5.3.3.8.1. Detailed calculations are shown in Appendix Section C3.2.1.
Table 5.3.3.8.1: Total Dead Weight of the Column Under Normal Operation
Weight load of Column (kN) 147.9676
Weight of Packed Bed (kN) 60.0259
Weight of Ladders (kN) 1.8084
Weight of Platforms (kN) 14.6094
Weight of Column Internal Fittigns (kN) 41.5897
TOTAL (kN) 266.0099
(ii) Hydraulic Testing Condition
Under hydraulic testing condition, the column is fully filled with water to simulate the
worst case scenario. Thus, weight of water is taken into account when calculating the
total dead weight of the column.
Weight of water was calculated based on the formula given in Sinnott & Towler (2009).
Total dead weight of the column under hydraulic testing operation is summarized and
tabulated in Table 5.3.3.8.2. Detailed calculations are shown in Appendix Section C3.2.2.
Table 5.3.3.8.2: Total Dead Weight of the Column under Hydraulic Testing
Condition
Dead Weight of Column (kN) 266.0099
Weight of Water (kN) 153.5978
TOTAL (kN) 419.6077
Full Report
5.3.3.9 Analysis of Stress

Stress analysis is the most important is a project as it is carried out to ensure that the designs
are able to withstand the worst combination of loadings without failure (Sinnott & Towler,
2009). Several stress analysis were carried out in accordance to ASME BPV Code Sec. VIII
D.1 Part UG-23, which include primary stresses due to internal pressure, dead weight stress
due to weight of the column, as well as bending analysis subjected to wind loading. Analyses
of stresses were conducted for both normal operation condition and hydraulic testing
condition.
At bottom tangent line, the longtitudinal stress (σL) and hoop stress (σh) was calculated by
using Equation 13.62 and Equation 13.61 in Sinnott & Towler (2009), respectively.
The dead weight stress (σw) was calculated by using Equation 13.63 in Sinnott & Towler
(2009).
The Bending moment (Mx) at bottom tangent line was calculated with Equation 13.75 in
Sinnot & Towler (2009).
For preliminary design studies, the wind speed can be taken as 160 km/hr, which is
equivalent to a wind pressure of 1280 M/m2 (Sinnott & Towler, 2009). Then, the bending
stress was calculated by using Equation 13.64 in Sinnott & Towler (2009).
Full Report
The resultant longitudinal stress (σz) was also calculated based in the formulas given in
Sinnott & Towler.
Detailed calculations are shown in Appendix Section C3.2.3 to C3.2.7.
The results of analysis of stresses are summarized and tabulated in Table 5.3.3.9.1.
Table 5.3.3.9.1: Results of Analysis of Stresses

Stresses Operating Condition
Normal Hydraulic Testing
Longitudinal Stress, σL (MPa) 28.9530 43.0158
Hoop Stress, σh (MPa) 57.9059 86.0316
Dead Weight Stress, σw (MPa) 2.3777 3.7506
Bending Stress, σb (MPa) 10.9751 10.9751
Resultant Longitudinal Stress 37.5503 50.2403

(Upwind) (MPa)
Resultant Longitudinal Stress, σz 15.6002 28.2902

(Downwind) (MPa)
Difference in Principle Stresses (MPa) 42.3057 57.7414
Maximum Compressive Stress, σc 13.3527 14.7256

(MPa)
Full Report
The maximum allowable stress of carbon steel Grade A285 is 88.9429 MPa, and the critical
buckling stress was calculated to be 489.3507 MPa. As shown in the table above, the
difference in principle stresses for both normal operating condition and hydraulic testing
condition is lower than the maximum allowable stress of carbon steel. In addition, the hoop
stress calculated for both the operating conditions was found to be lower than the maximum
allowable stress of carbon steel. The maximum compressive stress for both normal operating
condition and hydraulic testing condition was found to be well below the critical buckling
stress. Therefore, this can be concluded that the design is satisfactory.
5.3.3.10 Column Support

The type of support used to support will be depending on various factors, including the size,
shape and weight of the column, the design pressure and temperature, as well as the column
location and arrangement (Sinnott & Towler, 2009). The support must be able to withstand
the weight of the column and contents, and any other superimposed loads (Sinnott & Towler,
2009). As stated in Sinnott & Towler (2009), skirt support is normally used for tall, vertical
column that are subjected to wind loading. Thus, skirt support selected for the absorption
column is the straight cylindrical shell type.
The skirt support is welded to the end of the column, as shown in Figure 5.3.3.10.1. The skirt
thickness must be adequate to withstand the weight of the column and bending moment
applied on the column.
Figure 5.3.3.10.1: Skirt-Support Welds (Sinnott & Towler, 2009)

Full Report
The bending stress was calculated by using Equation (13.84) in Sinnott & Towler (2009).
The dead weight stress was calculated by using Equation (13.85) in Sinnott & Towler (2009).
The resultant stresses of the skirt must not exceed the required design criteria under worst
combination of dead-weight loading of the column and wind loading (Sinnott & Towler,
2009).
The required criteria can be checked by using Equation (13.86) and (13.87) in Sinnott &
Towler (2009).
The detailed calculations are shown in Appendix Section C. The results calculated are
summarized and tabulated in Table 5.3.3.10.1.
Table 5.3.3.10.1: Results of Skirt Design

Skirt Support with Straight Cylindrical Shell
Bending Stress, σbs (MPa) 5.0182
Dead Weight Stress, σws 2.2018
(Normal Operating Condition) (MPa)
Dead Weight Stress, σws 3.4732
(Hydraulic Testing Condition) (MPa)
Resultant Stress (Tensile), σs (MPa) 2.8163
Resultant Stress (Compressive), σs (MPa) 8.4913
Full Report
The resultant tensile stress was found to be well lesser than the design criteria, 128.2097 MPa.
In addition, the resultant compressive stress calculated was also much lower than the design
criteria, 535.6029 MPa. Therefore, the skirt support with straight cylindrical shell design is
satisfactory.
5.3.3.11 Nozzle Sizing

The nozzles are sized based on the guides provided in Sinnott & Towler (2009). The area of
the pipe was first calculated by using the formula as shown below:
Then the optimum velocity of the fluid was estimated from the Table 5.3.3.11.1 by
performing interpolation.
Table 5.3.3.11.1: Optimum Velocity in terms of The Fluid Density (Sinnott &
Towler, 2009)
Fluid Density (kg/m3) Velocity (m/s)
1600 2.4
800 3.0
160 4.9
16 9.4
0.16 18.0
0.016 34.0
Then the area and diameter required were calculated. Based on the calculated required
diameter, the nominal pipe size was chosen from ASME pipe schedule list provided by
ArcelorMittal.
Full Report
The results were summarized in Table 5.3.3.11.2 as shown below.
Table 5.3.3.11.2: Summary of Pipe Sizing Results
Actual Pipe
Vs Ap Dp
Stream
Nominal Schedule
(m/s) (m2) (mm)
Diameter (mm) Number
9.3684 0.03353 206.6246

Syngas Inlet 250 30
Treated 14.0709 0.01572 234.9913 200 30

Syngas
2.8855 0.0078 99.5944

Solvent Inlet 125 40
Rich Amine 2.8834 0.0092 107.7539 150 40

Outlet
Full Report
5.3.4 Mechanical Drawing and Data Sheet
5.3.4.1 Mechanical Design Drawing

Full Report
5.3.4.2 Mechanical Data Sheet
CARBON DIOXIDE ABSORPTION COLUMN

Company Alternis BioAmmonia Pvt. Ltd. Description CO2 Removal

Equipment No. AC-401 Sheet No. 1
Removes carbon dioxide from
Number Required = 1 Functions
sour gas (syngas)
Fluid Properties
IN OUT
Stream Sour Gas Sweet Gas
Lean Amine Rich Amine
Description (Syngas) (Treated Syngas)
Flow Rate kg/hr 19239.28 77094.01 5889.8254 90443.47
Temperature ºC 50.00 50.00 50.06 50.30
Pressure bar 23.79 23.70 23.50 23.50
Technical Data
Operating Temperature ºC 50 Design Temperature ºC 70
Operating Pressure bar 23.70 Design Pressure bar 26.07
Design Orientation - Vertical Design Code - ASME
Column Type - Packed Column Packing Type - Random
Reaction Type - Absorption Packing Model - INTALOX saddle
Column Diameter m 1.2439 Packing Height m 8.2684
Column Height m 12.6778 Insulation Material - No
Wall Thickness m 0.032 Domed Head Type - Ellipsoidal
Corrosion Allowance m 0.004 Weight (Normal) kN 266.0099
Weight
Construction Material - Carbon Steel (A285) kN 419.6078
(Hydraulic Testing)
Mechanical Design Data

Normal Operation Hydraulic Testing Operation
Longitudinal Stress MPa 28.9530 43.0158
Hoop Stress MPa 57.9059 86.0316
Dead Weight Stress MPa 2.3777 3.7506
Bending Stress MPa 10.9751 10.9751
Resultant Stress (Upwind) MPa 37.5503 50.2403
Resultant Stress (Downwind) MPa 15.6002 28.2902
Max. Compressive Stress MPa 13.3527 14.7256
Vessel Support Data
Support Type - Skirt Support Length m 0.3
Construction Material - Carbon Steel (A285) Thickness m 0.03
Connection - Welded at Joint Height m 0.2
Nozzle Specification Material Specification
Nozzle Description Nominal Size Parts Material
A1 Syngas Inlet 250 Column Carbon Steel A285
A2 Lean Amine Inlet 125 Packing Ceramic
B2 Treated Syngas Outlet 200 Column Support Carbon Steel A285
B1 Rich Amine Outlet 150 Nozzles Carbon Steel A285
- Manways 600 Manways Carbon Steel A285
Remarks
Prepared By Lee Leong Hwee Date 28th January 2014
Checked By Lydia Yap Li-Ya Date 28th January 2014
Full Report
5.4 Detailed Process and Mechanical Design: Methanator

The main purpose of the methanator is to remove the oxides in the process as oxides would
decrease the activity of ammonia synthesis catalyst, the oxides will react with the catalyzed
used in the loop (catalyst poisoning) and cause deposition of ammonium carbonate in the
synthesis loop.
This section focuses on the design of the fixed bed reactor for methanation process that will
be mainly used to remove all carbon oxides present in the syngas to avoid any deposition of
ammonium carbonate in the ammonia synthesis reactor. In methantion process, hydrogen
reacts with carbon monoxide and carbon dioxide to produce methane and steam.
The stream that is to be processed is the heated syngas outlet from the absorber of the CO2
removal section. The compositions, temperature, pressure and flow rates are shown in the
table below. One of the key requirements in this design is the target of the methanation unit to
reduce the outlet of carbon oxides to at least 0.1-0.5% ppm level (Hawkins, 2009).
The primary safety consideration would be over-pressurization occurring within the vessel
due to high operating pressure conditions, exothermic reactions of carbon oxides with
hydrogen and possibility of embritlement due to the presence of hydrogen. All these factors
were taken into consideration during the mechanical design of the vessel. Suitable material
was selected for the vessel and stress analysis was done to ensure that the stresses calculated
were well under the maximum design stress of the material selected.
Full Report
Table 5.4.1-1: Inlet and outlet flow rates, composition and conditions of syngas
Description Inlet Outlet
Vapour fraction 1 1
Temperature ( 300 347.01
Pressure (kPa) 2300 2300
Molar flow
694.57 690.997
(kmol/hr)
Mass flow (kg/hr) 5889.84 5889.84
mass flow Mass Mass
Component mass flow (kg/hr)
(kg/hr) composition composition
H2 1046.46 0.1777 1035.52 0.1758
CH4 28.13 0.0048 56.77 0.0096
CO 48.06 0.0082 Trace Trace
CO2 3.06 0.0005 Trace Trace
N2 4670.46 0.7930 4670.46 0.7930
H2O 93.69 0.0159 127.10 0.0216
5.4.2.1 Catalyst Selection

Nickel based catalyst was selected for the methanation reactions as Nickel is among the most
active metals for the process, suitable for high temperature process and it is relatively cheap.
Although Ruthenium is more active compared to Nickel, it is much more expensive than
Nickel and it is also short in supply. Another disadvantage of Ruthenium is that they produce
higher hydrocarbons under methanation conditions. Moreover, since the methanation
reactions are thermodynamically stable, it can be thought as non-selective; hence, there is no
advantage to be gained by using more expensive metals. (Ross,2013)
5.4.2.2 Catalyst Details

The catalyst selected is the Haldor Topsøe PK-7R methanation catalyst, which is a Nickel
catalyst with MgAl2O3 support (HaldorTopsøe, 2013). The operating temperature range of the
catalyst is 190-450 , with bulk density of 512kg/m3, porosity of 0.3, density of 976kg/m3
and void factor of 0.475 (Zhang, et al., 2006).
Full Report
5.4.2.3 Sizing of Reactor

A vertical vessel is to be designed to process the incoming syngas from the CO2 removal
absorber. The vessel will have a gas inlet and outlet, loading and cleaning manways.
Internally, there will be a catalyst support beam, support plate, wire mesh and a gas
distributor.
Several assumptions were made when sizing the reactor. This includes:
 The reactions involved in the reactor are:
Reaction1
Reaction2
Reaction3
 CO will undergo reaction first before CO2 in the reactor

 CO will undergo reaction 1 and 2, while CO2 will undergo reaction 2 and 3
 The selectivity of CO is 62.2%, while the selectivity of CO2 is 99.4% for Ni/Al2O3
catalyst (Fujita & Takezawa, 1997)
 The differential equation to determine the catalyst weight is:
and
Where and are the rate of reaction of carbon oxide and carbon dioxide, while and
are the inlet flow rates of carbon monoxide and carbon dioxide respectively.
 The differential equation to determine the pressure drop is:
To begin with the design of the reactor, the amount of catalyst required for the chemical
reactions was determined followed by the calculations on the height of catalyst bed required.
The amount of catalyst needed was found by plotting a graph of conversion vs. mass of
catalyst using Polymath Fogler Softwarre as shown in Figure C4.1.1 and C4.1-3 (Appendix).
The amount needed depends on the amount of incoming carbon oxides into the reactor. In
Full Report
this design, an extra 10% of catalyst was added to ensure that the targeted carbon oxides
outlet is achieved and also to take into account catalyst degradation. . This brings the total
catalyst bed mass to 1304 kg. The pressure drop of the reactor was determined by plotting a
graph of weight of catalyst vs. pressure using Polymath Fogler Softwarre as shown in Figure
C4.1-2 and C4.1.4 (Appendix) and it was found to be 25kPa. Next, the diameter of the vessel
was set to be 1m. The catalyst bed height was then calculated to be 3.24m. The reactor was
sized based on the standard of ASME BPV Code (Sec. VIII D.1 Part UG-27). Detailed
calculations are shown in Appendix C4.
Table 5.4.2.3-1: Summary of reactor size

Parameter Unit Value
Mass of catalyst kg 1304
Volume of catalyst m3 2.54
Height of catalyst bed m 3.24
Reactor diameter m 1
Pressure drop kPa 25
Extra height allowance (top & bottom) m 1.2
Total reactor height including ellipsoidal head m 4.44

This section focuses on the mechanical design of the fixed bed reactor for methanation.
Detailed description on the design pressure and temperature, material selection, reactor feed
positioning, support and vapor distributor, vessel end, shell thickness, vessel height,
insulation, external painting, stress analysis, skirt support and pipe sizing will be included in
this section. All the detailed calculations are shown in Appendix C4.2 to C4.11.
5.4.3.1 Design Pressure and Temperature

A vessel must be designed to withstand the maximum pressure to which it is likely to be
subjected in operation. The design pressure is the pressure at which the relief device is set.
The normal working pressure of this vessel is 23 bar, the design pressure will be set at 10%
above the operating pressure, which is 25.3 bar. It is important to set a margin of about 5-10%
Full Report
between the design pressure and operating pressure to avoid spurious operation of the relief
valve during minor process upsets.
The maximum allowable stress will depend on the material operating temperature. The
operating temperature of this vessel is 350 and this temperature depends on the amount of
carbon oxides entering the reactor. Under ASME BPV Code, the maximum design
temperature at which the maximum allowable stress is evaluated should be taken as the
maximum working temperature of the material, with due allowance for any uncertainty
involved in predicting vessel wall temperature (Sinnot & Towler, 2009). The design
temperature of the vessel to be designed is set at 400 .
5.4.3.2 Material Selection

Material selection for the vessel was done based on the process environment and operating
conditions in the vessel. Carbon steel (A285 Grade A) was selected as the material for
construction of the reactor. Carbon steel provides fair resistance to corrosion and it is widely
used in the industry. This material was selected because of its relatively low price and
acceptable material properties for methanation process. It provides the required strength,
good workability and welding properties for a pressure vessel. Although stainless steel has a
higher maximum allowable stress, the maximum allowable stress of carbon steel is sufficient
for this vessel. Moreover, stainless steel is more costly. Thus, carbon steel is chosen. A
corrosion allowance of 4mm will be added to the shell thickness to account for exterior
corrosion. The summary of the material properties selected is shown in the table on the next
page.
Table5.4.3.2-1: Summary of properties of material selected

Parameter Properties
Material Carbon Steel A285 Grade A
Conductivity 43W/m.K
Maximum allowable stress 70MPa
Density 7850kg/m3
Full Report
5.4.3.3 Reactor Feed Positioning and Support

The feed will enter the reactor from the bottom and exits from the top to create a pressure
gradient in the reactor.
A wire mesh and a spacing allowance of 0.6m is placed above the bed to ensure that no
catalyst is blown out the top of the reactor during operation. An allowance of 0.6m is also
included at the bottom of the reactor.
A steel support beam of 0.2m will also be placed under the catalyst bed to support the high
amount of catalyst in the reactor. A support plate above the steel beam will be placed to allow
gas to flow through the reactor.
5.4.3.4 Vapor Distributor

A gas distributor of 50mm thickness and 1m width will be allocated at the bottom of the
reactor to ensure that the gas is evenly distributed across the cross-section of the bed. The
common types of distributor are drilled plate, cap design, continuous horizontal slots,
standpipe design and sparge tube (Rhodes, 2008). In this design, drilled plate is selected.
5.4.3.5 Manway Sizing and Catalyst loading/Unloading

In this design, the manway of 0.5m is placed above the catalyst bed. A compensation ring
will be added to the manway to provide additional support.
The catalyst can be pumped into the reactor as a wet slurry. The water can then be drained
out of the system. This is to ensure that the catalyst is not damaged when it enters the reactor.
The same way will be used to remove the used catalyst.
5.4.3.6 Shell Thickness

According to Sinnot and Towler, the minimum thickness required for a vessel diameter of 1m
is 7mm not including corrosion allowance. This required minimum wall thickness is to ensure
that the vessel is sufficient to withstand its own weight. The vessel wall thickness was
calculated to be 22.47mm, inclusive of a 4mm corrosion allowance. The thickness equation
Full Report
used was specified by the ASME BPV Code (Sec. VIII D.1 Part UG-27). Detailed
calculations are shown in Appendix C4.3.
5.4.3.7 Vessel Ends

The reactor operates at a high pressure, thus it needs a suitable head shape. Ellipsoidal head is
the most suitable and economical choice as the reactor ends. Torispherical heads is lowest in
cost but it is only viable up to 15 bar, which is not suitable for the design of this reactor
which operates at a pressure of about 20 bar. Hemispherical heads are stronger and it is able
to work at much larger pressures but they are relatively expensive. Since ellipsoidal heads is
at the middle ground between pressure and cost, it is selected for this design. An additional
allowance of 4mm was added for corrosion. The thickness of the ellipsoidal head was
calculated to be 18.14mm. However, the thickness of the ellipsoidal heads was taken to be the
same as the vessel wall thickness (22.47mm) for consistency and to simplify the design.
5.4.3.8 Vessel Height

The vessel height was determined based on the required amount catalyst. A reactor diameter
of 1.0m was assigned to give a bed height of 3.24m. A section above and below the catalyst
bed is left void of catalyst to stop entrainment of catalyst particles. 0.6m is left above and
below the bed. The total height of the reactor including the ellipsoidal head is calculated to be
4.98m. Detailed calculations are shown in Appendix C4.2.
5.4.3.9 Insulation
Insulation is required due to the high temperature of the content in the reactor. It will be
added to the outer surface of the reactor to lower the outer wall temperature. This also helps
to prevent potential hazards from happening and also to prevent excessive heat loss to the
atmosphere. Mineral wool will be used as the insulation material for the reactor as it is a good
heat insulator and it is widely used in many industries. An insulation thickness of 0.075m is
used for the reactor. Heat transfer calculations shows that the heat loss is low as the outside
wall temperature is 38.6 , which is close to the estimated ambient temperature of 35 .
Full Report
5.4.3.10 Painting
Paints are used mainly to protect the external surface of the vessel from atmospheric
corrosion. Epoxy-based paint, which is a chemical resistant paint is selected for this vessel as
it undergoes chemical reactions.
5.4.3.11 Stress Analysis
5.4.3.11.1 Design Loads

The weight loadings will first be determined before performing the stress analysis. The
weight loadings of the reactor include the reactor shell, fittings, insulation and process fluid.
For fittings load, 2 ladders were included (inside and outside the reactor) and no platform was
included as the reactor is considered small in size. Steel caged ladders of 360N/m are selected.
Platforms were not added due to the small size of the reactor. A summary of the calculated
weight loadings is shown in the table below. The weight of empty vessel was doubled to take
into account the weight of steel support beam, plate and wire mesh.
Table 5.4.3.11.1-1: Summary of weight loadings of the reactor during operation
Loadings Weight (kN)

Dead load of vessel 69.64
Catalyst 12.79
Fittings 3.63
Insulation 5.99
Process fluid 0.0914
Total weight of vessel during operation, 92.14
Total weight of vessel for hydrotesting, 107.81
2 loading situations were analysed independently. This includes wind loading and hydrostatic
testing. The thickness of 19.46mm of the reactor satisfied the condition for wind loading and
hydrostatic testing. Detailed calculations are shown in Appendix C4.6.
Full Report
Table 5.4.3.11.1-2: Summary of stress values for both normal and hydrotesting
conditions
Design Code: AS1210 Units Design
Material type - Carbon steel
Specification - A285
Grade - A
Wall thickness mm 22.14
Corrosion allowance mm 4
Maximum allowable stress MPa 70
Longitudinal stress, MPa 28.15
Hoop stress, MPa 56.29
Bending stress, MPa
Critical buckling stress, MPa 430.11
Stress analysis Normal operating condition
Longitudinal stress (upwind) MPa 28.27
Longitudinal stress (downwind) MPa 25.47
Stress analysis Hydrostatic test condition
Longitudinal stress (upwind) MPa 28.06
Longitudinal stress (downwind) MPa 25.25
Table 5.4.3.11.1-3: Safety checking of stress values under normal operating

conditions
Safety Checking for normal operating condition Units Satisfied
Maximum allowable stress 70 MPa
Critical buckling stress 430.11 MPa
Maximum difference between principal 
30.82 < 70 MPa
stresses

Checking for critical buckling stress 2.68 < 430.11 MPa
Table 5.4.3.11.1-4: Safety checking of stress values under hydrostatic test

conditions
Safety Checking for hydrostatic test condition Units Satisfied
Maximum tensile stress 70 MPa
Critical buckling stress 374.65 MPa
Maximum difference between principal 
31.04 < 70 MPa
stresses

Checking for critical buckling stress 2.90 < 430.11 MPa
Full Report
Table 5.4.3.11.1-3 and 5.4.3.11.1-4 shows that the stress analysis carried out satisfied the
allowable stress of ASME standard. Detailed calculations are shown in Appendix C4.9.
5.4.3.12 Skirt Support

Skirt support is selected for the vessel support as it is suitable for vertical columns and it is
equally strong in all directions thus able to withstand wind loading. Saddle support was not
chosen as the maximum weight that it can handle is 90kN for a 1m diameter vessel, which is
not sufficient for this vessel design (Sinnot & Towler, 2009).
Table 5.4.3.12-1: Summary of skirt support design

Skirt support design
Skirt material Carbon steel A285 Grade A
Skirt outer diameter m 1.04
Skirt thickness mm 22
Skirt height m 1.0
To check axial compressive stress for skirt design
Bending moment at bottom of tangent line,
Nm 36560.64
Mx
Young Modulus of skirt material MPa 210000
Table 5.4.3.12-2: Safety check stress analysis for skirt support

Design conditions MPa Normal Hydrostatic
Bending stress, MPa 2.28
Dead weight stress, MPa 1.44 1.68
Maximum tensile stress MPa 0.84
Maximum compressive stress MPa 3.96
To check for satisfaction of design criteria Satisfied

Tensile 0.84 < 117.2

compressive 3.96 < 525
The maximum tensile and compressive stresses satisfies the design criteria; 2mm will be
added to the thickness of the skirt support for corrosion. Therefore, a skirt thickness of 22mm
and 1.0m is viable to support the reactor. Detailed calculations are shown in Appendix C4.10.
Full Report
5.4.3.13 Pipe and Nozzle Sizing

Two nozzles are required for this vessel, one for feed inlet and one for the product outlet. The
pipe size and standard is summarized in the table below.
Table 5.4.3.13-1: Summary of pipe size for methanator

Selected pipe standard
Volumetric Calculated
Nominal
Description Phase flow rate Schedule Schedule Velocity
Size
(m3/s) number number (m/s)
(mm)
Feed
Vapour 0.40 36.14 40 250 7.85
syngas
Product
Vapour 0.43 36.14 40 250 8.51
syngas
The pipe size was sized based on the velocity method detailed by Sinnot & Towler (2009).
Both inlet and outlet pipe size for the methanator selected are the nominal pipe size of
250mm with schedule number of 40. Both velocities with the mentioned pipe size are within
the range of gases velocities of 15-30m/s (Sinnot & Towler, 2009). Detailed calculations are
shown in Appendix C4.11.
Full Report
5.4.4 General Arrangement Drawing
5.4.4.1 Mechanical Drawing

Full Report
5.4.4.2 Data Sheet for Methanator
METHANATOR DATA SHEET

Company Alternis BioAmmonia Pvt. Ltd. Description Methanator

Number Required = 1 Function Converts CO and CO2 to CH4
OPERATING DATA
Stream No. IN: stream 69 OUT: stream 70
Reactor Type Adiabatic Fixed-Bed Reactor
Design Vertical
Orientation
Maximum 1.02m Total Height 4.44m
Diameter
Process Data Units Inlet Stream Value Outlet Stream Value
o
Temperature C 300 340.01
Pressure kPa (abs) 2300 2300
STREAM PROPERTIES
Component Units Inlet Stream Value Outlet Stream Value
Carbon kg/h 48.06 TRACE
Monoxide
Carbon Dioxide kg/h 3.06 TRACE
Water kg/h 93.69 127.10
Methane kg/h 28.13 56.77
Hydrogen kg/h 1046.46 1035.52
Nitrogen kg/h 4670.46 4670.46
Properties Units Inlet Stream Value Outlet Stream Value
Vapour Fraction - 1 1
Density kg/m3 4.093 3.9365
Viscosity Pa.s 2.0 x10-5 2.0x10-5
Thermal W/m.K 0.18 0.19
Conductivity
CATALYST PROPERTIES
Name Haldor Topsøe PK-7R
Type Ni/MgAl2O3
Shape Extruded ring
Properties Units Value
Catalyst Bulk 512.6
kg/m3
Density
Void Fraction - 0.475
Equivalent 6.8
mm
Sphere Diameter
Catalyst Mass kg 1304
Catalyst Bed 2.54
m3
Volume
Full Report
SHELL MATERIAL OF CONSTRUCTION

Shell Material Carbon steel A285 Grade A
Material Density kg/m3 7850
Material 43
W/m.K
Conductivity
Material 70
Maximum
Allowable Stress MPa
at Design
Temperature
Shell Inner 1.0
m
Diameter
Shell Thickness m 22.47x10-3
INSULATION
Insulation Mineral Wool
Material
Thickness of mm 75
Insulation
Density of kg/m3 130
Insulation
Conductivity of W/m.K 0.04
Insulation
Mass of kg 610.52
Insulation
Reactor Wall 38.6
Temperature
Temperature 3.6
Difference
(Reactor Wall -
Ambient)
REACTOR SUPPORT
Support Type Skirt
Support Material Carbon steel A285 Grade A
Material Density kg/m3 7850
Material W/m.K 0.04
Conductivity
Support mm 22
Thickness
(Including
Corrosion
Allowance)
Support Height m 1.0
NOZZLE-PIPE SIZING
Nozzle Material Carbon steel A285 Grade A
Properties Units Inlet Stream Value Outlet Stream Value
3
Material Density kg/m 7850
Material W/m.K 0.04
Conductivity
Volumetric Flow m3/s 0.40 0.43
Full Report
Rate
Velocity m/s 7.85 8.51
Schedule - 40 40
Number
Outside mm 273.1 273.1
Diameter
Wall Thickness mm 9.27 9.27
Nominal mm 250 250
Diameter
Full Report
5.5 Detailed Process and Mechanical Design: Waste Heat Boiler (WHB-101,
WHB-102, WHB-103)
5.5.1 Definition of Design and Specifications
Waste Heat Boiler (WHB) or also known as Heat Recovery Steam (HRSG) are generally
used to recover energy from high temperature syngas or flue gas produced in chemical plants
in order to generate steam that will be utilized in the power generation steam turbine system.
In the Alternis BioAmmonia plant, the dual fluidized bed gasifier produces of
syngas at a temperature of and at an elevated pressure of . This high
temperature syngas need to be cooled down before being supplied to the downstream utilities
of syngas cleaning which operates at temperature below . Therefore, it is wise to
recover the energy of the syngas to generate superheated steam which can be used to produce
electricity.
The Waste Heat Boiler is designed to be consisting of 3 sections, a super heater, kettle
evaporator, and an economizer operating at a single pressure. Since high pressure steam
should be generated for the efficient use in power generation process, the water inlet to the
economizer need to be supplied at high pressure, where by in this plant it is pumped up to
. The compressed water will be heated to the saturation temperature of the water at
and will be then introduced into the kettle evaporator which will then convert it to
saturated steam. The saturated steam will be superheated in the super heater and will be
supplied to steam turbine. The temperature profile of the syngas and steam as well as the
amount of steam generation is much affected by the pinch and approach point that to be
selected based inlet syngas temperature. The pinch point is the difference between the gas
temperature leaving the evaporator and the temperature of saturated steam. The approach
point is the difference between the temperature of saturated steam and the temperature of the
water entering the evaporator. Based on the suggested WHB temperature profile by
V.Ganapathy in his article titled “Heat Recovery Steam Generators: Understand the
Basics”(Ganapathy, 2001), the range of pinch and approach point is and
respectively for syngas inlet temperature ranging from to . Therefore, the
temperature profile that is used in this design based on the energy and mass balance is as
shown in figure below:
Full Report
Superheater Kettle Evaporator Economizer
Figure 5.5.1.1: Temperature Pofile
The waste heat boiler designed according to the temperature profile above is capable of
producing of superheated steam. One of the limitations that were considered
during the design is the final temperature of the syngas as need to be cooled down to a
temperature above the dew point, in order to prevent the condensation of the tar
content in the syngas. Another matter that was considered is the fluid allocation for the casing
or the shell side and the tube side, whereby, the syngas was allocated in the tube and the
water to flow in the shell. Since the compressed water is being heated to be a steam, it
requires more volume to expand as it acquires heat from the syngas.
Both economizer and superheater operate very similarly to heat exchangers and
therefore, they were designed using the Kern’s method as explained in the Chemical
Engineering Design book by Ray Sinnot and Gavin Towler (2009). For detailed mechanical
design, TEMA (Tubular Exchanger Manufacturers Association) (NPTEL, 2003) standards
were referred as well. Meanwhile, the kettle evaporator was designed based “Industrial
Boilers and Heat Recovery Steam Generators: Design, Applications and Calculation” book
(Ganapathy, 2003) and Chemical Engineering Design (Sinnott and Towler, 2009)was also
referred for this design.
5.5.2.1 Economizer and Superheater as Heat Exchanger

As mentioned earlier, the economizer and superheater are designed as shell and tube heat
exchangers as there is no phase changes take place. In economizer, the compressed
Full Report
water at is being heated to a temperature slightly lower than the saturation

temperature which is to . In Superheater, the saturated steam at temperature is
being superheated to a temperature of . The maximum allowed pressure drop in both
heat exchangers is taken as 0.5 bar. The standard step-by step Kern’s Method was used for
the design of these two sections of the Waste Heat Boiler (Sinnott and Towler, 2009). The
steps are explained below:
Step 1: Mass and Energy balance is carried out based on the temperature profile shown above
to define the Duty and the maximum flow-rate of steam to be generated.
Step 2: Physical Properties of both inlet and outlet of shell and tube side streams is obtained
and tabulated. The values were obtained from PRO-II Simulation. For calculation purpose,
the mean values of the physical properties of the shell and tube side streams were used. The
fouling factors for both streams were obtained referring to the Table 12.2 of (Sinnott and
Towler, 2009).
Step 3: Value of Overall Coefficient, is obtained from PRO-II Simulation
Step 4: Number of Shell and Passes is chosen. An even number of tube pass is preferred as it
allows the inlet and outlet nozzles to be at the same end of the HEX simplifying the
pipeworks. For both section, one shell pass and 2 tube pass is chosen as initial guess.
Step 5: Log Mean Temperature Difference, , Correction Factor, , and are

calculated referring to (Sinnott and Towler, 2009).
Step 6: Provisional Heat Transfer Area is determined by dividing the duty calculated by the
and
Step 7: The type of HEX and the TEMA Codes is decided. The economizer is designed as
AEL type and the superheater as type AES. The initial guess of the tube dimensions, inner
and outer diameter and tube length is decided. The tube arrangement for both sections is
decided to be arranged in equilateral triangular pattern as it gives higher heat transfer rates.
Tube pitch is taken as 1.25 times the tube outer dimension.
Step 8: The number of tubes are calculated. The tube side velocity at this stage is calculated
to check if it is reasonable. If the tube side velocity is very high meaning that the residence
time for the heat transfer to occur is very limited and therefore might not achieve the required
Full Report
duty. Therefore, iteration on the number of tube is done to achieve a reasonable velocity.
Increasing the number of tubes decreases the tube side velocity.
Step 9: The shell diameter is calculated.
Step 10: Tube side heat transfer coefficient is calculated
Step 11: Shell side heat transfer heat transfer coefficient is calculated
Step 12: This step is only for Superheater. Due to high operating temperature, the
nonluminous heat transfer plays a significant role. Therefore, the non-luminous radiation heat
transfer coefficient is evaluated. To estimate the radiation heat transfer coefficient the partial
pressure of the Carbon dioxide and Water vapor in the syngas and the beam length which
equals to the inner diameter of tube is obtained. The emissivity of the Carbon Dioxide and
water vapor at 1 bar is first determined and then multiplied with the pressure correction factor
calculated to obtain the emissivity at 5 bar (Chemieingenieurwesen and Gesellschaft, 2010).
The decrease in emissivity due to presence of carbon dioxide and water vapor is calculated to
obtain the overall emissivity of the syngas. Finally, non-luminous radiation heat transfer
coefficient is calculated using the Boltzmann coefficient.
Step 13: Overall heat transfer Coefficient, including the fouling factors is calculated and
compared to the initial . Reiterations were done by altering the tube dimensions and the
number of tubes.
Step 14: Tube and Shell side pressure drop is calculated and checked if they are within the
specification. If yes, the design is accepted.
For detailed hand calculation, please refer to Appendix C5.
5.5.2.2 Kettle Evaporator as Fired tube Boiler

The kettle evaporator converts the water at to saturated steam at and .
There is phase change occurring in the kettle evaporator. Since the water enters the kettle
evaporator at temperature slightly lower than that of the saturation temperature, the duty of
the kettle evaporator accounts for both sensible heat from the inlet temperature, to the
saturation temperature of and the latent heat to convert the saturated water to
saturated steam at constant temperature. The mode of heat transfer is assumed to be
convective boiling. The steps used to determine the design and sizing of kettle evaporator is
as follow:
Full Report
Step 1: Mass and Energy balance is carried out based on the temperature profile shown above
to define the Duty and the maximum flow-rate of steam to be generated.
Step 2: Physical Properties of both inlet and outlet of shell and tube side streams is obtained
and tabulated. The values were obtained from PRO-II Simulation. For calculation purpose,
the mean values of the physical properties of the shell and tube side streams were used. The
fouling factors for both streams were obtained referring to the Table 12.2 of (Sinnott and
Towler, 2009).
Step 3: One shell pass and 2 tube pass is chosen as initial guess.
Step 4: Log Mean Temperature Difference, , Correction Factor, , and are

calculated referring to (Sinnott and Towler, 2009)
Step 5: The initial guess of the tube dimensions, inner and outer diameter and tube pitch
decided. The tube arrangement is decided to be arranged in rotated square pattern as it gives
higher heat transfer rates and ease of cleaning. Tube pitch is taken as 1.5 times the tube outer
dimension. An initial guess of number of tubes is made.
Step 6: The type of HEX and the TEMA Codes is decided. The kettle evaporator type AKL is
decided.
Step 7: Convective Heat Transfer Coefficient at the tube side is calculated.
Step 8: The wall surface temperature is calculated. The shell-side heat transfer coefficient is
then calculated by using reduced pressure correlation given by Mostinski (1963) (Sinnott and
Towler, 2009).
Step 9: Overall heat transfer coefficient including the fouling factors is calculated and
compared with PRO-II Value.
Step 10: The total heat transfer area that is required to achieve the duty based on the overall
heat transfer coefficient is calculated.
Step 11: The minimum required length of tube is then calculated and was standardized
according to the available tube length in industry.
Step 12: Kettle evaporator layout is then decided starting with determination of the tube
bundle diameter which is the similar method as the general shell and tube heat exchanger.
Full Report
Step 13: The inner shell diameter is then calculated by taking it to be 1.5 times the tube
bundle diameter.
Step 14: Since the freeboard between the liquid level and shell should be at least 0.25 m
(Sinnott and Towler, 2009), an initial freeboard height is taken. The liquid level is then
calculated and ensured it to be higher than the tube bundle diameter.
Step 15: The width at liquid level is calculated as the chord length of the freeboard segment.
The surface area of the liquid level is calculated by multiplying the width of liquid level and
the effective length of shell.
Step 16: The vapor velocity at the water surface in the kettle evaporator is calculated. The
maximum allowable steam velocity is calculated. The actual vapor velocity is checked if it is
below the maximum allowable steam velocity. If not iteration on the shell diameter is done.
Step 17: Tube side velocity and pressure drop is calculated and checked if they are reasonable.
The tube side pressure drop is ensured to be below .
Step 18: Shell side velocity and pressure drop is calculated and checked if they are reasonable.
The tube side pressure drop is ensured to be below .
For detailed hand calculation, please refer to Appendix C5.
5.5.3.1 Materials of Construction

For all 3 section of the waste heat boiler, shell and tube side, the design pressure and
temperature were estimated. Since water or steam at operating pressure of 50 bar is flowing
in the shell side, the design pressure for the shell side is taken as 55 bar (10% greater that
operating pressure). As for the tube side, the syngas flows at 5 bar, therefore the design
pressure is taken to be 10% greater at 5.5 bar. As for the design temperature at both shell and
tube side was taken greater than the maximum temperature of the stream. The type of
material that to be used for construction were then determined based on the design pressure
and temperature depending on the tensile strength of the material. The table below lists down
the design pressure, temperature, type of material to be used and the tensile strength obtained
from Figure 25-22 and Table 25-11 from Perry’s Chemical Engineer’s Handbook(Wiebert et
al., 2008).
Table 5.5.3.1.1: Materials of Construction

Full Report
Equipment Superheater Evaporator Economizer

Unit shell tube shell tube shell tube
Design
kpa 5500 550 5500 550 5500 550
Pressure
Design
437 810 274 545 210 267
Temperature
Stainless Stainless Stainless Stainless Stainless Carbon
Material Steel AISI Steel AISI Steel AISI Steel AISI Steel AISI Steel A285
Type 310 Type 310 Type 310 Type 310 Type 304 Gr A
Max Allowed
kpa 275800 34475 310000 137900 241000 85380
Tensile Stress
5.5.3.2 Stress Analysis
The stress analysis was done for all three sections of the waste heat boiler and similar method
was used as explained in the Chemical Engineering Design by Ray Sinnott and Gavin Towler
(2009). For all three sections, the primary stresses, longitudinal, and circumferential,
stresses were done. This followed by the dead weight calculation for both shell and tube
and the total dead weight stress was calculated. Since all 3 equipment are designed as
horizontal shells and the maximum height that is reached are well below 2 m, the bending
and wind loading stress are assumed to be negligible. The earthquake or seismic loading is
also not considered as the location of the plant is not prone to earthquake. Buckling test was
done to ensure that the compressive stress applied by the fluid in the waste heat boiler is well
below than the maximum allowable axial compressive load.
5.5.3.2.1 Economizer
Based on the thermal design, the required inner shell diameter was calculated as
taking into consideration of the clearance diameter. The thickness of the shell however
depends on the design pressure and the maximum allowable stress that the Stainless Steel
AISI Type 304 could withstand which is at the operating temperature and pressure
obtained from Perry’s Handbook (Wiebert et al., 2008). The maximum initial shell thickness
that was calculated for the economizer is . The stress analysis is then done on the
shell to make sure that the shell thickness is adequate to withstand the stress being applied on
it. The calculation method is followed as provided in the Chemical Engineering Design by
Ray Sinnott and Gavin Towler (2009) and shown in the Appendix C5. The design of the
economizer with shell thickness of 21.7 mm results in all the stresses that are calculated
above are well below the maximum tensile strength, the design can be further optimized by
reducing the thickness of the shell as it will reduce the cost of the material used. Iterations
Full Report
were done with different thickness and the final iteration with The
Table below gives the final stresses and results.
Table 5.5.3.2.1.1: Stress analysis for Economizer

Shell Thickness,
Longitudinal Stress, 110035.8377
Circumferential Stress, 220071.6753
Shell mean Diameter, 0.8103
Shell Dead Weight, 1.4038
Total Dead Weight, 11.0232
Dead Weight Stress, 433.0443
Resultant Longitudinal stress, 109602.7934
principal stress, 220071.6753
principal stress, 2750
Maximum Stress intensity, 110468.8819
Maximum compressive stress allowed, 237115.0042
Design compressive stress, 109602.7934
5.5.3.2.2 Evaporator
obtained from Perry’s HandbookWiebert et al., 2008). The maximum initial shell thickness
that was calculated for the economizer is . The stress analysis is then done on the shell
to make sure that the shell thickness is adequate to withstand the stress being applied on it.
The similar calculation method is followed and shown in the Appendix C5. The design of the
kettle evaporator with shell thickness of 25 mm results in all the stresses that are calculated
above are well below the maximum tensile strength, the design can be further optimized by
Full Report
Table below gives the final stresses and results which satisfies the tensile test and buckling
test.
Table 5.5.3.2.2.1 : Stress Analysis Kettle Evaporator

Shell Thickness,
5.5.3.2.3 Superheater
obtained from Perry’s Handbook (Wiebert et al., 2008). The maximum initial shell thickness
that was calculated for the superheater is . The stress analysis is then done on the shell
to make sure that the shell thickness is adequate to withstand the stress being applied on it.
The similar calculation method is followed and shown in the Appendix C5. The design of the
superheater with shell thickness of 35 mm results in all the stresses that are calculated above
are well below the maximum tensile strength, the design can be further optimized by
Full Report
Table below gives the final stresses and results which satisfies the tensile test and buckling
test.
Table 5.5.3.2.3.1 Superheater Stress Analysis

Shell Thickness,
5.5.3.3 Piping and Nozzle

The flow of the inlet and outlet streams for all 3 sections of the waste heat boiler need to be
maintained at the required velocity and pressure as well as taking into consideration the
corrosion allowance. Detailed explanation and calculation is shown in Appendix C5. The
nozzle dimensions were obtained from the ASME Standards(Steels, 2008). Table below
shows the final nozzle dimensions that is used for the design.
Table 5.5.3.3.1: Nozzle Dimensions

Outside Inner
Thickness
Nozzle NP Pipe/Nozzle Pipe/Nozzle
Description Type of nozzle
No. S Diameter, Diameter,
(mm)
1 Syngas into Superheater SS Sch 5S 10 273.1 3.4 266.3
Full Report
2 Syngas out of Superheater SS Sch 5S 10 273.1 3.4 266.3

Saturated steam into SS Sch
3 10 273.1 9.27 254.56
superheater 40S
Superheated steam out of SS Sch
4 10 273.1 9.27 254.56
superheater 40S
5 Syngas into boiler SS Sch 5S 10 273.1 3.4 266.3
6 Syngas out of boiler SS Sch 5S 10 273.1 3.4 266.3
SS Sch
7 Saturated water into boiler 10 273.1 9.27 254.56
40S
Saturate steam out of SS Sch
8 10 273.1 9.27 254.56
boiler 40S
9 Syngas into economizer SS Sch 5S 10 273.1 3.4 202.74
10 Syngas out of economizer CS Sch 20 8 219.1 6.35 206.4
Compressed water into SS Sch
11 8 219.1 8.18 202.74
economizer 40S
Saturated water out of SS Sch
12 10 273.1 9.27 202.74
economizer 40S
5.5.3.4 Flange Dimensions
For the 12 nozzles as mentioned above, the flange class numbers required were determined
based on the design temperature and pressure, and the material of construction. The designs
and dimensions of standard flanges were obtained from the ASME B16.5 Annex F(B16.5-
2003, 2003). The Table below shows the nozzle, design pressure and temperature and the
respective flange class and the group of material selected.
Table 5.5.3.4.1: Flange Group Material an d Class

Nozzle Temperature Pressure Material Type of Material Flange
Description
No. ( ) (psig) Group Class
1 Syngas into Superheater 818 65.3 1.2 Carbon A350 150
2 Syngas out of Superheater 818 65.3 1.2 Carbon A350 150
Saturated steam into Carbon A182
3 1490 782.3 2.4 2500
superheater
Superheated steam out of Carbon A182
4 1490 782.3 2.4 2500
superheater
5 Syngas into boiler 1013 65.3 1.7 Carbon A182 300
6 Syngas out of boiler 1013 65.3 1.7 Carbon A182 300
7 Saturated water into boiler 525 782.3 1.2 Carbon A350 300
8 Saturate steam out of boiler 525 782.3 1.2 Carbon A350 300
9 Syngas into economizer 512 65.3 1.2 Carbon A350 150
10 Syngas out of economizer 512 65.3 1.2 Carbon A350 150
Compressed water into Carbon A350
11 410 782.3 1.2 300
economizer
Saturated water out of Carbon A350
12 410 782.3 1.2 300
economizer
Table below shows the dimensions of the dimensions of the flanges for each of the nozzles.
Full Report
Figure 5.5.3.4.1: Welding Neck Dimension (B16.5-2003, 2003)
Table 5.5.3.4.2: Flange Dimensions

Outside Hub
Nominal Length
Diameter Thickness Diameter of Diameter
Wielding Pipe through Bore, B
Class of of flange, Hub, beginning of
neck (M) Size hub, Y (mm)
flange, (mm) (mm) chamfer, A
(NPS) (mm)
(mm) (mm)
N1 10 150 436.96 30.5872 327.72 293.5825 107.6014 273.6462
N2 10 150 436.96 30.5872 327.72 293.5825 107.6014 273.6462
N3 10 2500 723.715 181.6115 402.8225 293.5825 450.615 273.6462
N4 10 2500 723.715 181.6115 402.8225 293.5825 450.615 273.6462
N5 10 300 477.925 49.4311 344.6522 293.5825 124.5336 273.6462
N6 10 300 477.925 49.4311 344.6522 293.5825 124.5336 273.6462
N7 10 300 477.925 49.4311 344.6522 293.5825 124.5336 273.6462
N8 10 300 477.925 49.4311 344.6522 293.5825 124.5336 273.6462
N9 10 150 436.96 30.5872 327.72 293.5825 107.6014 273.6462
N10 8 150 378.675 29.733 271.8045 242.0715 110.517 223.839
N11 8 300 420.75 43.758 287.5125 242.0715 121.176 223.839
N12 10 300 477.925 49.4311 344.6522 293.5825 124.5336 273.6462
Full Report
5.5.3.5 Support System: Saddles
Figure 5.5.3.5.1: BN-DS-M01 Saddles for horizontal vessels Dimension

Dimension
Superheater Evaporator Economizer
Label (mm)
A 950 760 580
B 1025 875 725
C 820 630 450
D 345 250 160
E 16 16 16
F 200 200 200
G 120 120 120
H 10 10 10
J 270 270 250
Bolt M24 M24 m24
Bolt hole 30 30 30
Mass(kg) 235 205 130
Full Report
5.5.4 Specific Data Sheet and mechanical design drawing

Full Report
5.5.4.2 Specific Data Sheet
Waste Heat Boiler

Mechanical Design Specification Sheet
Company Alternis BioAmmonia Pvt. Ltd. Description Waste Heat Boiler

Equipment No. WHB-101,WHB-102, WHB-103 Sheet No. 1
Recovery of heat syngas produced for
Number Required = 1 Functions
In-plant Power Generation
1st Unit - Superheater
Type of Shell and Tube Heat Exchanger AES
Side Shell Tube
Streams Inlet Outlet Inlet Outlet
Type of Fluid High Pressure Steam Syngas
Mass Flow Rate 7320.2506 15859.7461
Pressure 50 50 5 5
Temperature 263.9000 426.8910 840.0000 760.0000
Specific Heat 2.5294 2.3170 2.2550 2.2126
Thermal
0.0395 0.0632 0.1338 0.1260
Conductivity
Density 24.4681 16.5915 1.0500 1.1314
Viscosity 0.0180 0.0260 0.0336 0.0318
Fouling Factor 0.0001 0.0003
Allowed Pressure
0.5000 0.5000
Drop
Pressure Drop
0.4274 0.2305
Calculated
Heat Duty 803.1045
Cross Flow Area 0.0889 0.3295
Heat Transfer
172.9265 175.7791
Coefficient
Radiation Heat
72.5644
Transfer Coefficient
Overall Heat
92.9305
Construction and Material
Side Shell Tube
Stainless Steel AISI Type Stainless Steel AISI Type
Type of Material
310 310
Design Pressure 5500 550
Design Temperature 437 810
Number of Pass 1 2
Shell Tube
Inner Diameter 1.4906 Inner Diameter 45.8
Full Report
Outer Diameter 1.5306 Outer Diameter 50.0

Thickness 20 Thickness 2.10
Baffle Cut 25.00 Total Length 5.00
Connection
Baffle Spacing 298 0.12
Allowance
Shell Cover Type Torispherical Head Length 4.88
Inside depth 132 N tubes 400
Shell Side Velocity 1.1144 Pitch Pattern Triangle
Tube Pitch 62.5
Linear Velocity 12.2586
Stress Analysis
Maximum Tensile principal stress,
275800 204960
Strength,
Longitudinal Stress, principal stress,
102480 101396
Circumferential principal stress,
204960 2750
Stress,
Shell mean Diameter, Maximum Stress
1.5106 103563
intensity,
Shell Dead Weight, Maximum Stress
54.4166 202209
intensity,
Total Dead Weight, Maximum Stress
102.8245 98646
intensity,
Dead Weight Stress, Maximum
1083.3375 compressive stress 256483
allowed,
Resultant Design compressive
Longitudinal stress, 101396.2684 stress, 101396
2nd Unit - Kettle Evaporator

Type of Shell and Tube Heat Exchanger AKL
Side Shell Tube
Mass Flow Rate 7320.2506 15859.7461
Pressure 50 50 5 5
Temperature 200.0000 263.9000 760.0000 310.0000
Specific Heat 4.6285 2.5294 2.2126 1.9064
Thermal
0.6646 0.0395 0.1260 0.0738
Conductivity
Density 857.0380 24.4681 1.1314 2.0098
Viscosity 0.1338 0.0180 0.0318 0.0211
Allowed Pressure
0.5000 0.5000
Drop
Pressure Drop
0.000003 0.1189
Calculated
Full Report
Heat Duty 4082.8188

Heat Transfer
63097.3040 415.1638
Coefficient
Overall Heat
294.9408
Side Shell Tube
Stainless Steel AISI Type Stainless Steel AISI Type
Type of Material
310 310
Design Temperature 273.9 545
Number of Pass 1 2
Shell Tube
Shell Cover Type Torispherical Head Total Length 5.00
Connection
Inside depth 96.6 0.12
Allowance
Shell Side Velocity 0.0009 Length 4.88
Max Allowable
1.1667 N tubes 90
Velocity
Pitch Pattern Square
Tube Pitch 76.2
Stress Analysis
310000 224078
Strength,
112039 111306
224078 2750
Stress,
Shell mean Diameter, Maximum Stress
1.2372 112772
intensity,
29.2487 2213278
intensity,
42.7357 108556
intensity,
732.9860 compressive stress 235663
allowed,
111306 111306
Longitudinalstress, stress,
3rd Unit - Economiser
Type of Shell and Tube Heat Exchanger AEL
Side Shell Tube
Full Report

Mass Flow Rate 7320.2506 15859.7461
Pressure 50 50 5 5
Temperature 100.4749 200.0000 310.0000 204.0213
Specific Heat 4.0991 4.6285 1.9064 1.8269
Thermal
0.6809 0.6646 0.0738 0.0606
Conductivity
Density 949.6661 857.0380 2.0098 2.4631
Viscosity 0.2776 0.1338 0.0211 0.0174
Allowed Pressure
0.5000 0.5000
Drop
Pressure Drop
0.0065 0.0470
Calculated
Heat Duty 883.1172
Heat Transfer
1427.5766 365.8302
Coefficient
Overall Heat
240.3926
Side Shell Tube
Stainless Steel AISI Type
Type of Material Carbon Steel A285 Gr A
304
Design Temperature 210 267
Number of Pass 1 2
Shell Tube
Baffle Cut 25 Total Length 3.66
Connection
Baffle Spacing 0.1601 0.05
Allowance
Shell Side Velocity 0.0879 Length 3.61
N tubes 108
Pitch Pattern Triangle
Tube Pitch 62.5
Stress Analysis
110036 220072
Strength,
220072 109603
0.8103 2750
Stress,
Shell mean Diameter, 1.4038 Maximum Stress 110469
Full Report
intensity,
11.0232 217322
intensity,
433.0443 106853
intensity,
109603 compressive stress 237115
allowed,
110036 109603
Longitudinalstress, stress,
Remarks
Prepared By Nisha Thavamoney Date 14th January 2014
Checked By Jenny Yap Wee Li Date 15th January 2014
Full Report
5.6 Detailed Process and Mechanical Design: Synthesis Reactor

Ammonia synthesis is the major and first section in the manufacturing of nitrous chemical
products which include fertilizers, explosives component and plastics. Oil palm trunks from
oil palm plantation will be used as the main feedstock throughout the whole process until
the final product (anhydrous ammonia) is obtained. The direct synthesis process required
basis gases of hydrogen and nitrogen with the ratio of 3 to 1. The 2-bed ammonia reactor
has several advantages as it required low operating pressure (150bar) and low energy
consumption (28GJ/tonne NH3) as compared to other ammonia reactor (Appl, 2012). Also,
Topsoe-200 ammonia reactor has a relatively low catalyst (iron) cost compared to other
ruthenium catalyst based reactor. Ammonia synthesis reactor operates at a temperature of
500 and a pressure of 150bar, also it’s an exothermic reaction.
Gas from Methanator Topsoe-200

Ammonia product to
Ammonia refrigeration loop and
Synthesis storage
Unreacted Recycle gas
Reactor
Figure 5.6.1.1: Illustration of main process stream related to Topso e-200

Ammonia Reactor

For the basis of performance, main assumptions and calculations were performed in the
sizing of the main body of the ammonia reactor. All the detailed calculation and
assumptions which related to the sizing of the reactor can refer to the Appendix C6.

Iron catalyst will be used in ammonia synthesis reactor with its component listed in table
5.6.3.1. The type of iron catalyst used will be the AmoMax®-10RS type and catalyst particle
are spherical assumed to be 2 mm in diameter. The catalyst support in this reactor will be
Full Report
iron oxide. Based on the calculation done and shown in the appendix C6, the actual particle
density of the iron catalyst particle is 6635.5 kg/m3.
Table 5.6.3.1: Components of Amomax-10RS Catalyst

Component Percentage (w/w%) Density (kg/m3)
Iron 60 7874
Iron (III) Oxide 30 5240
Aluminium Oxide 4 3950
Calcium Oxide 3 3350
Potassium Oxide 2 2350
Vanadium Pentoxide 1 3360
5.6.3.1 Design of Catalyst beds

The weight of catalyst used was solved with the aid of Polymath software by using the
general formula of differential equation.
The rate of reaction for packed bed reactor can be determined as below:
Where rate of reaction is, is the equilibrium constant, is the partial pressure of
Ammonia, is the partial pressure of Nitrogen, is the partial pressure of hydrogen.
By conducting Polymath file, the total weight of catalyst required can be obtained at the
specific conversion of each bed. The weight of catalyst required in 1 st bed is 1132 kg and 372
kg of catalyst in 2nd bed. At the same time, the volume of each catalyst bed can be
calculated by dividing the calculated catalyst weight by the bulk density of the catalyst. The
thickness and height of each catalyst can be determined by using the following equation as
shown in Appendix C6.
Where is the density of the catalyst particle, is the voidage. The bed height to diameter
ratio was set to be 1.25 and thickness of the catalyst bed was set to be 1/10 of diameter.
Full Report
Based on calculation, it was obtained that the diameter and height of the 1 st bed is 1.2432 m
and 1.554 m. The optimal diameter for the 2nd bed was then also set to be 1.2432 m
whereas the height of 2nd bed is calculated as 1.073 m. Please refer to the Appendix C6 for
detailed calculation.
5.6.3.2 Design of Intercooler

Apart from designing the catalyst bed, the intercooler or the heat exchanger that located in the
hollow section of catalyst bed is required to be design as well. The mechanical design
calculation for the design parameters such as diameter and length of shell and tube was
calculated using Kern’s method. Outer diameter, inner diameter and length of the tubes was
initially assumed as 38 mm, 33.8 mm and 1.83 m respectively. Based on the calculation using
Kern’s method, the number of tubes that to be used in the intercooler was determined to be
196. The overall shell diameter of the intercooler heat exchanger was calculated to be 0.8478
m. From the design calculation of the 1st catalyst bed, the diameter of the catalyst bed left
after desucting the thickness of catalyst bed is 0.9946 m indicating that the intercooler design
can be fit into the hollow section of the catalyst bed.
5.6.3.3 Design of the reactor column

First the height of the reactor column was determined. An allowance of 1.2 m in the height
is allocated for the manhole at the top part of the vessel. Besides, allowance of space of 1 m
in between the catalyst beds will be included for the channel of piping system. Since the
intercooler/heat exchanger will be placed in between the catalyst bed, thus the heat
exchanger which is taller will be encounter as the cylindrical section of vessel height instead
of the catalyst bed itself. This is applicable to the 1st catalyst bed. The total cylindrical height
of the reactor can be summarized below:
Height of manhole : 1.200 m

Height of 1st bed : 1.830 m
Height of 2nd catalyst bed : 1.073 m
Space between beds : 1.000 m
Total : 5.103 m
5.6.4 Detailed Mechanical Design

Some major assumptions and detailed calculations for mechanical design of the ammonia
synthesis reactor are interpreted and discussed in the following sections. All the
assumptions and detailed calculations in designing the reactor are thoroughly explained in
Full Report
Appendix C6. For the detailed design purpose, the design pressure is determined to be 10%
higher than the oprating pressure after considering safety factor. The cylindrical shell will be
used in designing the reactor and the hemispherical heads are installed for the reactor both
at top and bottom. For safety factor and space allowance, the inner diameter of the reactor
is set to be 1.5 m slightly higher than the calculated inner diameter above.
5.6.4.1 Material of Construction

At ammonia synthesis section, all the pipelines are dealing with large amount of ammonia
and hydrogen at high pressure and temperature. Ammonia is a corrosive substance and large
amount of hydrogen in piping will cause hydrogen embrittlement. Thus, stainless steel 304
will be used for all the piping, reactor wall and heat exchanger for safety purposes. Grade 304
is the standard”18/8” stainless and it has the excellent forming and welding characteristic
when compared to carbon steel. Post-weld annealing is not required when welding is done at
the thin sections of the reactor. Properties of stainless steel 304 are listed as below:
Density 7999 kg/m3
Thermal Conductivity 16 W/m.K
Maximum Allowable Stress 103 MPa
For vessel support, a skirt support was used for the ammonia synthesis reactor and the
material of construction of skirt was Carbon Steel A285 as the flowing fluid around the skirt
is just the atmospheric air.
5.6.4.2 Operating Conditions and design parameter

The operating pressure and maximum temperature for this reactor are 15 MPa and 476 .
Therefore, the design pressure with 10% allowance is 16.5 MPa and the design temperature
is set to be 500 . The wall thickness of the reactor according to the design temperature
and pressure as well as the maximum tensile strength of the construction material, was
calculated to be 124.06 mm. Apart from that, 2mm of corrosion allowance is added to the
reactor wall thickness and giving the total thickness as 126.06 mm. Based on the Chemical
Engineering Design book (Sinnott & Towler, 2009), the minimum practical wall thickness for
a vessel with diameter of 1 to 2 m is 7mm. Therefore, the calculated wall thickness is
applicable in this design.
5.6.4.3 Mechanical design of hemispherical Heads

Hemispherical head for reactor will be used due to the high operating pressure. The internal
diameter of the reactor is assumed to be 1.5m and thus the height of the hemispherical head
Full Report
will be equal to the radius of the vessel which gives a value of 0.75m. Calculation is done to
determine the minimum thickness required for the hemispherical head and the calculation is
shown in the Appendix C6. The required thickness was determined to be 58.98 mm including
2 mm of corrosion allowance. However, for consistency and the purpose of simplifying the
design, the thickness of the head is set to be the same as the thickness of the vessel which is
124.06 mm. Applying a thicker wall than the minimum wall thickness calculated above
strengthens the safety of the equipment further.
5.6.4.4 Mechanical Design of Insulation

Insulation of the reactor is to prevent heat loss to the surrounding. All the detailed calculation
and assumptions on the insulation are showed in Appendix C6. Based on the calculated wall
temperature with the insulation thickness of 25 mm, it is sufficient to ensure the
temperature difference between the wall of the reactor and the surrounding to be less than
. The small temperature difference ensures that the heat loss from the reactor to the
surrounding is acceptable. Similar calculation to determine the wall temperature of the
reactor was done when the insulation thickness is assumed to be negligible. The
temperature difference is still lesser than which is . Therefore, the mineral wool
insulation thickness for the reactor is considered as negligible.
5.6.4.5 Dead weight of Ammonia Synthesis Reactor

The total dead weight of reactor can be divided into several sections. The detailed
calculations involved in determining the dead weight of the reactor are clearly shown in
Appendix C6. The dead weight of the column, is estimated as follow:
where, Dead weight of the shell,

Factor accounting for the weight of nozzles, internal supports, baffles
Density of column material,
mean diameter of column, m
Length of reactor column, m
The is assumed as 1.15 for a vessel with several man ways, plate support and internal
fittings. For the intercooler, is assumed as 1.08. The support of catalyst (wire mesh) will
be taken into account when calculating the dead weight of the reactor. While for the weight
of fittings, a caged ladders and platform will be installed to the reactor. By assuming the total
height of ladder is the same as the internal height of the reactor (5.103 m). Hydraulic pressure
Full Report
testing is done to check for any possible leakages in the reactor after fabrication. Therefore,
only the reactor weight and weight of water will be used for pressure testing. Table below
shows the total dead weight of the reactor.
Table 5.4.5.1: Dead weight of reactor during normal operation
Component Dead Weight
Weight of Reactor Column
Weight of Intercooler
Weight of Fittings
Weight of Process Fluid
Weight of Catalyst
Total Weight during normal operation
Weight of water for hydraulic pressure testing
Total Weight during hydraulic pressure testing
The wind loading per unit length, of the column can be obtained from the wind pressure,
by multiplying by the effective column diameter, : the outside diameter plus an
allowance for the thermal insulation and attachments such as pipes and ladder (Sinnott &
Towler, 2009). Based on Sinnott & Towler (2009), a wind speed of (100 mph) can
be used for preliminary design studies; equivalent to a wind pressure of . With
the effective column diameter of 2.052 m, the wind loading and bending moment of the
reactor was calculated to be 2626.71 N⁄m and 34200.61 Nm respectively.
5.6.5 Analysis of stresses
5.6.5.1 Analysis of stress of Ammonia Reactor Column

Analysis of different stresses on the reactor will be carried out. Table 5.6.5.1 below shows the
considered stresses analyzed in this mechanical design evaluation. All the calculation
involved in determining the stresses are shown in the Appendix C6.
Table 5.6.5.1: Summary of stress values

Stress Dimension Value (MPa)
Longitudinal Stress 45.81
Circumferential stress 91.62
Dead Weight Stress 0.645
Dead Weight Stress (Hydraulic Testing) 0.717
Bending Stress 0.1595
Critical Buckling Stress 1526
Normal Operation
Full Report
Stress Dimension Value (MPa)

Resultant longitudinal stress (upwind) 45.3245
Resultant longitudinal stress (downwind) 45.0055
Differences between the Principal Stresses (upwind) 46.2955
Differences between the Principal Stresses 46.6145
(downwind)
Maximum compressive stress 0.8045
Hydraulic Pressure Testing
Resultant longitudinal stress (upwind) 45.2525
Resultant longitudinal stress (downwind) 44.9335
Differences between the Principal Stresses (upwind) 46.3675
Differences between the Principal Stresses 46.6865
(downwind)
Maximum compressive stress 0.8765
Since all the stresses determined are below the maximum allowable stress of stainless steel
(103 MPa), thus this shown that the reactor will not fail even it reach the worst case scenario.
Both of the maximum compressive stress during normal operation (0.8045 MPa) or pressure
testing (0.8765 MPa) did not exceed the critical buckling test (1526 MPa), thus this reactor
will be strong enough to withstand the buckling.
5.6.5.2 Analysis of stress of skirt support

For a tall vessel, skirt supports are recommended as they do not impose concentrated loads
on the vessel shells; they are particularly suitable for use with tall columns subject to wind
loading. Based on Figure 13.23 from Sinnot & Towler (2009), for a vessel with diameter of
, the suggested skirt thickness, is . It is assumed that the height from the
bottom tangent line to ground to be considering the allowance for piping. The skirt
thickness should be such that under the worst combination of wind and dead weight loading
the following criteria are not exceeded:
where is the maximum allowable design stress for the skirt material SS304, normally
taken at ambient temperature, which is determined to be , is the welded-joint
efficiency assuming to be , is the young modulus which is and is
the base angle of a conical skirt, taking it to be . From the result obtained, both criteria
are bigger than the maximum resultant stress (tensile and compressive). Thus, a 10mm skirt
Full Report
thickness is enough to support the wind and dead weight loading of reactor. Lastly, a 2mm
of corrosion allowance will be added to the skirt thickness and gives a total skirt thickness of
12mm.
5.6.6 Sizing of pipe for the inlet and outlet

There are two inlet and two outlet stream on the ammonia synthesis reactor. The inlet stream
includes the reactor feed and inlet of cold gas into intercooler in the reactor. The outlet stream
will be the outlet of the intercooler and the product outlet stream. For the calculation for pipe
sizing, based on Section 5.6 of Chemical Engineering Design by Sinnot & Towler (2009), the
velocities of the streams are estimated based on the density of the streams. Similarly,
stainless steel 304 will be used as the construction material for the pipes due to the present of
ammonia and large amount of hydrogen component. The inner diameter of the pipe can be
calculated based on the volumetric flow and liquid velocity. After the pipe size has been
determined, a large pipe diameter was usually chosen based on the nominal diameter in the
ASME Pipe Schedules. The reason for choosing a larger pipe size is to account for the
flowing fluid fluctuation in flow rates.
Table 5.6.6.1: Pipe sizing based on ASME Standard dimensions
Description Calculated Calculated Standard ASME
Diameter Diameter (inch) Diameter (inch)
(mm)
Feed 165.685 6.523 8.0
Inlet to 174.108 8.0
intercooler 6.855
Outlet from 158.922 8.0
Intercooler 6.257
Product Outlet 164.600 6.480 8.0
Full Report
5.6.7 Specific Data Sheet and mechanical design drawing

Full Report
Ammonia Synthesis Reactor

Mechanical Design Specification Sheet
Alternis BioAmmonia 2-bed Ammonia Synthesis

Company Description
Pvt. Ltd. Reactor
To produce Ammonia across 2
Number Required = 1 Functions catalyst bed
Process Data
Fluid Properties Unit Feed Product
Mass Flow Rate 17500.736 17500.736
Volumetric Flow Rate 708.307 698.472
Temperature 340 427.463
Pressure 150 149.8
Density 24.708 25.057
Catalyst
Number of Bed 2
Catalyst Type AmoMax®-10RS
Particle Shape Spherical
Particle nominal diameter 2
Particle Density 6635.5
Bed 1 Bed 2
Bed Height 1.554 1.073
Bed Thickness
Bed Diameter m 1.2432 1.2432
Void Fraction 0.7739 0.7739
Equipment Properties
Number of Bed 2
Number of Intercooler 1
Design Pressure 16.5
Design Temperature 500
Pressure Drop 20
Type of Support Skirt
Column Properties
Design Orientation Vertical
Domed Head Type Hemispherical
Shell Material Stainless Steel 304
Shell Head Material Stainless Steel 304
Shell Inner Diameter 1.4
Shell Thickness 126.06
Column Height 6.751
Skirt Material Stainless Steel 304
Skirt Thickness 12
Intercooler Shell and Tube Heat Exchanger
Shell Tube
Full Report

Outer Diameter 0.8478 Outer Diameter 38
Thickness 6.4 Thickness 2.1
Baffle Cut 25 Total Length 1.83
Shell Side Velocity 6.8667 Number of Tubes 196
Tube Side Velocity 1.9648
Stress Analysis
Maximum Tensile Strength 103 Hydraulic Dead Weight Stress, 0.717
Normal Dead Weight, 389.84 Bending Stress, 0.1595
Hydraulic Dead Weight, 433.08 Critical Buckling Stress, 1526
2626.7 Maximum Compressive Stress,
Wind Loading, 0.8045
1
34200. Hydraulic Maximum Compressive
Bending moment, 0.8765
61 Stress,
Longitudinal stress, 45.81 45.3245
(upwind),
Circumferential stress, 91.62 45.0055
(downwind),
Hydraulic Resultant longitudinal
Dead Weight Stress, 0.645 45.2525
stress (upwind),
Hydraulic Resultant longitudinal
44.9335
stress (downwind),
Remarks
Prepared By Nisha Thavamoney Date 14th January 2014
Checked By Jenny Yap Wee Li Date 15th January 2014
Full Report
5.7 Detailed Process and Mechanical Design: Design of Vapour-Liquid

Separator (S-702)
The last ammonia product recovery process involves flashing the process streams to a lower pressure
and temperature within its dew and bubble point. The resultant mixture is then separated by a vapour-
liquid separator in which its detailed mechanical design will be performed. After being flashed, the
properties of the streams going into the vapour-liquid separator are listed out in Table 5.7.1.1 below.
Table 5.7.1.1 Properties of gas and liquid obtained
Overall Vapour Phase Liquid Phase

o
Temperature ( C) -34.15 -34.15 -34.15
Pressure (kPa) 305 305 305
Mass Flow (kg/hr.) 6539.183 1320.6 5218.583
ρ (kg/m3) 545.335 1.748 679.882
Volumetric flow, Q (m3/hr.) 2630.067 2621.603 8.464
Phase Fraction (molar basis) - 0.2763 0.7237
Since the separator to be designed is the last out of the three separators. Its performance is the
determining factor in meeting high purity product specification of >95% mass fraction of Ammonia.
As such, it should be fitted with appropriate internals and appropriate sizing to improve separation
efficiency.
The design methods and heuristics for this vapour-liquid separator are results of a review of literature
source and accepted industrial guidelines. The design methods are obtained from a combination of
Red-Bag (2013), Sinnott (2009) and Svreck WY & Monnery WD for vapour-separator dimensioning,
whereas Mechanical Design is done in accordance to Boiler and Pressure Vessel Code- ASME.
5.7.2.1 Orientation
The two-phase separator was decided to be orientated vertically as its advantages overweigh that of a
horizontally-oriented one. This is because it requires a smaller land plot area which is favourable in
the view of its economic implications as well as more efficient for high gas-liquid volume ratio which
applies to this case (Red-Bag, 2013). Furthermore, it performs better in handling liquid slugs since,
due to its shape, enough surge space is provided to ensure no liquid carry-over in the gas outlet
(Mulyandasari, 2011). This is crucial as the gas outlet is fed to a downstream compressor which
cannot tolerate any liquid inlet as it may damage it. Added to that, vertical-oriented vessel is less
sensitive to liquid level fluctuations since change in liquid volume per unit level change is only small
and thus this allows for a better level control (Guo et al. 2007). Since gases moves vertically, changes
in liquid level does not affect cross-sectional area of gas flow and thus liquid removal efficiency
remains constant with varying flow rate unlike with horizontal separator (Red Bag, 2013). However,
because of the natural up flow of gas against the falling droplets of liquid, a sufficiently large
diameter is required to slow the gas down to below velocity at which liquid droplet will settle out
(Guo et al. 2007). However, there should not need to concern about the diameter being too large, as
Full Report
ammonia liquid density much exceeds that of vapour density which is 679.882kg/m3 and 1.748kg/m3
respectively. As a result, terminal velocity of liquid ammonia droplets will be relatively high which
generally gives small vessel volume.
5.7.2.2 Vapor-Liquid Separation

In the vapour-liquid separator, vapour and liquid segregation is achieved through three consecutive
separation stages in which each is intended to separate decreasing liquid droplet sizes. The initial
separation occurs at the separator inlet aided by a diverter device to dissipate the energy and limit the
feed momentum in order to re-direct flow of large liquid droplets to the liquid accumulation section
(KW International, 2008). The next stage is through gravity separation for smaller liquid droplets as
the vapour flow upwards across the vessel diameter (Svrcek, WY & Monnery WD, 1993). The final
stage is elimination of smallest liquid droplet or ‘mist’ by a mist eliminator known as demister pad
which is typically a knitted wire mesh. It works in a way that, as the gas streams passes a torturous
path of the mesh, the contact surface of the demister become wetted with the entrained liquid.
Continuous contact will cause particle to build-up and coalesce into larger droplets, big enough to
overcome the drag force and drops to the liquid accumulation section. Added to that, vapour-liquid
separator has a liquid accumulation area to hold liquid attained from all three separation stages for a
short period of time. This retention time should be enough to allow degassing of dissolved gases in
the liquid. The design of the separator with regards to each of these sections is discussed below.
5.7.2.2.1 Primary Separation: Inlet Diverter

Internal inlet nozzle diameter is taken equal to that of the feed pipe. Its sizing is done from using the
optimum velocity method based on fluid density from (Sinnott, 2009). The optimum velocity is such
that it is below that at which erosion is likely to occur. Furthermore, it is ensured that the momentum
criteria as per guidelines provided from (Svrcek, WY & Monnery WD, 1993) is being conformed to.
Calculations are presented in Appendix section C.7. Furthermore, depending on material of
construction and their allowable stress at operating temperature, calculations are done to ensure that
pipe thickness is able to withstand operating pressure. The resultant inlet nozzle pipe specifications
being used based on ASTM A106M standard sizes are tabulated below:
Table 5.7.2.2.1 Inlet Nozzle Pipe Specifications
Material Nominal Pipe Outside Schedule Thickness Inner Diameter

Size (mm) Diameter (mm) Number (mm) (mm)
Carbon Steel 500 508 10 6.35 495.3
Pipe flanges temperature-pressure rating class is determined from ASME B16.5-2003 standard with
regards to the material Carbon Steel. Based on the rating, standard flange sizes for the pipe’s nominal
size are then obtained. The same approach is used for other nozzles as well.
5.7.2.2.2 Secondary Separation: Vessel Diameter

The vessel diameter is sized in a way to allow an efficient degree of separation in the secondary
separation section. The terminal velocity, UT, of the liquid droplet is obtained when a force balance on
the liquid droplet falling under gravity through a fluid, gives zero net force (Gravity – Buoyancy –
Drag = 0).
Full Report
This liquid droplets will settle as long as gas velocity, UV < UT. The upward velocity of gas velocity is
ensured to be low enough by ensuring a sufficiently large enough diameter. The K value with mist
eliminator in eq 6.1.1 is typically empirical since coalesce droplet diameter is hard to predict. In this
separator design, the K value with demister is obtained from (York Mist Eliminator, n.d.) in which its
K-values have been curve fitted and translated into an equation in terms of the operating pressure
(Svrcek, WY & Monnery WD, 1993).
The resultant minimum vapour disengagement diameter is found to be 0.6712m. Calculations are
presented in Appendix section C.7. This represents the mist eliminator diameter and thus, the inner
diameter of the vessel should be made larger to accommodate space for its support ring. Typically the
calculated value is taken up to the next six inch (Svrcek, WY & Monnery WD, 1993). However, a
safety margin to allow gas flow fluctuations are being considered in the design and the diameter is
thus taken to another following 6in increment. Thus, the resultant required diameter of the vessel is
taken at 36in which is 0.9144 m.
After having decided the vapour disengagement diameter, the disengagement height from the
centreline of the inlet nozzle to the demister pad follows the rules of thumb in which
Or a minimum of
Where dN is the inner diameter of the inlet nozzle
The latter was used, since the first equation is lesser than the minimum, HD calculated to be 0.8573m.
Calculations in Appendix section C7.3.
5.7.2.2.3. Tertiary Separation: Mist Eliminator

Since high liquid separation efficiency is required, a mist eliminator in the form of demister pad is
installed. The demister pad is constructed in the form of knitted wire mesh as it is cost effective,
versatile and provides a high efficiency separation to droplet sizes of 2 to 3μm (Sulzer Chemtech,
n.d.). The height is taken as 4in with an additional 2in for grid which totals to 0.1524m as this is the
typical standard size. This demister should be horizontally fixed and supported. At both the top and
bottom, dual support rings are used to mount mist eliminators whereby one of the rings at the top will
have a removable segment for mounting and demounting the pad. Since diameter of vessel is small,
one piece eliminators are used which is removable from the top. The demister can be accessed
through the vessel’s removable flanged head.
Top of demister pad to the top tangent line of the vessel is taken to be 1ft which is equivalent to
0.3048m as per recommended from (Svrcek, WY & Monnery WD, 1993).
Full Report
5.7.2.2.4 Liquid Accumulation: Liquid level

The retention time and, in turn, liquid level is dependent on the following parameters
Table 5.7.2.2.4.1 Liquid retention time parameters and their definition
Definition
Hold-up time Time taken to reduce liquid level from normal liquid level (NLL) to empty
/low liquid level (LLL) while maintaining a normal outlet flow without feed
make up. Hold-up time is based on the reserve required for good level control,
efficient degassing and safe operation of downstream facilities.
Surge time Time taken to rise from normal liquid level (NLL) to maximum/ high liquid
level (HLL) while maintaining a normal feed without the outlet flow. Surge
time is based on requirements to accumulate liquid as a result of upstream and
downstream variations or upsets.
Hold-up time and surge time are dependent on the type of the service of the vessel, in which for our
case, is a separator which feeds to a tank with a downstream pump. The hold-up and surge time
allocated for this type of service is shown in Table C7.5.1 in Appendix. Moreover, assuming trained
personnel and standard instrumented separator, a factor of 1.2 is used. As a result, after accounting
this factor, the required hold-up up time required is 6 minutes whereas surge time is 2.4 minutes. Each
respective liquid level height are calculated as per guidelines from (Svrcek, WY & Monnery WD,
1993) and shown in the Appendix section C7.6 and is summarized in the table below
Table 5.7.2.2.4.12Heights of the liquid levels
HLIN Height from centreline of the inlet nozzle to high 0.80m

liquid level
HS Height from high liquid level (HLL) to normal 0.52m
liquid level (NLL)
HH Height from normal liquid level (NLL) to low 1.30m
liquid level (LLL)
HLLL Height from low liquid level (LLL) to bottom 0.38m
tangent line (BTL). Allows space for positioning
level controllers
Total 3m
Full Report
5.7.2.3 Total vessel Height
This L/D ratio is within the common range 2-6 for separators (Couper et. al. 2012).
5.7.2.4 Outlet Nozzle Sizing

Calculations procedures are similar to Inlet Nozzle and are covered in Appendix Section C7.2 the
specifications of the separator outlet nozzles are tabulated in the following tables.
Table 5.7.2.4.1 Vapour Outlet Nozzle Pipe Specification

Carbon Steel 250 273.1 10 4.19 264.72
Table 5.7.2.4.2 Liquid Outlet Nozzle Pipe Specification

Carbon Steel 32 42.2 10 2.77 36.66
5.7.2.5 Vent and Drainage

(Red-bag, 2013) stated that the size of vents and drains on vessels as defined from Engineering Guide
for the Preparation of Engineering Flow Diagrams (BN-EG-UE208), para. 5.2.4 for our vessel volume
of
The size of vents and drains on vessel shall be;
Table 5.7.2.5.1 Vent and Drain Connection Sizes
Vessel Volume Drain Connection Vent Connection
Up to 17m3 2” or 50.8mm 2” or 50.8mm
Since of vessel is of low pressure at 305kPa, Sch10 thickness is used.

Full Report
5.7.2.6 Vessel Head

As the last separator is a small vessel of 36 in. or 0.9144m diameter, a removable flanged top head is
used, which serves to provide access to the vessel in which no man-way is needed. Since the vessel
operates at low pressure rating, ellipsoidal head will prove to be the most economical and reliable
closure to use. Flat ends was also considered due to its low cost and due to the low pressure rating,
however it is less structurally efficient (Sinnot & Towler, 2009) and thus the thought is discarded. A
standard ellipsoidal head with major and minor axis ratio of 2:1 is being used.
Once again to obtain flange dimensions, pressure-temperature flange rating with regards to the
material of construction is determined. A standard size for flange in particular for 36in (vessel
diameter) is then found. The dimensions shown in general drawing includes flange thickness, outer
diameter, bolt circle diameter, bolt hole diameter and bolt numbers. The ellipsoidal head at the bottom,
on the other hand, is welded to the body and the weld line to tangent line gives an additional height of
51mm.
Height of vessel = Height of cylindrical vessel + height of ellipsoidal head (top + bottom)
= 4.3 + (Di/4) + (Di/4)
= 4.3 + (0.9144/4) + (0.9144/4) = 4.757 m
Flange height welded bottom = 51mm
Removable Flange height = 120.9mm
Top to Bottom tangent height = 4.3 m + 51mm +120.9mm = 4.473 m
Ellipsoidal head height = top head + bottom head = 0.4572m
Total height of vessel + Ellipsoidal heads = 4.93 m
5.7.3.1 Material Selection
5.7.3.1.1 Vessel Material

Table 5.7.3.1.1 Material Selection for Vessel Body and Head
Material Carbon Steel ASTM A516

Lowest Service Temperature -45oC
Min Yield Strength (MPa) 221
Tensile Strength (MPa) 379-448
Maximum Allowable Stress (MPa) 95.147
The material selected in the construction of the vessel is carbon steel ASTM A516 Grade 55 with
specifications for pressure vessel plates for moderate to lower temperature service. Its lowest usual
Full Report
service temperature is at -45oC (Key to Metals, 2010). This choice is justified due to carbon steel’s
availability and their excellent performance in lower ambient temperature services as a result of their
excellent notch toughness (Macsteel, 2012). This is important since in the design of low-temperature
system, which applies to this design case (at -27oC), notch toughness ranks high in importance, since
it is likely that part of the structure will fail as a result of notch or other stress concentration (Key to
Metals, 2010). Added to that, carbon steel has excellent weldability and low coefficient of thermal
conductivity and has the big advantage of low initial cost.
On the other hand, its shortcomings lie mainly in their corrosion performance. There have been
incidences that ammonia in its anhydrous state has caused stress-corrosion cracking (SCC) on carbon
steel. The root cause of SCC is the presence of oxygen from air and residual stresses in the metal. As
a result, enforced safety measures should be incorporated in the design. For ammonia equipment that
has been opened, there should be application of nitrogen purging system as a start-up procedure as to
prevent air getting into the ammonia system. Only when all the air is out, the separator can start
operating. The system should be integrated with regular system analysis on oxygen content to assure
the absence of oxygen. On the other hand, to prevent SCC as a result of residual stress in the material,
post weld heat treatment must be performed to reduce residual stress during the carbon steel’s
fabrication (Fertilizer Europe, 2008). Carbon steel can also suffer irreversible damage due to
hydrogen cracking. In hydrogen rich environment, such as the case, hydrogen can diffuse into the
carbon steel and react with carbon to form methane. This loss of carbon results in a loss of mechanical
strength and the formation of cracks. However, this only happen under certain conditions of typically
above 350oC which is not the case and thus the separator is safe from this nature of corrosion
(Nitrogen+Syngas, 2011). Added to that, it was stated in (Committee of Stainless Steel Producers,
1978) that generally anhydrous liquid ammonia is considered to be non-corrosive to carbon steel and
all classes of stainless steel. Although stainless steel has a higher corrosion performance and is not
susceptible to SCC, it is much more expensive and it is still more economical to use carbon steel by
including a considerable corrosion allowance on the thickness of the column (Cheremisinoff &
Davletshin, 2010).
As for the outer surface of the separator column, corrosion is even less likely to happen because the
column was designed to have a thermal insulation of 0.116 m, whereby moisture content on the outer
surface of the column will be greatly reduce. On top of that, the unique attributes of polyurethane
foam also make it naturally resistant to corrosion and staining. Besides, the corrosion performance of
carbon steel column can be greatly improve by applying layers of painting to slow down the corrosion
process.
5.7.3.1.2 Skirt Material

Since vessel is small and comparatively light, the material used for skirt is Carbon Steel A285 with
low- and intermediate tensile strength.
Table 5.7.3.1.2.1 Material for Vessel Skirt
Material Carbon Steel A285

Maximum Allowable Stress 89 N/mm2
Young’s Modulus 200,00 N/mm2
Full Report
5.7.3.1.3 Insulation Material

The liquefaction of Ammonia to -34.15oC at atmospheric requires much higher thermal insulation
efficiency than usual industrial insulations and apparently polyurethane foam has been the insulation
material used typically in this area of field. This is because it offers advantages such as having low
thermal conductivity which allows efficient retention of heat flow and also a good balance between
weight and mechanical strength. Due to these qualities, it can be used in applications that require
combination of insulation with load-bearing, impact resistance, weight and space saving, as well as
providing ease of installation and maintenance (DUNA-Corradini, 2011). Moreover, polyurethane
foam has the highest R-value (6.5 per in.) of all readily available, cost effective insulations available
in the market today (DWYER’s Foam System, 2004). Furthermore, the closed cells structure of rigid
polyurethane foams minimize water absorption which prevents corrosion on outer metal surface and
ice formation for very low temperature application. Added to that, it has a near-zero air permeability,
which is relevant in the prevention of SCC on separator’s material. On the other hand, its drawback is
that polyurethane foams are combustible and non-fireproof as they are organic compounds. However,
their ignitability and rate of burning can be made to improve in order to suit a variety of insulating
application.
Table 5.7.3.1.3.1 Material for Vessel Insulation
Material Polyurethane Foam

Thermal Conductivity 0.023 W/m2.K
Density 200,00 N/mm2
5.7.3.2 Load and Stress Analysis

Detailed mechanical design calculations for vapour-liquid separator vessel can be found in Appendix
C7.7-C7.12. Two loadings situations were analysed:
1. Normal Operation Loading

2. Hydrostatic Test Loading
The summarized calculations’ results are presented in Table 5.7.3.2.1 below;
Table 5.7.3.2.1 Mechanical design values for new thickness
New thickness 8 mm
Corrosion Allowance 4mm
Total Thickness 12mm
Di 0.9144m
Do 0.9384m
Loads on Column
(1) Dead Weight Load
Vessel Shell 15.10 kN
Top & Bottom Ellipsoidal Head 1.248 kN
Demister 0.184 kN
Insulation 1.01 kN
Full Report
Total 17.546 kN
(2) Hydrotesting Load
Water 30.69 kN
Total 48.237 kN
(3) Wind Load 11.420 kN
Primary Stresses
Design Pressure 0.477 MPa
Hydrostatic Test Pressure 1.241 MPa
(1) Pressure Stresses
Normal Operation
Hoop Stress, σh 27.282 MPa
Longitudinal Stress, σl 13.641 MPa
Hydrotesting
Hoop Stress, σh 70.932 MPa
Longitudinal Stress, σl 35.466 MPa
(2) Weight Stress, σw
Normal Operation 0.757 MPa
Hydrotesting 2.081 MPa
(3) Bending Stress, σb 5.013 MPa
(4) Net Longitudinal Stress, σz = (σl -σw σb)
Normal Operation
Upwind 17.897 MPa
Downwind 7.871 MPa
Hydrotesting
Upwind 38.398 MPa
Downwind 28.373 MPa
Maximum Stress Intensity
Normal Operation
(σ1-σ2) 9.385 MPa
Upwind (σ1-σ3) 27.520 MPa
(σ2-σ3) 18.135 MPa
(σ1-σ2) 19.411 MPa
Downwind (σ1-σ3) 27.520 MPa
(σ2-σ3) 8.110 MPa
Hydrotesting
(σ1-σ2) 32.534 MPa
Upwind (σ1-σ3) 71.553 MPa
(σ2-σ3) 39.019 MPa
(σ1-σ2) 42.560 MPa
Downwind (σ1-σ3) 71.553 MPa
(σ2-σ3) 28.993 MPa
Full Report
In determining the thickness, the principal stress (σ1-σ3) during Hydro testing becomes the decisive
variable, since it gives the highest stress values amongst the rest. The thickness is increased until (σ1-
σ3) < 95.8MPa (maximum allowable stress for Carbon Steel ASTM A516) and Hydro testing is
satisfied. Furthermore buckling check was done to ensure the design will stand under vacuum
conditions when it is subjected to external pressure. For both normal operations and Hydro testing,
maximum compressive stress is well below the critical buckling stress of 172.043 MPa. Hence, the
design won’t buckle under vacuum conditions.
5.7.3.3 Vessel Support

A straight cylindrical skirt support (ϴ = 90oC) was selected due to its suitability for use in vertical
vessels, as they do not impose concentrated loads on the vessel shell (Sinnott & Towler, 2009).
Openings are provided in the skirt for access and for any connecting pipes in the event of inspection
or maintenance. The skirt is welded flush to the shell as this type of welding is usually preferred. The
loads carried by the skirt are transmitted to the foundation slab. Skirt thickness mechanical design
calculations are found in section C7.13 of the Appendix and are summarized in the table below.
Table 5.7.3.3.1 Skirt Mechanical Design Calculations
thickness 2mm
Corrosion Allowance 4mm
Total Thickness 6mm
inner diameter of skirt, Ds = Di 0.9144m
DSO 0.9264m
Skirt height 1.5 m
Stresses on Skirt
(1) Bending Stress, σb 30.52 MPa
(2) Dead Weight Stress, σws
Normal Operation 3.047 MPa
Hydro testing 8.377 MPa
(2) Resultant stress in the skirt

Maximum σS (tensile) 27.471 MPa
Maximum σS (compressive) 38.896 MPa
Design Criteria
> Maximum σS (tensile) 75.65 MPa > 27.471 MPa
> Maximum σS (compressive) 54.68 MPa > 38.896 MPa
Full Report
5.7.4 Summarize of the detailed design specification of Vapor Liquid Separator (S-702)
5.7.4.1 Mechanical Design

Full Report
EQUIPMENT SPECIFICATION SHEET

Vapour Liquid Separator
Company Alternis BioAmmonia Pvt. Ltd. Description Vapour-Liquid Separator (Flash

Vessel)
Equipment No. S-702 Sheet No.
Number Required = 1 L/D 1171mm/ 4473mm
VESSEL DETAILS
VESSEL CLASSIFICATION Division 1
DESIGN CODE Boiler and Pressure Vessel Code – ASME
LOCATION OUTDOOR
HAZARDOUS CONDITIONS Y
PRESSURE CYCLIC N
TEMPERATURE CYCLIC N
FLUID CONTAINED H2O, H2, N2, CH4, NH3
TOTAL VAPOUR PRODUCT FLOW kg/hr. 2294.345
TOTAL LIQUID PRODUCT FLOW kg/hr. 4244.84
DESIGN PRESSURE kPag 477.4
NORMAL OPERATING PRESSURE kPag 305
RELIEF VALVE SET PRESSURE kPag 477.4
HYDROSTATIC TEST PRESSURE kPag 1861.9
O
DESIGN TEMPERATURE C -38
O
NORMAL OP. TEMPERATURE C -34.15
DEAD WEIGHT kN 17.546
HYDROSTATIC TEST WEIGHT kN 48.237
CONSTRUCTION AND MATERIALS
SHELL MATERIALS ASTM A 516 Grade 55
SHELL INTERNAL DIAMETER mm 914.4
SHELL LENGTH ( CYLINDRICAL ) mm 4300
SHELL TAN-TAN LENGTH (+ FLANGE) mm 4473
CORROSION ALLOWANCE mm 4
INSULATION mm
INSULATION MATERIAL Polyurethane Foam
PAINT Epoxy-based
STRESS REDUCTION YES
JOINT EFFICIENCY % 100%
COLUMN INTERNALS
INLET DEVICE INLET DIVERTER
DEMISTER PAD TYPE KNITTED WIRE MESH
3
MESH DENSITY kg/m 192
MESH SURFACE AREA m2/m3 650
VOIDAGE % 97.5
NOMINAL MICRON RATING 3μ
PAD DIAMETER mm 934.4
GRID DIAMETER mm 864.4
GRID THICKNESS mm 35.4
GRID CLEARANCE mm 25
PRESSURE DROP ACROSS DEMISTER kPa 100
DISSHED END SUPPORT
TOP END BOTTOM END SKIRT MATERIAL ASTM A 285 Gr A
TYPE ELLIPSOIDAL ELLIPSOIDAL SKIRT THICKNESS mm 6
MATERIAL ASTM A516 Grade 55 A516 Grade 55 SKIRT HEIGHT mm 1500
THICKNESS mm 7 7 INNER DIAMETER mm 914.4
CORROSION ALLOWANCE mm 4 4 NO. OF ACCESS HOLES 1
HEIGHT mm 239.6 239.6 HOLE DIAMETER mm 850mm
INNER DIAMETER mm 914.4 91.4 JOINT EFFICIENCY % 85
JOIN EFFICIENCY % 100 100
NOZZLE SCHEDULE
FLANGE FLANGE WIND LOADING
REF. NO SIZE SERVICE REF. NO SIZE SERVICE
RATING RATING
A1 1 250 Feed 150 L3,L4 2 40 Switch 150 DESIGN WIND SPEED km/hr. 180
B1/B2 2 250/25 Outlet 150 D1 1 50 Drainage 150 WIND LOAD kN/m 2.66
V1 1 50 Vent 150 I1 1 - Inspection -
L1,L2 2 40 Transmit 150 P1 1 40 Gauge 150
STREAM DETAILS
STREAM NUMBER 140 143 141
DESCRIPTION VAPOUR-LIQUID MIXTURE VAPOUR OUT LIQUID OUT
O
TEMPERATURE C -34.15 -23.16 -23.16
PRESSURE kPag 224.4 124.4 124.4
VOL FLOWRATE m3/hr. 2630.067 2621.603 8.464
Remarks
PREPARED BY Fatimah Azizah Riyadi DATE 23/1/14
CHECKED BY Lee Leong Hwee DATE 23/1/14
Full Report
CHAPTER 6 | PIPING AND INSTRUMENTATION DIAGRAM (P&ID)
6.1 Piping & Instrumentation Diagram for Post-Treatment of Syngas

Section
6.1.1 P&ID Flow Sheet
Full Report
6.1.2 Brief Description of Flow Sheet
6.1.2.1 Start-Up Process

During the startup process, the globe valve at the suction side of the pump, GLV106 and the
bypass valve at the at the discharge line BV101 is fully open. At the same time the discharge
valve, GV112 is fully closed to prevent the water flow reach the waste heat boiler at the
moment. The driver is switched on. Since the temperatures of the pumped fluid will exceed
93°C, the pump is warmed up prior to pump start-up. A small amount of the boiler feed water
is circulated through the pump until the casing temperature is within 38°C of the water
temperature prior to pump start-up to avoid thermal shock to the liner and impeller and
prevent damage of mechanical seal of the pump. The discharge globe valve is then open
slowly till the pump reaches the desired flow rate capacity which is 8 m3/h. The pressure
indicator is checked to ensure the pump quickly reaches the correct discharge pressure. The
shell side vent of the economizer, DV-107 is fully open and the GLV109 is partially open to
slowly fill the shell side in order remove the air present in the shell. When the shell is full, the
shell vent-valve is shut and the GLV109 is open and the shell-side outlet is open as well. As
for the syngas inlet at the superheater, the tube side vent should be open and the tube side
inlet valve, GV107 should be partially open to fill the tube-side slowly with the syngas and
remove any foreign gases present. The tube side vent valve is closed once the tube side is
filled with syngas and the inlet valve is fully open. Same method applies to the evaporator,
WHB-102 as well where the cold fluid, steam should enter first then followed by the hot
syngas.
6.1.2.2 Normal Operation

During normal operation, the boiler feed water from the storage tank is passes through the Y-
Strainer before entering the multistage pump whereby it get pumped up to 50 bar. The
pressure indicator, PI-108 is placed at the discharge side of the pump to monitor the pressure
of the water. The check valve (CHV107) is provided on the discharge side of the pump to
avoid reverse flow of the water when the pump is not in operation and downstream to the
check valve, the bypass valve (BV101) is provided as recirculation line in case the upstream
control valve is closed. The temperature of the syngas outlet from the economizer is
controlled to the temperature that it need to be cooled down before entering the scrubber and
reformer by adjusting the flow-rate of the water entering the economizer using the flow
control valve, CV-102. The flow control valve is set to be a fail open valve as during the loss
of power, the control valve is open to make sure that the flow of the compressed water to the
Full Report
waste heat boiler to cool down the syngas or else it might result in high temperature of the
syngas flowing out of the waste heat boiler disrupting the downstream process. Since the
temperature and pressure of the streams in this section are very crucial properties that need to
be monitored continuously, the pressure transmitter and indicator as well as the temperature
transmitter and indicator is displayed with the operator access to adjustment.
6.1.2.3 Abnormal Operations, Emergency Shutdowns and Maintenance

Pressure Relief valves are installed for the kettle evaporator and the superheater whereby it
will be open when there is a constant increment of pressure build-up in the shell side to vent
out the excess pressure out of the system. A Steam trap is installed at the saturated steam
stream leaving the kettle evaporator and the superheated steam stream leaving the superheater
in order to trap and remove steam condensate with the minimum loss of steam itself. The
pressure of the syngas stream leaving the cyclone need to be monitored carefully as it hugely
depends on the performance of the cyclone and it will affect the heat transfer mechanism in
the waste heat boiler if the pressure is greater than the design pressure or very low than the
design pressure. All control valves have bypasses to ensure operation can proceed if any
equipment required immediate repair work.
Full Report
6.2 P&ID (Autothermal Reformer, Syngas and Air Compression)

Full Report
6.2.2 Brief Description of Flow Sheet
6.2.2.1 Start-up Process

During the start-up, the system within the reformer is operating but the steam which is the
one of the main feeds of autothermal reformer is ready. At this situation, fuel and oxidant
which is air will be used to perform partial oxidation in the reformer.
Possible scenario for Autothermal Reformer Start-up:
1. Start-up of the plant after the annual plant shut-down

2. Start-up procedure after emergency shut down
Before start-up, the operators should ensure that all the closed valves and inlets entering the
autothermal reformer in the P&ID are closed. At this stage, the feed will be allowed to
undergo compression and preheat to the desired temperature. The start-up of autothermal
reformer is carried out by heating the reformer with natural gas or inert to a temperature
between 110°C and above. Once the autothermal reformer is above the boiling point of water
at the operating pressure, the inlet for reformer will be opened. Once the partial oxidation in
the combustion zone is established, air inlet is introduced incrementally up to the desired
flowrate. Therefore, the entire system will achieve the desired operating condition. CV-222
will open to purge the extra natural gas or inert used in the start-up process. It will be stored
in the storage tank and discharge.

At normal operation, Syngas will be compressed to the desired pressure and preheated in the
fired heater. The fired heater is operated by using fuel gas and air as feedstock into the burner.
The closed globe valve BV-202 and BV-203 will be used during the maintenance of CV-202
and CV-203. The proportion of fuel gas and air entering the fired heater is adjusted using
flow ratio controller FFC-201. Flow of syngas and air will be transmitted by FT-202 and FT-
203. The ratio is set and controlled by CV-203. A by-pass with GV-209 is used for safety
purposes of using fuel gas in the fired heater. TI-201 is to indicate the inlet temperature of
syngas. The flow of the fuel gas will be controlled by CV-202 in order to achieve the desired
preheated temperature. Preheated syngas will be mixing with the steam. However, the
proportion of syngas and steam entering the reformer are adjusted using flow ratio controller.
The amount of steam required will be controlled by CV-204 and mix with syngas at MV-201
then enters the reformer. However for air, it will be preheated by using the fuel gas from fired
heater through a heat exchanger. The fuel gas from the heat exchanger will be sent to stack.
Full Report
Check valves CHV-202 and CHV-203 are to prevent the feed from flowing backward. The
pressure drop in the catalytic bed in equilibrium reactor will be indicated in PDI-207. The
outlet temperature of syngas is indicated in TI-209 and the conversion of the methane is
checked and indicated in AI-209.
6.2.2.3 Abnormal Operatiom, Emergency Shutdown and Maintenance

Since autothermal reformer is operated under high temperature and pressure. Precautions and
concerns need to be emphasized on the temperature and pressure of the reactor. The pressure
of the reformer will be shown in control room. High and low alarm (PAH-206 and PAL-207)
will be stimulated if the reformer is operated at abnormal pressure to notify the operators and
therefore pragmatic approaches can be taken. It is known that, the temperature of the
reformer is mainly affected by the air entering the reformer. Therefore, the temperature of the
reformer will be measured and send through TIC-208. Then, the flow of air entering the
reformer can be adjusted by CV-201. Besides, that high temperature and low temperature
alarm (TAH-208 and TAL-208) will be stimulated and notify the operators if the reformer is
operating at abnormal operating temperature.
Pressure relief valve is attached at the top of the reformer. Pressure relief valve (PRV-201) is
used when there is pressure build up in the reformer. GV-219 and GV-220 is located between
PRV-202. These two valves will only be opened during the maintenance of PRV201.
Shut down process is carried out by terminating the air flow entering the reforme followed by
terminating the steam flow. After that, the reformer will be purge with fuel,natural gas or
nitrogen. The reformer will then be allowed to cool down to approximately 50°C.
Full Report
6.3 P&ID (Water-Gas Shift Reactors)

6.3.1 P&ID Diagram Flow Sheet with Legend
Full Report
6.3.2 Piping and Instrumentation Diagram (P&ID) Explanation
The design of the P&ID is done based on the optimization of the carbon monoxide shift system and
the operating conditions of each equipment were controlled by flow control, pressure control,
temperature control and ultimately cascade control. The safety aspects during the operation of the
plant are also considered by including pressure relief systems, alarms, safety trips and interlocks.
To begin, the start-up of the plant is required as the cold catalyst bed should be warmed to a
temperature above due point before the process gas is introduced to both the high temperature water-
gas shift reactor (R-301) and low temperature water-gas shift reactor (R-302). Hot inert nitrogen from
the fire heater is used in this case in order avoid temperature peak which normally occurs when the
process gas comes into contact with the catalyst bed for the first time. The inert nitrogen will then be
vent out at the R-302reactor exit. Eventually, the inert gas is gradually replaced by the process gas
while the inlet and exit valves on the process gas pipelines are fully opened with a slow closing of the
vent in order to commission the reactor.
A temperature transmitter is located at the process gas inlet of R-301 is used to monitor the
temperature entering the reactor while the chromatograph carbon monoxide and hydrogen analyzers
are used to monitor the inlet conditions which is the steam to carbon monoxide (CO) ratio that will
affect the conversion of CO and the purity of hydrogen produced. Since R-301 is an isothermal
reactor, the cooling water flow control is essential in order to maintain the reactor temperature at
350 . In this case, a cascade control is used whereby the output of Temperature Indicator Control
303 is used to adjust the set point of the Flow Indicator Controller 303. This control will give a
smoother control in situations where direct control of the variable would give rise to unstable
operation. In terms of safety, this control also ensures a lower possibility of runaway temperature.
Again, the compositions of process gas at the reactor outlet are determined. However, this time,
chromatograph analyzers for carbon monoxide and hydrogen are used as the operators will be more
interested in knowing the conversion of CO and the amount of hydrogen produced as this will also
indicate that it is time to replace the deactivated catalysts in the reactor. Similar controls are used for
the low temperature water gas shift reactors (R-302) with the same reasoning.
As the process gas exits R-301 and enters a heat exchanger (HX-301) with cooling water on the shell
side to cool the process gas from 350 to 200 in order to meet the operating condition of R-302.
Temperature transmitter 304 and temperature indicator control 304 are used to regulate the
temperature of the heat exchanger whereby if the reactor temperature is too hot, the flow control valve
304 will open allowing more cooling water to flow into the heat exchanger. On the contrary, if the
temperature of the process gas is already cooled down, the opening of flow valve 304 will be closed
or restricted allowing less cooling water to cool the process gas.
Pressure relief valves are installed on both the high temperature water-gas shift reactor (R-301) and
low temperature water-gas shift reactor (R-302). This is done to make sure that the excessive pressure
in the both the reactors can be vented to the flare header as excessive pressure build-up in the reactors
can be very dangerous and can lead to major accident. It was decided that one pressure relief valve
for each reactor is sufficient as spare pressure relief valves are normally ready for installation and use
whenever there is a pressure relief valve malfunction. Besides that, a by-pass valve is installed in case
of a faulty pressure relief valve whereby, the pressurized process gas can still be vent to the flare
header during emergency when that faulty valve is replaced by a new valve.
Full Report
On top of that, alarms are used as one of the safety features in this plant as they will warn the
operators and workers of serious and potentially hazardous, deviation in process conditions. Software
alarms with shared displace device are used to alert when the pressure difference across the reactor is
too high (more than 3 bar). The other alarm such as the Temperature High High Priority One alarm
which is only triggered during the critical stage will immediately alert the operator to trigger an
emergency shut-down using a hand switch causing the venting valves located at the outlet pipelines of
R-301 and R-302 to open due to the deactivation of solenoid valves. All the process gas in the reactors
will immediately be vent to the flare header.
During the shut-down of the plant, inert nitrogen which by-passes the fired heater is used in order to
purge out all the process gases without damaging the catalysts as it will avoid condensation of water
on to the catalyst.
Full Report
6.4 Piping and Instrumentation Diagram (P&ID): Carbon Dioxide Removal

Section
6.4.1 P&ID Flow Sheet for Carbon Dioxide Removal Section

Full Report
6.4.2 Brief Description of P&ID Flow Sheet for Carbon Dioxide Removal Section
The equipments that are included in the P&ID flow sheet are listed as follow:
 Major Equipment – Absorption Column (AC-401)

 Minor Equipments – Shell and Tube Heat Exchanger (HX-402)
– Pump (P-401)
 Intermediate Equipment – Storage Tank (T-401)
6.4.2.1 Start up
During startup process, all the equipments and pipelines are purge with inert gas to remove
all the impurities that might cause contamination to the process stream. After that, the gas is
vented to relief venting header through valve GV-437. Check valve (SCV-401) is installed on
the purging pipeline to ensure that the process fluid does not flow into the pipeline and
contaminate the inert gas. After that, the absorption column is loaded with the required
amount of amine solvent prior to the entry of the syngas so that maximum CO2 absorption
and removal can be achieved from the beginning of the process.

In normal operation of the process, the CO2 removal section is monitored and controlled by
various control schemes suggested. On the absorption column (AC-401), a temperature
indicator transmitter is installed to monitor the temperature of the column. If the temperature
goes too high, signal will be sent to the high temperature switch and the high temperature
alarm will be stimulated to alert the operators. Moving on, a pressure indicator transmitter is
also installed on the absorption column to ensure that the pressure is well within the column
operating pressure range. In case the pressure rises above the maximum operating pressure of
the column, a signal will be sent to the high pressure switch to stimulate the high pressure so
that the operators are aware. Similarly, signal will be sent to the low pressure switch if the
pressure of the column drops below the minimum operating pressure, and the low pressure
will be stimulated. Pressure relief valve is also instilled on the absorption column. Pressure
relief valve is mainly used for prevention of pressure build up in the column. The relief valve
is set to open when the pressure in the column is too high until it hits the predetermined set
pressure. Besides that, pressure differential control scheme is installed across the packed bed
Full Report
of the absorption column to measure and monitor the pressure drop across the packed bed.
The efficiency of CO2 absorption decreases as the pressure drop across the packed bed
increases, thus high pressure drop is undesired. The high pressure drop might be due to
clumping of the packings or damaged packings. Signal will be send by the pressure
differential transmitter to the high pressure differential switch to stimulate the high pressure
differential alarm. With this, operators will be alarmed and maintenance works can be carried
out as soon as possible.
Level of liquid solvent in the absorption column is monitored by adjusting the flow rate rich
amine (stream 403) outlet. The liquid level in the column is displayed on the level indicator
and the signal is sent by the level transmitter to the level control so that the flow rate of rich
amine exiting the column will be adjusted accordingly by the control valve (CV-401). If the
liquid level in the column is too high, opening of the control valve will increase to promote
the outlet liquid flow. Contrary, the opening of the control valve will decrease to restrict the
outlet liquid flow of the liquid level in the column is too low. Moving on, Stream 401 is the
sour gas (syngas) stream existing from the separator (S-401) where the water content in the
stream is reduced to minimal in the separator. A pressure indicator is placed on the stream to
monitor the pressure to make sure that the stream is entering the absorber is within the
operating pressure condition. As for Stream 402, the sweet gas (treated syngas) stream
exiting the absorption column, a pressure indicator is installed to monitor the pressure in the
pipeline. A CO2 analyzer is also installed on Stream 402 to monitor the concentration of CO2
as the maximum amount of CO2 that can present in the stream is 100ppm. The operators will
be notified by the alarm if the concentration of CO2 in the stream is high. CO2 analyzer is
commercially available on the market.
On the other hand, for the shell and tube heat exchanger (HX-402), it is used to further cool
the lean amine stream exiting from the stripping column to the operating temperature of the
absorption column. The temperature of the lean amine stream (Stream 411) exiting the heat
exchanger is controlled by adjusting the flow rate of the cooling water. A temperature
indicator transmitter is installed to display and transmit the signal to the temperature
controller so that the flow rate of the cooling water can be adjusted accordingly by the control
valve (CV-402). If the temperature of the lean amine stream is too high, the control valve
opening will increase to allow higher flow of cooling water to cool the lean amine stream.
Oppositely, if the temperature of the lean amine stream is too low, the control valve opening
Full Report
will decrease to restrict the flow of cooling water so that the lean amine temperature can be
maintained at the column operating temperature range. On top of that, every inlet and outlet
streams of the heat exchanger are installed with a gate valve so that the heat exchanger can be
removed for maintenance purposes. A check valve (SCV-402) is installed on the stream
entering the absorption column (Stream 411) to ensure one direction flow and to prevent back
flow of the lean amine. It is also used to avoid the syngas from mixing into the lean amine
stream.
An intermediate storage tank is included in the system. The storage tank is used to store the
lean amine temporarily if there are too much of liquid amine flow out from the absorption
column to the stripping column and to ease the adjustment of flow rate of the lean amine.
Besides that, fresh amine solvent will be made up into the storage tank in case there are any
lost of amine solvent in the system. A sampling valve (GV-438) is installed at the bottom of
the storage tank so that the liquid solvent in the tank can be taken for further analysis. If the
amount of amine solvent is lesser than the required amount, the fresh make up amine solvent
will be added into the tank by adjusting the manual valve (MV-401) manually.
As for the pump (P-401), a pressure indicator is installed after the pump to monitor the
pressure of the stream and to make sure that the pressure is at the required value. A flow
control scheme is installed across the pump to regulate the flow of the lean amine from the
storage tank and into the heat exchanger. Flow indicator transmitter is installed after the
pump to display and measure the flow rate of lean amine exiting from the pump. Signal will
be sent by the transmitter to the flow controller positioned before the pump so that the flow
rate can be adjusted accordingly by the control valve (CV-403). If the flow exiting the pump
is too high, the opening of the control valve will decrease to restrict the flow, and the opening
of the control valve will increase to promote the flow if the flow exiting the pump is too low.
A check valve (SCV-403) is installed to ensure one direction flow and to prevent back flow
of the lean amine stream.
Two pressure relief valves and two pumps are installed in the system. One is used during the
normal process and the other one is used as backup. On top of that, the advantage of
installing two pumps and pressure relief valves is that the process can still be carried on by
using the backup equipment while the first equipment is undergoing inspection and
maintenance. For every control valve installed in the system, 2 isolation valves and a bypass
valve are added for maintenance purposes.
Full Report
6.4.2.3 Shut Down and Maintenance

For shut down process, the liquid in the column, heat exchanger, pump, storage tank and also
pipelines is drained by allowing the liquid to flow through the drain valve. The liquid will be
sent to waste water treatment plant before discharging into the surrounding. The syngas or
vapour in the column is vented by opening the valve GV-437 that connects the pipeline to the
relief venting header. As for maintenance process, the backup equipment for pressure relief
valve and pump is used so that the process can be continued. The isolation valves and bypass
valve come into place when maintenance works are needed to be carried out on the control
valves. Isolation valves are used to stop the process fluid from flowing out from the pipelines
while maintenance is happening and bypass valve is in place for the process fluid to flow
continuously without passing through the control valve.
Full Report
6.5 Piping and Instrumentation Diagram (P&ID): Methanation Section
6.5.1 P&ID Flow Sheet for Methanation Section

Full Report
6.5.2 Brief Description of P&ID Flow Sheet for Methanation Section
6.4.2.1 Overall Process Description

Syngas from CO2 removal section’s absorber will first be heated by HX-501 and then enter
the reactor (R-501). The methanation product will then be cooled by HX-501. This part
involves heat integration. The water in cooled methanation product will then be condensed in
HX-502 using refrigerant.
6.5.2.2 Start-up
All the equipment will first be purged with N2 gas to remove all the unwanted air in the
equipment. This is important especially for the fixed bed reactor as air will poison the catalyst
that will be used in the process. At this stage, GV-501 and GV-518 will be open while GV-19
will be close. Nitrogen gas will pass through HEX-501, V-501, R-501 and HX-502 and
finally leave by being sent to flare. A check valve (SCV-501) was also placed to ensure that
there is no backflow of gases during the process.
At the starting point of the process, an electric heater (HT-501) will be used to heat up the
first stream of syngas to the desired temperature as there will not be any hot gas present in the
heat exchanger. GV-502 and GV-505 will be open while GV-503 will be closed to heat up
the syngas at this point. A temperature transmitter and indicator will be place at stream 504 to
display the heated temperature of that stream.
6.5.2.3 Normal operation

A composition analyzer will be placed at stream 502 to monitor the composition of carbon
oxides. There will be a steam drum (V-501) after the heat exchanger (HX-501) and its
purpose is to store syngas and to prevent syngas build-up during the operation. Since the
reactor was designed based on a specific amount of catalyst, the reactor will not be able to
support excessive carbon oxides as it will lead to low conversion reactions. Thus, the flow
rate leaving the steam drum will be controlled by CV-506 to prevent overheating and
overloading in the reactor.
Steam trap will be placed at the outlet of HX-501 to improve heat transfer process in the heat
exchanger by prolonging the heat transfer time of the process fluids between the shell and
tube side.
Full Report
A pressure indicator transmitter will also be placed at the bottom of the reactor. The switch
will be turned on and alarm will sound whenever the pressure is too high or too low to alert
the operators.
Pressure relief valves are also installed at the top of the reactor to regulate the pressure of the
reactor to ensure that the pressure of the reactor is not too high.
The desired temperature of syngas leaving HX-502 is set to be 5 . Such low temperature is
to ensure that water is condensed almost completely and ready to be separated in the
separator in the next section. In order to achieve this temperature, a temperature transmitter is
installed to send signal to the temperature control valve to regulate the flow rate of incoming
refrigerant.
6.5.2.4 Shutdown and maintenance

For shut down system of the reactor, valves GV-501 and GV-518 will be open while GV-19
will be close. Nitrogen gas will be fed to purge out all the contents in the equipment.
Operators can also access the reactor through the manway side by using the ladder. For every
control valve in the system, there will be isolated valves in case of any repair or maintenance
conducted without the need of shutting down the whole plant.
Full Report
6.6 Piping & Instrumentation Diagram of Ammonia Synthesis Reactor

Section
Full Report
6.6.2 Brief Description of P&ID Flow Sheet

The piping and instrumentation diagram shows the detailed arrangement of
equipment’s, instruments, piping, valves, and fittings in the real plant. Level control, flow
control, temperature and pressure control are some common control system that exists in a
plant. To start up an ammonia reactor, hydrogen and nitrogen component from methanator
are needed where required to be in a proportion of 3 to 1 respectively. From ammonia
synthesis section, the main feed into the reactor will be feed gas from methanator unit and
recycle (unreacted H) from ammonia refrigeration cycle.
Feed gas from methanator unit will be compress to 150 bar using a two-stage
compressor (K-602) with inter-stage cooler. The reason for using a multistage is mainly due
to high compression ratio (>5). A pressure control is needed to control the exit pressure and
at the same time, temperature of the stream can be control as well. A pressure indicator will
be place after K-602 to indicate the discharge pressure. When the detected discharged
pressure is deviates from 150 bar, signals will be sent to a pressure indicator control room
(PIC) where it further sent the signal to an electric motor (EM-602). An electric motor will
control the shaft of the compressor to achieve the desired discharge pressure.
Before the recycle stream can be feed into the first mixer (MX-601) for further
operation, a flow ratio control loop between the recycle stream and purge stream will be
control. Flow transmitter will be place at both recycle and purge stream and electric signal
will be send to the feed-forward controller (FFC) and thus, signal will be again send to the
control valve (CV601) to adjust and maintain the ratio throughout the process. A control
valve is always follow by a set of valve system which include two gate valve (GV), two drain
valve (DV), a reducer, an expander and a globe valve (GLV) as bypass for maintenance
purposes. In fact, control valves are usually selected in a size which is smaller compare to the
adjacent piping for economic reasons. However, sizing of pipe is still subject to the typical
selection process with respect to gas mass flow rate and density. With a smaller size of
control valve, a reducer or expander is usually needed to mate the control valve with the
piping and problems can occur with this configuration because the velocity related turbulence
generated by the expander at the control valve outlet creates its own noise. Normally this
noise will be noisier than the noise of abatement trim.
Mixer (MX601) will mixed up the feed gas containing hydrogen and nitrogen and the
recycle stream and again further divided into 2 streams. First stream will be used as cold gas
to cool the catalyst bed in ammonia synthesis reactor (R601) and the last stream will be used
as a feed for the reactor. Ammonia synthesis reaction is an exothermic reaction and thus heat
Full Report
will be release during the reaction. The feed gas temperature will increase when it pass
through the catalyst bed and therefore the cold gas that mentioned earlier will be used to
bring down the temperature before the reacted gas go into second catalyst bed. In order to
make sure the temperature at each bed does not exceed the optimum temperature (500 ),
temperature controller will be place at the first catalyst bed of the reactor. A temperature
control system will be placed at the outlet stream of the intercoolers. When the temperature
indicated that the temperature of the catalyst bed about to exceed the operating temperature,
control valve (CV602) will control the cold gas to stay for a longer time in the intercooler for
further cooling. Similarly, a set of valves will be installed together with the control valve.
Since R601 is dealing with high operating pressure, a pressure control will be used for
the reactor. When pressure in the reactor is about to exceed the operating pressure, a pressure
transmitter will transmit the electric signal to the pressure indicator alarm and the high alarm
will be rang to alert the operators. For safety purposes, pressure relief valve (PRV601 &
PRV602) will be installed at the top part of reactor which connected to the venting system in
the plant.
Since the main product of ammonia process is anhydrous ammonia, further cooling is
required to bring the hot ammonia product to approximate -33 . Thus, a heat exchanger
(HX-602) can be used at this section. Cooling water will be channel through the tube side and
ammonia product from reactor will be transfer through the shell side. Similarly, a temperature
control loop will be used to control the ammonia product with the flow of cooling water in. A
temperature transmitter will send the electric signal to the temperature indicator control room.
From the control room, electric signal will again being send to control valve (CV603) to
control the flow. When the outlet temperature is not being cooled to the desired temperature,
CV603 will adjust the opening and let the cooling water stay for longer time in the heat
exchanger for further cooling.
Full Report
CHAPTER 7 | PROPER DEFINITION OF BASIS, CRITERIA AND LIMITS OF

DESIGN
7.1 Definition of Design Basis
7.1.1 Functional Goals

Alternis BioAmmonia plant intends to design a 30 kilotonne per annum (ktpa) anhydrous
fertilizer-grade ammonia production plant using oil palm trunks as a potential feedstock
which is expected to be more economical and sustainable. Processes included in the design
are gasification of oil palm trunk with steam, secondary reforming, shift conversion, carbon
dioxide (CO2) removal, gas purification, NH3 synthesis and NH3 Separation. The anhydrous
fertilizer grade ammonia production plant was designed as a low carbon facility by analyzing
its technical, social, environmental and economic viability.
The plant site, which is located in Langkap, Perak was chosen based on the feedstock
specifications, plant capacity, site characteristics and utilities. Availability of link roads,
infrastructure and the site being a sufficient distance away from residential areas are also
taken into consideration.
The relevant process technologies for each end to end process inside the system boundary
were compared, evaluated and selected based on economic, safety and environmental
considerations. Since the sustainability of the technology would play a significant role in the
decision making process, therefore, the technologies selected are expected to produce the
least amount of harmful emissions, capable of generating high feedstock conversion
efficiencies and would be cost effective to operate and maintain.
The development of process flow diagram (PFD) comprised all of the equipment involved in
the processes were presented with the intention to provide the most sustainable plant design.
This would be achieved by following the sustainability concepts, one of which by reducing
the cause of emission within the process. In addition, the process design will aim at
recovering heat and water whenever possible and to treat waste in order to be reused in a
different process through.
Moreover, mass and energy balance were also performed for the processes inside the system
boundary to obtain required parameters inclusive of process efficiency, conversion, yield and
the amount of waste and by product produced. Heat and water recovery will be practiced
wherever applicable, and the carbon footprint will be maintained at a low and acceptable
Full Report
values to enhance sustainable practice. For instance, the ammonia plant designed in a way
that the CO2 removed at the Carbon Dioxide Removal Stage is being sold to the nearby
Glycol Plant to be used for other application and further reduce emission to environment.
Thorough sizing of equipment items with detailed specifications were also carried out to
enable capital and operating cost estimations. Thus, evaluation of the detailed mechanical of
all main process equipment along with its completed calculations were perfomed. Factors
such as corrosion allowance, design loads, minimum practical wall thickness, internal
pressure, external pressure, combined loading and vessel supports were taken into account for
the design calculations. By utilizing proper design and material of construction, such as
employing appropriate wall thickness and utilizing proper corrosion protection in the design
of the equipment, would allow high performance operating life of the plant with low
maintenance requirements and smooth operation of the plant. As to ensure the safety aspect
of the plant, piping instrumentation diagrams (P&IDs) of all the equipment, which accounts
for the safety of the process equipment, were analysed and included in the report as well.
Aside from that, safety assessment for the main equipment of the plant and an environmental
assessment for the operational phase of the entire plant were also performed in the design
project. For the safety assessment, possible hazards occurring in the process were identified
and mitigation methods were proposed. Also, emergency response protocols will be
discussed following the execution of a bow-tie diagram for at least one hazardous scenario
identified in the P&ID of the main equipment.
In addition, the design of the plant would also be illustrated through a plant layout drawing to
understand and get a clear look on the complete plant design. Besides, economic evaluation
of the entire project including market evaluation, capital cost estimation, operating cost
estimation, profitability evaluation and a critical overview will also be performed in order to
determine the economic viability of project.
Full Report
7.1.2 Budgeting
Engineering costs generally includes all the contractor charges, home office costs, costs
associated with detailed design and other engineering that is essential to carry out the project.
Contingency charges are the extra costs that have not undergone any of the previous
categories, and that should be added to the project budget to allow for variations in cost
estimates. Generally, engineering costs and contingency charges are each taken as a
minimum of 10% of the ISBL (Sinnot & Towler, 2009). Thus, the total engineering costs
and contingency charges were calculated to be $9.68 million.
7.1.3 Reliability and Durability

In terms of reliability, the design of the equipment focuses on minimizing the effect of the designed
equipment on the environment. In terms of durability, on the other hand, the structural design of the
equipment aims to enhance the resistance and mechanical strength of the equipment that are subjected
to the inevitably harsh environment.
7.1.3.1 Codes and Standards

Table 7.1.3.1 Summary of codes and standards relating to the reliability and durability of the equipment
Equipment Codes and Standards
Pressure vessels  The pressure vessels were designed in such a way that
they are able to endure maximum exerted pressure and
handle extreme fluctuations in operating pressure.
 The design pressures of vessels that are subjected to
respective internal pressures were set at 10% above the
normal operating pressures following the API RP 520
design practice
 In compliance with the BS EN 13445 code, given a
design pressure, the design temperature of the vessel
was set at a temperature which was more than the
maximum fluid temperature with the maximum design
temperature being the maximum working temperature of
the material at which the maximum allowable stress is
determined under the ASME BPV code, and the
minimum design temperature being the lowest
temperature expected during operation as stated in the
ASME BPV Code Sec. VIII D.1 part UW
 The maximum allowable stress values in ASME BPV
Code were estimated using the ASME BPV Code Sec. II
Part D, Mandatory Appendix 1 where creep and stress
rupture strength do not govern the selection of stresses
Full Report
Distillation Column  The detailed design of absorption column was carried

out in accordance to the American Society of
Mechanical Engineers (ASME) standard. The
procedures for detailed design of the absorption column
and support based on ASME standard were taken from
Sinnott & Towler (2009). The set of codes covered by
ASME standard are listed as follows:
 Minimum thickness of the vessel
 Type of head and end and the minimum thickness
 Maximum allowable stress of the material at given
temperature
 Corrosion allowance
 Internal and external stresses
 Vessel support
Heat Exchanger/ Waste  Standards of American Tubular Exchanger
Heat Boilers/ Condenser Manufacturers Association, TEMA, which identify heat
exchanger by a three letter code were used
 The TEMA standards provide the preferred shell and
tube dimensions, the design and manufacturing
tolerances, corrosion allowances and the recommended
design stresses for materials of construction
Pumps/ Compressors  Coulson et al. (1999) method was used for the
calculation of pressure drop through pipes and fittings
 Kerns (1975) method was used for the practical design
of pump suction piping
7.1.3.2 Quality of Materials and Construction

Some of the components present in the syngas are very corrosive due to the presence of a
variety of dissolved materials, typically the most aggressive being CO2 and hydrogen. Hence,
in order to achieve a high service factor, the material selection based on quality is important.
The choice of materials is dependent on the species present, their concentration, temperature,
fluid velocity, as well as the type of equipment (vessels, furnace, pump, piping,
etc.) .Essentially, complete identification of all the materials potentially present in the feed is
necessary for the proper selection of materials. In some cases, carbon steel with a 4mm
corrosion allowance was used instead of costly stainless steel in order to save cost. By
considering the safety aspect, mechanical properties and corrosion resistivity of material, a
quality assessment is summarized in the table below.
Full Report
Table 7.1.3.2: Summary of the quality of materials used for the design of the equipment
Material Type of process Safety Aspect Mechanical Properties Corrosion Resistivity
Alumina  Reforming at  Ability to resist  Ability to withstand  High alumina bricks
Silicate high temperature distortion at high the operating of 60% SiO2 are
Refractory up to 1300°C temperature temperature at a range resistant to attack by
of 1300-1700°C alkalis
(ToolBox, 2013)
 Its lower thermal
conductivity of
1.3W/m.K minimizes
heat loss to the
surrounding
Stainless Steel  Reforming at  Are able to  Gives an austenitic  High corrosion
310S high withstand stress structure which has resistance with
temperature up corrosion greater strength in chromium content of
to 1300°C cracking at comparison with plain more than 12% and
 Carbon extremely high carbon steels, the addition of nickel
monoxide temperatures particularly at elevated
conversion to temperatures
carbon dioxide
at a temperature
range between
200-350
 Ammonia
synthesis
Low Alloy  Carbon  Ability to  Provides mechanical  Corrosive resistance
Steel A387 Gr monoxide withstand strength that is slightly is similar to that of
22 conversion to external higher than that of the plain carbon steel
carbon dioxide pressures and carbon steel at
at a temperature impacts caused elevated temperatures
range between by the
200-350 environment
Carbon Steel  Carbon  High resistance  Good weld-ability  Low resistance to
Grade A285 monoxide to ignition in  Provides good corrosion except ub
conversion to oxygen and slow tensile strength and certain specific
carbon dioxide rate of has high toughness environments, such
at a temperature combustion (Gandy, 2007) as sulphuric acid and
range between  Susceptible to caustic alkali and
200-350 alkaline stress suitable for most
 Carbon dioxide corrosion organic solvent
removal by cracking in except chlorinated
amine solvent environment solvent
where both containing CO2
components are
corrosive
 Methanation
process
Full Report
 Used mainly for

vessel supports
Carbon Steel  Ammonia  Has a lowest  Excellent notch  Safe from nature of
ASTM A516 product service toughness for lower corrosion due to the
recovery temperature at - ambient separator that is
products, which 45 temperatures operating at a sub-
involves  Susceptible to services
zero temperature
flashing the stress-corrosion
process streams cracking
to lower  Irreversible
pressure and cracking due to
temperature hydrogen rich
within its dew environment as
and bubble point hydrogen can
diffuse into the
carbon steel and
react with
carbon to form
methane at
temperatures
above 350
Epoxy-based  Weather attack  Provides  Good adhesion  Rust-inhibitive/
Paints  Chemical protection corrosion resistant
abrasion on against the
surfaces of surrounding and
vessels impacts due to
change in
weather
Mineral  Contamination  Ability to absorb  Resistant to wear  Provides little
Wool from weather noise generated and tear induced by protection against
conditions ad by the reactor negligence and site corrosion
monsoonal thus reducing conditions
changes noise pollution
 Inert to the
components
present in
syngas in case
of minor leaks
from the
equipment
Full Report
7.1.4 Flexibility
Alternis BioAmmonia plant is design for a specific production capacity, and the processes are
optimized with regards to investment and operating costs for the specific capacity. This is
mainly due to the limitation in the availability of the raw materials required by the plant,
which is the palm oil trunk which only available during the felling process of oil palm tree
cultivation. Therefore, to satisfy the economic stability of the plant, the plant is designed,
such that it could be operated cost-effectively following the trend in the market demand. The
overall design and selection of technologies explained earlier were all completed considering
this issue as well. Furthermore, in the case of increasing the market demand of ammonia,
extra plots of land are provided for plant expansion, as stated in plant layout section, which
enables future modifications of the plant. The current production rate is considered low, as
this is only the first stage of the project development. Therefore, before suggesting expansion
of the plant, Alternis BioAmmonia plant would first need to further evaluate the viability of
the plant based on technical, economic and environmental considerations.
7.1.5 Maintainability
Maintenance on a plant is normally carried out to prevent problems from occurring, to put
faults right, as well as to make sure that the equipments are functioning effectively, so that
the operation of the plant can be run smoothly. An effective maintenance program of the
plant will make the operation and equipments more reliable. Therefore, maintenance schedule
and program need to be planned and carried out efficiently.
The engineers or staffs who are responsible for their respective section must have a working
knowledge of the equipments in that section, the required maintenance process for each
equipments, as well as spare parts to be stored. Besides, a record must be kept whenever
repairs are made to each equipment. This will allow the supervisor or senior engineers to
understand and to make appropriate judgments about the maintenance program, the quality
and condition of equipment, as well as the replacement time of the equipment.
In addition, all routine procedures must be grouped and kept together on a checklist
according to the scheduled frequency. The procedures are normally scheduled for specific
time periods so that the maintenance works on the equipment can be uniform over the
calendar year. Besides that, all the maintenance works will be conducted by engineers who
are qualified and knowledgeable in the operation and maintenance of the equipment. Most
Full Report
importantly, all the maintenance works and procedures will be conformed to the
manufacturers’ recommendations.
The purity of the product will be checked daily by collecting the samples and further examine
in the laboratory to ensure that the purity is maintained. On weekly basis, inspection on the
pipelines and lubricating oil tank level will be carried out. This is to make sure that the
pipelines are in good condition and the lubricating oil are always above the recommended
level. Furthermore, all the measurement detectors and indicators will be examined and
inspected regularly based on monthly basis. Lastly, the plant will be shut down for about two
to three weeks annually for overall inspection and maintenance. Every equipment in the plant
will be evaluated to make sure that the equipment are in good condition so that smooth
operation of the plant can be promised.
7.1.6 Environmental Evaluation
7.1.6.1 Introduction of Environmental Aspect and Impact Register

The main objective of this environmental evaluation is to study and evaluate the
establishment of the process plant as well as determine the potential environmental impact
from the sub-processes of the Anhydrous Ammonia Plant. Besides, control and mitigation
measures as well as legislation guidelines will be included to ensure that activities from the
process plant are under control and thus prevent and reduce the significant environmental
impacts. The environmental evaluation will be done by developing an Environmental Aspect
and Impact Register Table.
7.1.6.2 Scope of Environmental Aspect and Impact Register

The main focus of the environmental aspect and impact register is prioritized on the
environmental aspects and impacts of the normal operating phase, abnormal operating phase,
maintenance/cleaning operations and startup and shutdown in the Anhydrous Ammonia Plant.
Furthermore, all the possible emissions and its identifications and significances are included
in the evaluations.
The evaluations of environmental aspect and impact will be conducted based on the processes
in anhydrous ammonia plant such as Pre-treatment, gasifier, waste heat boiler, post-treatment,
Full Report
autothermal reformer, gas shift reactor, carbon dioxide removal, methanator, ammonia
synthesis and purification of ammonia.
7.1.6.3 Methodology
The studies of environmental aspect and impact register will be performed by approaching
the following steps:
Step 1: Select activities or process
The activities and processes that are large and worth for examination and small enough to
understand.
Step 2: Identify the environmenal aspects
Based on the selected activity and process, identify the environmental aspects.
Step 3: Identify environmental impacts
With the aforementioned environmental aspects, identify the actual and potential
environmental impacts.
Step 4: Impact identification and significance
Evaluate and quantify the significance of impacts. The issues that should take into
consideration are the scale of impact, severity of the impact, probability of occurrence and the
duration of impact.
7.1.6.4 Determination of Significance
7.1.6.4.1 Ranking of the Probability of Aspect

Table 7.1.6.4.1.1 Ranking of the Probability of Aspect
Ranking Probability of Aspect
1 Unlikely in 5 years
2 Unlikely in a year
3 Probable in a year
4 Probable in a month
5 Probable in a week
6 Probable in a day
Full Report
7.1.6.4.2 Ranking of the Consequences of Impact

Table 7.1.6.4.2.1 Ranking of the Consequences of Impact
Ranking Consequences
1 Unlikely to have an effect
2 Limited effect
3 Effect but not lasting
4 Potential for serious and lasting harm
7.1.6.4.3 Evaluation Significance of Impact

The significances of impact are calculated using the formula below:
The calculated significances are tabulated in the Table 8.2.4.3.1. However, the degrees of the
significance impact are tabulated in Table 8.2.4.3.2. Figure 8.2.4.3.1 shows the risk contour
based on the multiplication of Probability of Aspect and the Consequences of Impact.
Table 7.1.6.4.3.1 Significance of the Impact

Consequences of Impact
Probability of Aspect 1: 2: 3: 4:
Unlikely to have Limited effect Effect but not Potential for
an effect lasting serious and lasting
harm
1: Unlikely in 5 years 1 2 3 4
2: Unlikely in a year 2 4 6 8
3: Probable in a year 3 6 9 12
4: Probable in a month 4 8 12 16
5: Probable in a week 5 10 15 20
6: Probable in a day 6 12 18 24
Table 7.1.6.4.3.2 Degrees of Significance Impact

Range of Calculated Significance Significance of Impact
1-6 Low Risk
7-12 Moderate Risk
13-18 High Risk
19-24 Critical
Full Report
Figure 7.1.6.4.3.1 Risk Contour based on the Significance of Impact

Full Report
Phase of Project: Normal Operation

Activity Environmental Impact Impact On Risk Control/Mitigation Legislation/Guidelines
Aspect Mechanism Probabilit Consequence Significan Measures
y of ce of
Aspect Impact
Pre-treatment (Transportation, Shredder and Dryer)
Transportatio Usage of Emission of Air quality 4 2 8 Imply bulk Environmental quality
n of oil palm heavy duty greenhouse gases Probable Limited Moderate transportation to (Control of Emission
trunk to vehicles (GHGs) by the in a month Effect Risk reduce the emissions from Diesel Engines)
anhydrous heavy duty Regulations 1996/
ammonia plant vehicles Direct and Regional
contributes to air
pollution
Depletion of Natural 5 2 10 Maintenance of the -
fossil fuel due to resources Probable Limited Moderate transport
usage in vehicles in a week effect Risk
Transport Increase the Noise 5 2 10 Transportation of Laws of Malaysia Act
route nearby traffic load quality Probable Limited Moderate feedstock at 333 – Road Transport
the process around the and traffic in a week effect Risk specified time Act 1987/
plant process plant and volume Direct and Local
causes noise
pollution
Shredding and Shredding of Generation of Air and 6 1 6  Monitor and Regulations 36 and 38
drying of the oil palm dust, fine water Probable Unlikely to Low Risk ensure the dust Environmental
feedstock trunk into particles and quality in a day have effect and fine particles Quality (Clean Air)
smaller size water content generated are Regulations 1978,
and removal of from the within of below Environmental
Full Report
water content shredder and the allowable limit Quality (Sewage and
from chips dryer  Install an air filter Industrial Effluents)
near the outlet of Regulations 1979/
the machinery Direct and Local
 Operate the
machinery in an
enclosed area and
well-ventilated
area
 Treat and ensure
the water
discharge at an
allowable limit
Shredding and Noise from the Noise 6 1 6  Ensure that the First Schedule of the
drying of oil shredder and quality Probable Unlikely to Low Risk machinery is Factories and
palm trunk dryer contribute in a day have effect running properly Machinery (Noise
to noise and generate noise Exposure)
pollution level below the (regulation), 1989,
limit Occupational Safety
 Employees should and Health Act 1974/
use earplugs when Direct and Local
working within
area surrounding
by machinery
Gasifier
Combustion of Combustion of Huge amount of Air quality 6 2 12  Well-ventilated Regulation 36 and 38
feedstock feedstock in heat, flue gas Probable Limited Moderate environment is Environmental
Full Report
the gasifier and particulate in a day effect Risk provided to Quality (Clean Air)
dust will be decrease the Regulations 1978
generated and surrounding
may contribute temperature and
to air pollution heat radiation
 Install an air
filter near the
outlet of the
machinery
Particulate dust Air 6 2 12  Monitor and Regulation 36 and 38
from the gasifier Quality Probable Limited Moderate ensure the dust Environmental
might fly to the in a day effect Risk and fine particles Quality (Clean Air)
nearby generated are Regulations 1978
residential area within of below
the allowable
limit
Waste Heat Boiler
Generation of Steam is High Air quality 6 1 6  Well-ventilated Regulation 36 and 38
steam generated in temperature is Probable Unlikely to Low Risk environment is Environmental
waste heat required to in a day have effect provided to Quality (Clean Air)
boiler generate steam decrease the Regulations 1978
and thus surrounding
generate a lot of temperature and
heat heat radiation
Post-Treatment (Cyclone and Scrubber)
Removal of Cleaning of Effluent and Air and 6 1 6  The solid waste Regulations 36 and 38
solid particles, solid particles solid waste will water Probable Unlikely to Low Risk discharge from the Environmental
Full Report
tar and water in the syngas be generated quality in a day have effect cyclone will be Quality (Clean Air)
in syngas from the post- recycle back to Regulations 1978,
treatment gasifier and use as Environmental
process combustion agent Quality (Sewage and
 Optimize the Industrial Effluents)
process and Regulations 1979/
reduce the effluent Direct and Local
generated
Autothermal Reformer Section (Fired Heater and Autothermal Reformer)

Combustion of Combustion of Heat and flue gas Air quality 6 2 12  Well-ventilated Regulation 36 and 38
natural gas in natural gas in and particulate Probable Limited Moderate environment is Environmental
fired heater fired heater in dust will be in a day effect Risk provided to Quality (Clean Air)
order to heat generated and decrease the Regulations 1978
up the syngas may contribute surrounding
to desired to air pollution temperature and
temperature heat radiation
 Flue gas generated
will be utilized in
the heat exchanger
to heat up the air
entering the
reformer
subsequently send
to stack
Particulate dust Air 6 2 12  Monitor and Regulation 36 and 38
from the gasifier Quality Probable Limited Moderate ensure the dust Environmental
Full Report
might fly to the in a day effect Risk and fine particles Quality (Clean Air)
nearby generated are Regulations 1978
residential area within of below
the allowable
limit
Gas Shift Reactor
Water-gas Usage of Leakage of Water 6 1 6  Monitor and take Environmental
shift reaction Ferum catalyst during quality Probable Unlikely to Low Risk prompt action if Quality (Sewage and
Chromium the process in a day have effect there is leakage Industrial Effluents)
catalyst and of catalyst Regulations 1979/
zinc oxide Direct and Local
catalyst in the
reaction
Carbon Dioxide Removal
Carbon Separation Effluent will be Water 6 1 6  Optimize the Environmental
dioxide process to leaving the quality Probable Unlikely to Low Risk water usage to Quality (Sewage and
removal remove the separator in a day have effect reduce effluent Industrial Effluents)
water content Regulations 1979/
Direct and Local
Methanator
Separation Separation of The water Water 6 1 6  Water discharged Environmental
process syngas and separated out quality Probable Unlikely to Low Risk will be sent to Quality (Sewage and
water will be in a day have effect wastewater Industrial Effluents)
discharged from storage Regulations 1979/
the separation Direct and Local
tank
Ammonia Synthesis
Full Report
Transferring Fugitive Process stream Air quality 6 2 12  Pipe clips and Regulation 36 and 38
the syngas into emission due will be leaked Probable Limited Moderate flanges can be Environmental
ammonia to the leakage and released to in a day effect Risk installed to tighten Quality (Clean Air)
reactor of process the environment and this minimize Regulations 1978/
stream near the the leakages Indirect and Regional
gaskets joint
Ammonia Purification
Refrigeration Fugitive Efficiency of the Air quality 5 2 10  Pipe clips and Regulation 36 and 38
cycle emission due system decreased Probable Limited Moderate flanges can be Environmental
to leakage of as R-717 is in a week effect Risk installed to tighten Quality (Clean Air)
R-717 from leaked and this and this minimize Regulations 1978/
the compressor lead to higher the leakages Indirect and Regional
and valve seals power  Pungent smell of
into the consumption as the R-717 can be
environment well as the detected easily
emission of therefore prompt
carbon dioxide action can be
taken if there are
leakages
 Sensitive
electronic leak
detector can be
installed to
identify the
leakages
Purge gas Methane and Release of Air quality 5 2 10  Monitor the Regulation 36 and 38
traces of methane, Probable Limited Moderate process and Environmental
Full Report
ammonia, ammonia, in a week effect Risk prompt action if Quality (Clean Air)
hydrogen and hydrogen and leakage happen Regulations 1978/
nitrogen nitrogen to Indirect and Regional
environment lead
to air pollution
Overall Plant Operation and Maintenance
Fresh water Consumption Depletion of Natural 6 1 6  Minimize the Environment Quality
consumption of fresh water natural resources resources Probable Unlikely to Low Risk wastage and usage Act, 1974/
in a day have any of water Indirect and Regional
effect
Disposal of Solid waste Causes Water 5 1 5  The solid waste Environmental
solid waste from overall groundwater quality Probable Unlikely to Low Risk can send to Quality (Prescribed
plant and pollution due to and land in a week have an landfill or Premises) (Scheduled
maintenance highly soluble effect incineration Wastes Treatment and
residue  Propose waste Disposal Facilities)
management plan Regulations 2006/
The solid waste which consider all Indirect and Regional
increase the details such as the
waste and causes period for waste
soil pollution disposal,
minimization of
waste, method of
handling and
storage,
transportation and
method of
disposal
Full Report
Waste water Generation of Wastewater Water 4 1 4  Propose suitable DOE Recommended

generation wastewater generated from quality Probable Unlikely to Low Risk and efficient Emission Limits/
from the the plant in a month have an discharge system Indirect and Regional
overall plant contaminated the Biodiversi effect
and surface and ty
maintenance groundwater
period
Electricity Electricity Depletion of Natural 6 1 6  Optimize the -
Usage usage during fossil fuel resources Probable Unlikely to Low Risk usage of
operation and in a day have any electricity
maintenance effect
Operating Fugitive Leakage gas will Air 6 2 12  Monitor the Regulation 36 and 38
Process emission due be releasing to Quality Probable Limited Moderate process and take Environmental
to leakage of environment in a day effect Risk prompt action if Quality (Clean Air)
pipe there is leakage Regulations 1978/
detected Indirect and Regional
Full Report
Phase of Project: Abnormal Operation

Activity Environmental Impact Impact On Risk Control/Mitigation Legislation/Guidelines
Aspect Mechanism Probability Consequence Significance Measures
of Aspect of Impact
Power failure Usage of All the Interruption 3 3 9 Generators need to Occupational Safety and
electricity in machinery, to the process Probable Effect but not Moderate be backed up to health Act 1994/
the process equipment and plant in a year lasting Risk ensure there are Indirect and Local
plant operation operations sufficient energy
processes are supply to plant
shut down due especially in plant
to the cut off of safety system
electrical
supply to the
plant
Autothermal Hydrogen High pressure Embrittlement 1 4 4 Monitor the process Occupational Safety and
reformer and embrittlement and temperature hydrogen can Probable Potential for Low Risk plant and imply health Act 1994/
Ammonia reactor, lead to in 5 years serious and emergency plant Indirect and Local
syhthesis hydrogen explosion lasting harm shut down system if
reactor embrittlement the reactor works
might happened abnormally
Full Report
Phase of Project: Start-up and Shut Down Operation

Activity Environmenta Impact Impact Risk Control/Mitigatio Legislation/Guideline
l Aspect Mechanis On Probabilit Consequenc Significanc n Measures s
m y of e e of Impact
Aspect
Start-up operation
Electricity Consumption Depletion Air 3 1 3 Imply the most DOE Recommended
consumptio of electricity of natural quality Probable Unlikely to Low Risk energy efficient Emission Limits and
n from during startup resources and in a year have an machinery, Malaysian Ambient
Malaysia (preheater, and depletion effect equipment and Air Quality
National pumps, indirect of plant system Guidellines
Grid compressors, emission natural (MAAQG)/
gasifier) of GHG to resource Indirect and Global
atmosphere s
Venting and Fugitive Air Air 3 1 3 Installation purge Malaysian Ambient
emptying emission Pollution quality Probable Unlikely to Low Risk recovery system Air Quality
the during the in a year have an and closed loop Guidellines
equipment shutdown effect system (MAAQG)/
in the process Indirect and Global
process
plant
Shut-down Operation
Venting and Fugitive Air Air 3 1 3 Installation purge Malaysian Ambient
emptying emission Pollution quality Probable Unlikely to Low Risk recovery system Air Quality
the during the in a year have an and closed loop Guidellines
equipment shutdown effect system (MAAQG)/
in the process Indirect and Global
process
plant
Full Report
7.1.7 Safety
7.1.7.1 Inherent and Extrinsic Safety

Processes are divided into those that are intrinsically safe, and those which safety has to be
engineered in (extrinsic). An intrinsically safe process is one whereby safe operation is
inherent in the nature of the process; a process which causes no danger, or negligible danger,
under all foreseeable circumstances (all possible deviations from the design operating
conditions / worse case scenarios). Intrinsic safety involves designing the process to be
inherently safe, thereby eliminating the safety risk. Wherever practicable and economic, an
engineer should opt for a process that is deemed inherently safe. However, many industrial
processes are more often than not, inherently unsafe, and deviations from process parameters
could result in extremely hazardous situations. When the team has reached the end-point of
inherent safety, extrinsic safety systems are used to reduce the risk to the tolerable risk
level. Extrinsic safety systems are add-on devices, included in the design for the explicit
purpose of preventing or mitigating risk. Safety of operation of these processes depend on
the design and provision of engineered safety devices, and on good operating practices, to
prevent a dangerous situation developing, and to minimise the consequences of any incident
that arises from the failure of these safeguards. The term “engineered safety” covers the
provision in the design of control systems, alarms, trips, pressure-relief devices, automatic
shut-down systems, duplication of key equipment services; and fire-fighting equipment,
sprinkler systems and blast walls, to contain any fire or explosion.
Layers of protection methodology as illustrated in Figure 7.1.7.1.1 was used as the tool for
risk assessment. The specific hazards associated with the major and minor equipment within
the primary reforming section of the plant and how the hazards are mitigated are assessed
and summarized in section 7.1.7.5 below. The specific hazards associated with the major and
minor equipment within the shift reaction section of the plant and how the hazards are
mitigated are assessed and summarized in section 7.1.7.6 below. On the other hand, section
7.1.7.7 summarizes the specific hazards and mitigation measures associated with the carbon
dioxide removal section of the plant whereas section 7.1.7.8 is for methanation section of
the plant. Lastly, section 7.1.7.9 illustrates the specific hazards and mitigation measures
associated with ammonia synthesis section of the plant.
Full Report
Figure 7.1.7.1.1: Layers of protection
7.1.7.2 Emergency Response Plan (ERP)

The first step when developing an emergency response plan is to conduct a risk assessment to
identify potential emergency scenarios. An understanding of what can happen will enables us
to determine resource requirements and to develop plans and procedures to safeguard against
potential hazards. The emergency plan should be consistent with the performance objectives
of the plant. The objective of this section is to provide guide measures for all employees in
the ammonia production plant on the actions that should be taken upon emergency or worse
case scenarios in order to mitigate the hazards to avoid major incidents. At the very least, the
facility should develop and implement an emergency plan for protecting employees, visitors,
contractors and anyone else in the facility. This part of the emergency plan is called
“protective actions for life safety” besides also minimizing the damage to the equipments in
the plant. The emergency response plan was developed for an event whereby ammonia gas
leaks from the plant and is illustrated in Figure 7.1.7.2.1.
Full Report
Leakage of ammonia gas
Figure 7.1.7.2.1: Emergency response plant for ammonia gas leakage. Source: (Tseng, 2008)
Full Report
7.1.7.3 Introduction of Process Hazards
7.1.7.3.1 Fire Hazards

Release of syngas containing methane may result in fire upon ignition by any fire sources.
Under high pressure, intense flames with high heat radiation intensity can be produced. This
may lead to forming of large vapour cloud that will cover a large area in the plant. Fire or
flame might cause damage to both the workers and the equipment.
7.1.7.3.2 Corrosive Hazards

Amine solvent used in the absorption column contains MDEA and piperazine which are
corrosive. Exposure to the amine solvent might cause injury to the workers and sometimes
can even cause fatal, depending on the extent of leakage of the amine solvent, the
concentration of amine solvent as well as the duration of the exposure. On top of that, the
corrosive nature of the amine solvent might also cause damage to the equipment.
7.1.7.3.3 Thermal Hazards

Thermal hazards are possible to occur when energy that is released during desired reaction or
undesired reaction cannot be controlled. This may cause by deviation of the processes from
the normal operating conditions. Thermal hazards may lead to serious damage to the workers
in the plant. The extent of the impact depends on the extent of the deviation of the process
from normal operation and spreading of the heat energy.
7.1.7.3.4 Explosion Hazards

Several equipments in the main process sections are operating under high pressure.
Overpressure of these process equipments might lead to explosion hazards. Gas cloud formed
after explosion hazard might cause combustion if it falls within the flammable concentration
range and an ignition source is present. The pressure and heat generated from combustion
might cause damage and injury to both the workers and the process equipment.
Full Report
7.1.7.4 General Mitigation Measures for Hazards

There are several general mitigation measures applied to the whole plant as general safety
precautions. The general mitigation measures are summarized below:
 First Aid response team is ready in case of emergency

 Fire respond team is available in case of fire hazards or explosion hazards
 Appropriate Personal Protection Equipment (PPE) are provided for the workers and
operators during start up and shut down process, normal operation, as well as
maintenance of the process equipment in the plant.
 Good emergency response plan is established in case of evacuation
 Proper start up and shut down process
 Provide intensive training to the operators to make sure safe operations of the plant
 Annual shut down of the plant for maintenance purposes
Full Report
7.1.7.5 Safety Assessment for Primary Reforming Section
7.1.7.5.1 Hazards Identified, Causes, Outcomes and Mitigation Measure
Table 7.1.7.5.1.1: Possible hazards, their causes, outcomes and mitigation measures for major equipment.
Major Operating Hazard Preventive and Mitigative Safety
Causes Possible Outcomes
Equipment Conditions Identified Measures
Autothermal Start-up Fire and  Inappropriate start-up  Release of syngas gas, which  Provide sufficient trainings for
Reformer explosion procedure. contains flammable workers on start-up procedure.
(ATR-201) hazard  A sudden high inflow of components such as methane.  Provide a start-up checklist on safe
syngas.  Release of the hot syngas to procedure and audited annually by
 Malfunction of the valves at the atmosphere causing skin professional.
the inlet of the autothermal burns and harm to the  Install temperature and pressure
reformer. environment. indicator to monitor temperature and
 Pressure-build up due to the pressure in vessel, tubes and in
high inflow of the gas. pipeline.
 Vessels and pipelines rupture.  Installation of temperature control
detecting the temperature of the
syngas leaving the autothermal
reformer and control valve to
control the entering flow rate of
water.
 Monthly inspection of the
controllers, valves and equipment.
Normal Thermal  High operating temperature  Causes burn due to direct  Appropriate personnel protective
Operation hazard of syngas. contact. equipment such as glove is supplied
to employees.
Full Report
 Mineral wool is used as insulation

layer to prevent heat loss to the
atmosphere and at the same time to
avoid possible exposure of high
temperatures.
Fire and  Faulty temperature control  Rupture of the vessel and  Installation of pressure indicator to
explosion system and controller valve. pipelines if pressure becomes monitor pressure.
hazard  Fluctuation on inlet flow much higher than the design  Installation of check valve to
into the autothermal pressure. prevent the backflow of the process
reformer causing rise in  Abnormal operating condition gas.
temperature and pressure. causing damage to subsequent  Temperature and pressure indicator
 Malfunction of the valves at equipment. to monitor temperature and pressure
the outlet of the waste heat  Release of syngas gas, which of the autothermal reformer.
boiler causing back flow of contains flammable  Monthly inspection of the
syngas causing pressure components such as methane. controllers, valves and equipment.
build-up in vessel.  Release of the hot syngas to  First aid and firefighting system.
 Failure of the pressure-relief the atmosphere causing skin  Automatic fire sprinkler system.
valve. burns and harm to the
 Overflow/no flow of syngas environment.
at high temperature in the  May cause explosion in the
reformer. vessel due to the high
temperature and high pressure
contents.
Maintenance Fire and  Rupture of the vessel and  Proper venting before maintenance
and shut- explosion  Incorrect shut down pipelines. and shutdown.
Full Report
down hazard procedure.  Release of syngas gas, which  First aid and firefighting system.
 Valves fail to close contains flammable  Automatic fire sprinkler system.
completely resulting in components such as methane.  Proper maintenance and shutdown
presence of process fluid  Release of the hot syngas to procedure.
within the autothermal the atmosphere causing skin
reformer during burns and harm to the
maintenance. environment.
 May cause explosion in the
vessel due to the high
temperature and high pressure
contents.
Full Report
Table 7.1.7.5.1.2: Possible hazards, their causes, outcomes and mitigation measures for minor equipment.
Minor Operating Hazard Preventive and Mitigative Safety
Heat Start-up Fire and  Inappropriate start-up  Release of syngas gas, which  Provide sufficient trainings for
Exchanger explosion procedure. contains flammable components workers on start-up procedure.
(HX-207) hazard  A sudden high inflow of such as methane  Temperature and pressure indicator
syngas.  Release of the hot syngas to the to monitor temperature and
 Malfunction of the valves at atmosphere causing skin burns pressure of the heat exchanger.
the inlet of the heat and harm to the environment.  Make sure the vent valve is
exchanger.  Pressure-build up due to the completely closed before start up.
 Vent valve is opened. high inflow of the gas.  Monthly inspection of the
 Vessels and pipelines rupture. controllers, valves and equipment.
Normal Fire and  High operating temperature  Burns from direct contact.  Provide appropriate personnel
Operation explosion of the heat exchanger  Release of syngas gas, which protective equipment.
hazard  Fouling in tube side of heat contains flammable components  Mineral wool insulation is installed
exchanger such as methane around the heat exchanger to
 Blockage or leakage in  Release of the hot syngas to the prevent heat loss to the atmosphere
pipelines. atmosphere causing skin burns and to avoid possible exposure of
and harm to the environment. high temperatures.
 Vessels and pipelines rupture.  Installation of pressure indicator to
monitor pressure.
 Automatic fire sprinkler system.
 Use stainless steel that can
withstand high temperature and
pressure.
Full Report
Maintenance Fire and  Incorrect shut down  Presence of methane may cause  Proper drainage and vent system
and shut- explosion procedure. fire. before maintenance and shutdown.
down hazard  Valves fail to close  May cause subsequent pipeline  First aid and firefighting system
completely resulting in or equipment failure.  Automatic fire sprinkler system.
presence of process fluid  Proper maintenance and shutdown
within the heat exchanger. procedure.
Full Report
Table 7.1.7.5.1.3: Possible hazards, their causes, outcomes and mitigation measures for minor equipment
Minor Operating Hazard Preventive and Mitigative Safety
Centrifugal Start-up Fire and  Inappropriate start-up  Failure of equipment.  Provide sufficient trainings for
Compressor explosion procedure.  Abnormal operating condition workers on start-up procedure.
(K-205) hazard  Presence of liquid in causing damage to subsequent  Install a liquid sensor at the suction
Mechanical compressor suction. equipment. of compressor to detect presence of
failure  Faulty valve and controller  Possible pipeline rupture on liquid.
 Malfunction of electric downstream.  Use a more durable material for the
circuits or motors. pipelines.
Normal Fire and  Blocked at suction and  Failure of equipment.  Installation of pressure indicator to
Operation explosion discharge.  Abnormal operating condition monitor pressure.
hazard  Failure of the compressor causing damage to subsequent  Installation of spare compressor for
motor. equipment. use in the case of malfunction of
Mechanical  Presence of liquid in  May lead to subsequent main compressor.
failure compressor suction. pipeline and equipment  Installation of check valve to
 Backflow of the rupture. prevent the backflow of the process
compressed gas. gas.
 Monthly inspection of valves and
the compressor.
Maintenance Fire and  Valves and controller  Presence of methane may cause  First aid and fire response team
and explosion failure. fire.  Backup system with proper bypass
Shutdown hazard  Valves fail to close  May cause subsequent pipeline  Regular inspection and
completely resulting in or equipment failure. maintenance.
Mechanical
presence of process fluid
failure
within the compressor.
Full Report
7.1.7.5.2 Bow Tie Diagram for Hazardous Scenario in Primary Reforming Section
Full Report
7.1.7.6 Safety Assessment for Shift Reaction Section
Table 7.1.7.6.1.1: Possible Hazards, Causes, Outcomes and Mitigation Measures for Major Equipment
Operating
Equipment Hazard Causes Possible Outcomes Mitigation
Conditions
Water gas Start up Thermal  Malfunction of heater  Damage and  Appropriate start-up
shift hazard  Temperature of deactivation of procedure
reactors nitrogen used for catalyst due to  Temperature and flow rate
start-up is too excessive high control system
high/low temperature
 Process gas is  Catalyst is not fully
introduced too fast activated due to
insufficient
temperature
 Low conversion in the
reactors
Normal Thermal  Insufficient flow of  Overheating of  Temperatureand flow rate
operation hazard cooling water into the catalyst tube leads to control system for cooling
reactors results in deactivation of water
rapid increase of catalyst  Appropriate temperature,
reactor’stemperature  Insufficient cooling pressure and flow rate
 Excessive flow of leads to low alarm system on the inlet
cooling water into the conversion rate in low of reactor
reactors results in low temperature water gas  Appropriate insulation
temperature of reactor shift reactor around reactor
 Abnormal inlet syngas  Excessive cooling  Provide appropriate
Full Report
condition leads to slow reaction Personnel Protection

(temperature, pressure rate Equipment (PPE) for
and flow rate)  Causes burns or workers
injuries when comes  Restrict the contact of
in contact workers with the reactor by
outlining a safety distance
Fire and  Leakage of flammable  Explosion due to rapid  Appropriate temperature
explosion process gas increase in reactor’s and pressure alarm system
hazard  Pressure build up due temperature and on the reactor
to failure of pressure pressure  Regular inspection of
relief valve  Rupture of reactor’s catalyst and pressure relief
 Presence of ignition wall and catalyst tube valve
source nearby due to high pressure  Proper ventilation to flare
causing leakage system if excessive
 Fire caused by pressure occurs
leakage of process gas  Install methane gas
 Causes burns or detector
injuries when comes  Automated fire sprinkler
in contact alarm and water spray
system
 First aid and emergency
response team
Toxic  Fouling in catalyst  Rupture of reactor’s  Regular inspection on
hazard tube wall and catalyst tube catalyst tube thickness and
 Blockage or leakage  Poisoning of catalyst proper replacement of tube
in pipelines  Release of hazardous  Temperature, pressure and
 Abnormal inlet syngas gases such as methane flow rate control system
Full Report
condition  Process gas leaked together with appropriate

(temperature, pressure and affect the alarm system
and flow rate) downstream operation  Apply coating on the
 Pressure build up due  Leakage of process reactor wall to reduce
to blockage fluids contaminate the corrosion and proper
 Overflow inside fresh water source monitoring of the coating
reactor  Impact on human thickness
 Equipment failure health and  First aid and emergency
 Excessive corrosion environment response team
 Loss of asset
Mechanical  Abnormal inlet syngas  Release of hazardous  Regular inspection on
failure condition gases such as methane catalyst tube thickness and
(temperature, pressure  Causes flooding of proper replacement of tube
and flow rate) cooling water  Appropriate pressure alarm
 Rupture of reactor system on the reactor
wall / catalyst tube  Proper sizing and
due to sudden increase installation of fitting
in inlet pressure
 Improper fitting
Shut down & Toxic  Inappropriate shut  Contamination of  Appropriate shut down
maintenance hazard down procedure process gas in the next procedure
 Pressure inside reactor batch of reaction  Vent the reactor with
is not properly vent  Release of hazardous nitrogen gas before carry
 Cooling water is gases such as methane out shut down procedure
drained out first  Leakage of process  Regular inspection on
instead of hot process fluids contaminate the drainage/venting valve
gas fresh water source due  Appropriate maintenance
Full Report
 Process gas is not to improper drainage procedure

properly vent causing
a vacuum pull
 Cooling water is not
properly drained
Fire hazard  Leakage of flammable  Fire caused by  Regular inspection of
process gas leakage of process gas drainage / venting valve
 Malfunction of  Causes burns or  Install methane gas
venting valve injuries when comes detector Automated fire
 Presence of ignition in contact sprinkler alarm and water
source nearby spray system
response team
Full Report
Table 7.1.7.6.1.2: Possible Hazards, Causes, Outcomes and Mitigation Measures for Minor Equipment
Operating
Equipment Hazard Causes Possible Outcomes Mitigation
Conditions
Heat Start up Thermal  Inappropriate start-up  Causes burns or  Appropriate start-up
Exchanger hazard procedure injuries when comes procedure
 High temperature on in contact  Cooling medium is
shell side allowed to flow through
 Hot medium is filled the heat exchanger first
in first instead of cold before the hot medium.
medium leads to
overheating
Normal Thermal  High temperature on  Insufficient cooling  Temperature and flow rate
operation hazard shell side leads to low control system together
 Malfunction of conversion rate in low with appropriate alarm
control valve leads to temperature water gas system
insufficient cooling shift reactor  Restrict the contact of
 Abnormal inlet syngas  Causes burns or workers with the heat
condition injuries when comes exchanger by outlining a
(temperature, pressure in contact safety distance
and flow rate)  Regular inspection of
valves
 Appropriate insulation
around piping
 Provide appropriate
Personnel Protection
Equipment (PPE) for
workers
Full Report
Fire hazard  Leakage of flammable  Fire caused by  Install methane gas

process gas leakage of process gas detector
 Presence of ignition  Causes burns or  Automated fire sprinkler
source nearby injuries when comes alarm
in contact  First aid and emergency
response team
Toxic  Fouling in tube side  Tube, shell and pipe  Let material which is more
hazard  Blockageor leakagein rupture corrosive to flow in the
pipelines  Release of hazardous tube side and replace the
 Abnormal inlet syngas gases such as methane tube when necessary
condition  Process gas leaked  Regularinspection on tube
(temperature, pressure and affect the side thickness and proper
and flow rate) downstream operation replacement of tube
 Pressure build up due  Leakage of process  Temperature, pressure and
to blockage fluids contaminate the flow rate control system
 Equipment failure fresh water source together with appropriate
 Impact on human alarm system
health and  Automated fire sprinkler
environment alarm
response team
Mechanical  Abnormal inlet syngas  Release of hazardous  Regular inspection on tube

failure condition gases such as methane side thickness and proper
Full Report
(temperature, pressure  Causes flooding of replacement of tube

and flow rate) cooling water  Include a pressure
 Rupture of shell/tube transmitter to monitor the
due to sudden increase inlet pressure
in inlet pressure
Shut down & Toxic and  Inappropriate shut  Release of hazardous  Appropriate shut down
maintenance thermal down procedure gases such as methane procedure
hazard  Cool medium is  Leakage of process  Regular inspection on
drained out first fluids contaminate the drainage/venting valve
instead of hot medium fresh water source due  Appropriate maintenance
to improper drainage procedure
 Overheating of heat
exchanger
 Loss of asset
Pump Start up Mechanical  Inappropriate start up  Equipment failure  Appropriate start-up
failure procedure  Pipeline failure procedure
Normal Mechanical  Inappropriate  Equipment failure  Appropriate operating

operation failure operating condition  Vessel and pipelines condition
 Abnormal inlet rupture  Temperature and flow rate
conditions  Low flow rate of control system together
(temperature, pressure cooling water entering with appropriate alarm
and flow rate) water gas shift reactor system
 Blockage in suction  Backflow of fluid in  Regular maintenance
and discharge the pump  Backup pump is readily
 Blockage or leakage available
in pipelines  Include a pressure
Full Report
 Pump motor failure transmitter to monitor the

 Pump operating inlet pressure
without inlet fluid  Install check valve to
leads to cavitation prevent back flow of
process fluid
Shut down Mechanical  Inappropriate shut  Equipment failure  Appropriate shut down
failure down procedure  Pipeline failure procedure
 Process gas or fluid is  Regular inspection on
not fully drained drainage valve
Malfunction of
drainage valve
Full Report
7.1.7.6.2 Bow Tie Diagram for Hazardous Scenario in Shift Reaction Section
Full Report
7.1.7.7 Safety Assessment for CO2 Removal Section
Operating
Equipment Hazards Possible Causes Possible Outcomes Mitigation Measures
Conditions
Absorption Start Up Thermal  Lower down the CO2
Column hazard due to  Temperature of removal efficiency, causing  Provide appropriate Personnel
(AC-401) upstream heat incoming syngas poisoning of catalyst in Protection Equipment (PPE)
exchanger stream is excessively downstream process  Install temperature indicator and
failure high due to  Cause skin damages (burns) alarm to alert the operators
malfunction of upon contacting the hot  Monitor the exiting syngas
upstream heat surfaces temperature from the heat exchanger
exchanger in upstream process
 Ensure that the column is operating
Hazardous  Release of syngas or under steady state condition by
chemical amine solvent from the  Cause injuries to human as monitoring the temperature change
exposures absorption column due the contents of the absorption  Ensure that workers and operators
to leakage column are hazardous and have gone through thorough training
harmful if released  Regular maintenance of control
 Dissolved CO2 that leaked valves
from the column is very  Regular checking of the equipment
corrosive and may cause  Install pressure indicator and alarm
damage to the equipment and to ensure the pressure of the column
operators upon contacting in within the operating pressure
 Amine solvent leaked from range.
the column may pose thread
Full Report
to both humans and  Install level controller and alarm to

equipment as it is corrosive ensure that flooding does not occur
 Syngas that are released from  First aid and fire response team in
the column may cause fire case of emergencies
and explosion as some  Establish proper start up procedure
components contained in the
syngas is flammable
Rupture /  Rupture of the absorption
Explosion  Sudden high-in flow of column
pressurized syngas and  Pressurized syngas and
amine solvent solution amine solvent is released
into the absorption from the column to the
column surrounding
 Improper start up
procedures
Normal Fire Hazard
Operation  The syngas entering  Absorption column rupture  Install 2 pressure relief valves where
the absorption column  Major and minor pipelines 1 of it acts as backup in case of
from upstream rupture malfunction
contains methane,  Release of syngas containing  Regular checking and maintenance
which is flammable highly flammable on the valves, including control
Hazardous components, which is valves and pressure relief valves
chemical  Release of syngas or methane  Periodic checking on all the
exposures amine solvent from the  Cause fire or explosion when pipelines and the equipment
absorption column due there is ignition  Install pressure indicator and alarm
to leakage  Amine solvent that are to ensure the pressure of the column
Rupture / leaked from the column may in within the operating pressure
Full Report
Explosion  Flow control valve pose thread to both humans range.

failure which lead to and equipment as it is  Install level controller and alarm to
accumulation of amine corrosive ensure that flooding does not occur
solvent in the  Syngas that are released from  First aid and fire response team in
absorption column thus the column may cause fire case of emergencies
causing pressure and explosion as some  Ensure that the absorption column is
buildup in the column components contained in the not close to major ignition source
 Pressure relief valve syngas is flammable  Provide appropriate Personnel
fail to function thus  Cause serious impacts to Protection Equipment (PPE) for the
causing pressure human health workers and operators
buildup in the column  Cause environmental
pollution
 Plant shut down
Maintenance Hazardous  Amine solvent that are  Establish proper maintenance and
and Shut chemical  Release of syngas or leaked from the column may shut down procedures
Down exposures amine solvent from the pose thread to both humans  Provide proper training to the
absorption column due and equipment as it is workers and operators to ensure
to leakage corrosive correct and safe sampling
 Cause serious impacts to procedures are followed
human health  First aid and fire response team in
 Cause environmental case of emergencies
pollution  Provide appropriate Personnel
 Cause pollution to ground Protection Equipment (PPE) for the
water if the amine solvent is workers and operators
not contained after leakage  Install backup of auxiliary
equipment
 Install the absorption column on
Full Report
cement ground
Operating
Conditions
Centrifugal Start Up Rupture /  Improper or incorrect  Discharge valve fails to close
Pump Mechanical start up procedures during pump startup operation  Establish proper start up procedures
(P-401) damage  Entrapped air is present which may lead to damage of  Ensure that no entrapped air is
(releasing of in the pump which is the pump’s motor present in the pump before operating
chemical undesired  Entrapped air is present in the the pump
hazards)  Discharge valve fails to pump which is undesired as the  Ensure that the discharge valve is
close during pump entrapped air will accumulate at close before operating the pump
operation the pump suction point Provide proper training to the
inhibiting flow workers and operators to ensure safe
 May lead to negative pressure at and correct start up procedures are
the pump inlet, causing pump followed
failure  Regular or periodic checking and
 Cause damages in major and maintenance on pump
minor pipelines  Shut down the plant for annual
maintenance
Normal Rupture /
Operation Mechanical  Blockage of pump  May lead to backflow of lean  Install drain valve at the pump
damage suction and discharge amine solutions in the pump upstream to remove the
(releasing of  Failure of the pump  Cause damages in major and contaminants from the pipelines
chemical motor minor pipelines  Install two pump where one of it is
hazards)  Pump cavitation  Lean amine solution that are used for backup in case of
 Abnormal operating released from the heat malfunction of the main pump
conditions from the exchanger may pose thread to  Install check valve to prevent back
upstream process both humans and equipment as flow of the lean amine solution
Full Report
it is corrosive  Regular or periodic checking and

maintenance on pump
 Shut down the plant for annual
maintenance
 Install pressure indicator at pump
upstream to monitor the pressure of
the pump outlet
 First aid and fire response team in
case of emergencies
 Provide appropriate Personnel
Protection Equipment (PPE) for the
workers and operators
Maintenance Rupture /  Cause damages in major and
and Shut Mechanical  Improper or incorrect minor pipelines
Down damage shut down procedures  Lean amine solution that are  Establish proper shut down
(releasing of  Pump suction is fail to released from the heat procedures
chemical left open during shut exchanger may pose thread to  Ensure that the drain valve is open
hazards) down process both humans and equipment as and pump is completely de-
 Failure of drain valve it is corrosive pressurized
which results in the pump  Regular or periodic checking and
is not de-pressurized maintenance on pump
completely  First aid and fire response team in
 Failure of pump motor case of emergencies
 Provide appropriate Personnel
workers and operators
Shell and Start Up Thermal hazard
Tube Heat  Failure controller and  Lean amine solution that are  Establish proper start up procedures
Exchanger valves which results in released from the heat including a checklist for startup
(HX-402) improper cooling of lean exchanger may pose thread to process
Full Report
amine within the heat both humans and equipment as  Provide proper training to the
exchanger it is corrosive workers and operators to ensure safe
Rupture /  Improper cooling of the lean and correct start up procedures are
Mechanical  Improper or incorrect amine stream which will reduce followed
damage start up procedures the CO2 removal process in the  Regular or periodic checking and
(releasing of  Sudden high in-flow of absorption column maintenance on the heat exchanger
chemical lean amine solution  Injury and fatality  First aid and fire response team in
hazards) which lead to failure of  Fouling in the heat exchanger is case of emergencies
the heat exchanger excessive which might result in  Provide appropriate Personnel
 Failures of the control subsequent failure Protection Equipment (PPE) for the
valves at the inlet of tube  Downstream processes, workers and operators
side of the heat especially ammonia synthesis  Shut down the plant for annual
exchanger process, will be affected if CO2 maintenance
removal process in the
absorption column is inefficient
Normal Hazardous
Operation chemical  Release of lean amine  Lean amine solution will be  Regular or periodic checking and
exposures solution from the heat released which may cause maintenance on the heat exchanger
exchanger due to leakage damages to human and  First aid and fire response team in
or mechanical failure equipment as it is corrosive case of emergencies
 Utility stream which has been  Provide appropriate Personnel
Rupture / heated up may cause injuries Protection Equipment (PPE) for the
Mechanical  Abnormal operating (skin burns) to human upon workers and operators
damage conditions from the contacting  Shut down the plant for annual
(releasing of upstream process  Fouling in the heat exchanger is maintenance
chemical  Sudden high in-flow of excessive which might result in
hazards) pressurized lean amine subsequent failure
solution from the pump
 Blockage of the major
and minor pipelines
Full Report
Maintenance Thermal hazard  Lean amine solution will be  Establish proper shut down
and Shut  Failure controller and released which may cause procedures including a check-list for
Down valves which results in damages to human and shut down process
improper cooling of lean equipment as it is corrosive  Provide proper training to the
amine within the heat  Utility stream which has been workers and operators to ensure safe
exchanger heated up may cause injuries and correct shut down procedures
(skin burns) to human upon are followed
Hazardous contacting  Regular or periodic checking and
chemical  Release of lean amine  Affect the downstream process maintenance on the heat exchanger
exposures solution from the heat operations  First aid and fire response team in
exchanger due to leakage case of emergencies
or mechanical failure  Provide appropriate Personnel
Rupture /  Incorrect or improper workers and operators
Mechanical shut down procedures  Shut down the plant for annual
damage maintenance
(releasing of
chemical
hazards)
Full Report
7.1.7.7.2 Bow Tie Diagram for Hazardous Scenario in CO2 Removal Section
Full Report
7.1.7.8 Safety Assessment for Methanation Section
Operating
conditions
Methanator Start up Thermal hazard  Inappropriate start up  Skin burn due to direct  Imply proper start up
R-501 procedure contact with the high procedure guidelines
 Excessively high temperature electric heater  Provide proper
temperature transfer of or pipes safety/protective attire
process gas from electric  High temperature causes for workers
kettle catalyst degradation  Constant monitoring
during start-up
operation
 Install temperature
indicator at the outlet of
the electric kettle to so
that operator can
monitor the temperature
of syngas flowing
Normal Toxic hazard  Leakage of syngas in  Affect the production of  Install syngas leakage
Operation pipeline desired product detector
 Faulty pipelines  Leakage of syngas to  Install good ventilation
surroundings system to vent off the
 Affect the health of syngas away from the
workers who are exposed plant
Full Report
to syngas leaked  Accessible first aid kit

in the plant
Fire and  Leakage of syngas which  Burning and explosion  Install emergency stop
explosion contains methane in occur due to flammable button
hazard pipeline lead to ignition syngas  Install automated water
 Pressure build up in  Injury or fatality sprinkler to extinguish
reactor due to  Vessel rupture fire
malfunction pressure  Conduct regular
relief valves checking and
 Excessively high flow maintenance on pressure
rates of process gas in relief valves and control
reactor due to valves
malfunction of control  Install good ventilation
valves (overloading in system to relief
reactor) excessive pressure build
up in the reactor
 Install pressure alarms
 Use a steam drum to
store the incoming
syngas before sending to
the reactor
Mechanical and  Inappropriate selection  Pressure build-up occurs  Conduct regular

instrumentation of material and  Leakage of process gases inspection and
error instrumentation during maintenance
design stage  Good material selection
Full Report
 No maintenance during design stage

conducted
Thermal hazard  Higher carbon oxides  Skin burn or injury due to  Use insulation on the
flow rate than estimated direct contact with the reactor
during reactor design reactor  Provide appropriate
which lead to high Personnel Protection
temperature of the Equipment (PPE) for
reactor potential workers
exposed to hot reactor
Maintenance Fire and  Inappropriate shut  Damage to equipment  Provide proper shut
and shut explosion down and  Injury and fatality down guidelines and
down hazard maintenance procedure for operators
procedure  Shutdown and
 Ignition of leaked maintenance checklist
syngas with the  Purging the contents of
presence of fire the reactor using
ignition source nitrogen gas
 Vents not open to  First aid response team
remove contents of and fire fighter team in
reactor case of emergency
 Automated fire sprinkler
system
 Venting valves to vent
of gases in reactor
Full Report
Hazardous  Malfunction of  Affect the health of  Provide appropriate

chemical venting valves or workers who are exposed Personnel Protection
exposures pipeline damage to syngas leakage Equipment (PPE) for
during shutdown or potential workers
maintenance lead exposed to hot reactor
to leakages of  Install syngas leakage
hazardous syngas detector
to atmosphere  Accessible first aid kit
within the plant in the plant
Full Report
Operating
Equipment Hazard Possible Causes Possible outcomes Mitigation Measures
conditions
Heat Normal Thermal Hazard  High temperature of  Skin burn or injury due to  Placed the heat
exchanger operation equipment due to high direct contact with the exchanger at a safe
HX-501 temperature process heat exchanger distance in the plant
HX-502 fluid  Provide appropriate
 High carbon oxides Personnel Protection
flow rate into heat Equipment (PPE) for
exchanger potential workers
exposed to hot heat
exchanger
 First aid and emergency
response team
 Install automated water
sprinkler to extinguish
fire
 Install composition
transmitter and analyzer
Fire and  Leakage of syngas  Heat exchanger damage  Provide appropriate
explosion which contains  Release of high Personnel Protection
hazard methane in pipeline temperature of syngas to Equipment (PPE) for
lead to ignition the surroundings workers
 Pressure build up by  Injury or fatality  First aid and emergency
process fluid in the  Affect the production of response team
heat exchanger due to desired product  Install flow indicator
blockage in inlet or transmitter to detect
outlet
Full Report
 Continue feeding of high flow in the heat

syngas lead to exchanger
overloading in the heat  Install gas leakage
exchanger detector
Mechanical  Heat exchanger rupture  Flooding of syngas  Use strong material for
failure heat exchanger to
withstand high pressure
or stress
Toxic hazard  Leakage of syngas  Affect the health of  Provide appropriate
from heat exchanger to workers who are exposed Personnel Protection
surroundings to syngas leakage as Equipment (PPE) for
syngas contains workers
hazardous methane  Install gas leakage
detector
Maintenance Fire  Inappropriate shut  Affect the production of  Imply proper shut down
and shut down and maintenance desired product procedure guidelines
down procedure due to
ignition of flammable
syngas
Full Report
7.1.7.8.2 Bow Tie Diagram for Hazardous Scenario in Methanation Section

Full Report
7.1.7.9 Safety Assessment for Ammonia Synthesis Reaction Section
Operating
Equipment Hazards Causes Possible Outcomes Mitigation Measure
Conditions
Ammonia Start-Up Mechanical -Initiation with a bad batch of catalyst -Pressure drop increases -Regular inspection and
Synthesis Failure -Faulty preconditioning of catalyst -Temperature failures could maintenance
Reactor (R- -Larger catalyst size result in runaway reaction -Appropriate training for
601) -Catalyst fines produced during leading to fire or explosion operators
loading or poor loading -Injuries/loss of life as a -Corrosion-resistant materials,
Thermal and Fire All of the above could result in poor result of reactor blow up and/or adequate corrosion
Hazard selectivity and conversion achieved to -Plant has to be shut down allowances.
be lower than required standards after for repair works
start-up catalyst replacement.
-Maldistribution due to faulty flow
distribution design or plugging of
flow distributors with fine solids
-Rapid heating at reaction initiation
-Axial variation in temperatures
-Faulty inlet and exit flow distributor
-Setting of temperatures and pressures
are incorrect
-Transmitters are left in test mode
-Corrosion in pipework and reactor
vessel
Full Report
Normal Thermal and Fire -Unfavourable shift in equilibrium at -Rapid/gradual decline in -Use transmitters with remote
Operation Hazard operating temperature conversion seals
- Catalyst poison present in feed -Lost in catalyst activity -Institute procedures for
-Temperature sensor error or high -Sintered catalyst operation to inspect transmitters
temperature trip fails -Poisoned catalyst during routine rounds
-Electronic error in instrument or -Loss in surface area of -Earthing of electrical
controller for pressure and catalyst equipment
temperature -Reactor instability -Ensure no major ignition
-Impulse line leak/crimped - Thermal runaway can sources are placed nearby
-Sensor deformation occur because a runaway -Pressure-relief devices.
-Loss of seal fluid in transmitter exothermic reaction can -Fail-safe instrumentation
-Faulty feed and discharge port design have a range of results from -Provision of block valves on
- Leakage in reactor/valves the boiling over of the lines to main processing areas
-Feed temperature too high or exits reaction mass, to large - Install fire detection, alarm
threshold/ extraneous component increases in temperature and and control systems
reacts exothermically/instrument error pressure that lead to an - Proper insulation for pipeline
leading to temperature hotspots explosion. Such violence can transferring gases at high
Rupture/Explosion -Reaction is carried out at too low a cause blast and missile temperatures and pressures.
leading to release temperature which results in damage. The ammonia gas -Provide personnel with the
of chemical accumulation of reactants released could trigger a fire appropriate personal protective
hazards -Poor controller tuning or a secondary explosion. equipment (PPE) in accordance
-Contamination in feed (oxygenated Hot liquors and toxic to national codes and standards
compounds/sulphur) materials may contaminate
-Upstream process or equipment the workplace or generate a
upsets toxic cloud that may spread
-Fluctuations in feed from upstream off-site.
process
-Malfunction of in-line filters
Full Report
resulting in dust or corrosive products

being brought in from upstream
-Pipework rupture upstream of the
reactor
-Actuator sizing may be insufficient Potentially leads to the -Preventive maintenance
to actuate valve in emergency reaction vessel being at risk - Spare pressure relief valves
conditions from over-pressurisation due are installed as well as a bypass
-Actuator diaphragm ruptures or leaks to violent boiling or rapid vent valve
-Air line to actuator could be blocked gas generation. The elevated -Adequate, and secure, water
All of the above potential failures temperatures may initiate supplies for fire fighting.
could lead to relief valve malfunction secondary, more hazardous -Provision for access of
resulting in failure in purging runaways or decompositions. emergency vehicles and the
procedure -There can be serious risk of evacuation of personnel.
injuries, even death, to plant -Adequate separation of
operators, and the general hazardous equipment.
public and the local
environment may be
harmed. At best, a runaway
causes loss and disruption of
production.
Full Report
Maintenance Fire and explosion -Closed isolation valve -Inaccurate low reading with -Improve maintenance
and no response to process procedure and re-check to
Shutdown variations which could lead ensure that transmitter isolation
to hazards being left valve are returned to open state
unidentified resulting in after service or testing
major catastrophe such as -Consider redundancy with
reactor blow up each transmitter on separate
isolation valves with signal
comparison
Full Report
Operating
Equipment Hazards Causes Possible Outcomes Mitigation Measure
Conditions
Heat Start-Up Mechanical -Excessive clearance between - Fouling or scaling of -Impingement baffles are
Exchanger Failure baffles and tubes, high inlet gas heat exchanger which included at shell inlet nozzles to
(HX-601, velocities, surges in cooling potentially results in prevent erosion of tubes and
HX-602) water causes tube vibrations mechanical failure flow-induced vibration
resulting in noise - Affects all processes -Care must be taken to account
-Poor heat exchanger fabrication downstream for the larger heat exchange that
or faulty design results in - Injury and fatality occurs for clean tubes/surfaces as
unexpected corrosion, excessive the design was based on
pressure drop in heat exchangers reduced heat-transfer coefficients
-Bypass is left open that accounts for ultimate dirty
-Shell side is filled up first with film resistance
hot medium instead of tube side - Establish safe and proper start-
first with cold medium leading to up procedures, (e.g. provide a
overheating and pressure build- checklist for operators/workers
up during reactor start-ups)
Normal Thermal and Fire -Change in pH of coolant -Release of gases into -Pressure relief is provided to
Operation Hazard (water), high cooling water surrounding which may allow for system where block
temperatures, precipitation of cause severe skin burns, valves could isolate trapped
soluble compounds, presence of injury and/or fatality fluids
fungi or corrosion products -Fire -Ensure the air is vented
contribute to fouling - Affects all processes - Liquids being heated leaves at
-Lack of support for tube bundle, downstream the top of the exchanger
cavitation, improper tube -Plant shutdown to prevent the build-up of gases
finishing, vibrations, corrosion coming out of solution and vice
Rupture/Explosion and erosion all lead to leaks from versa for liquids with suspended
Full Report
leading to release the gasket at the tube sheet joint solids or viscous fluids
of chemical -Damaged insulation -Baffle windows are oriented to
hazards -Poor tuning of controller facilitate drainage
-Sensitivity to high flow rates, -Vents should be added to bleed
local turbulence with particles or off trapped gases
entrained gas bubbles resulting -Regular inspection and
in erosion of heat exchanger maintenance
material -Appropriate training for
-Fouling due to high service fluid operators
temperature -Corrosion protection
-Install fire detection, alarm and
control systems
-Provide personnel with the
appropriate personal protective
equipment (PPE) in accordance
with national codes and
standards
Maintenance Mechanical -Heat exchanger not designed for -Affects human health -Appropriate training for
and Failure leading to transient state resulting in and creates pollution to operators
Shutdown release of mechanical failure environment - Install fire detection, alarm and
chemical hazards - Injury/Fatality control systems
- Equipment failure
- Affects all processes
downstream
Full Report
7.1.7.9.2 Bow Tie Diagram for Hazardous Scenario in Ammonia Synthesis Reaction Section
353
Full Report
7.1.8 Plant Layout
7.1.8.1 Site Selection and Locations

The selection of plant location was done by considering a few criteria which includes
feedstock availability, plant accessibility, land cost, utilities availability, transportation
availability, water supply, availability of labour, effluent handling, environmental and social
impact. The chosen site location is (3o58’40.54’’ N, 101o06’12.72’’), with a total area of
83592m2 (387m× 216m) as shown in Figure 10.1 after considering all the factors above.
Figure 7.1.8.1.1 Location of Plant
The reasons of selection of this site can be summarized as follows:
1. The raw material for manufacturing of ammonia is readily available from the
plantation sites nearby.
2. Access roads are available for transportations of raw materials and manufactured
goods.
3. An area of low population density so its activities have minimum impact on the
neighborhood.
354
Full Report
7.1.8.2 Layout Approaches

Plant layout refers to the arrangement and allocation of industrial facilities which takes into
account of efficiency, economics and safety. Plant layout plays an important role in loss
prevention, control of hazard and improves process efficiency. A well designed plant layout
involves efficient utilization of spaces available without compromising safety.
Process unit plot plans can be categorized into two types, namely Structure-Mouted or Grade-
mouted Horizontal Inline arrangement (Bausbacher and Hunt, 1993). The Grade-mouted
Horizontal Inline arrangement is selected as it provides an easier plant construction and also
more convenience for operation and maintenance. Several factors, as shown below, are
required to be considered when designing the plant layout:
7.1.8.2.1 Economic Considerations

The construction cost should be as minimal as possible by adopting a plant layout that is able
to maximize space utilization to save the land cost. Each sections of the plant are arranged as
close as possible according to a safety distance margin. Plant layout that requires the least
amount of steel work and shortest run of pipe connection is also preferred.
7.1.8.2.2 Process Requirements

Process requirements involve the arrangement of each section to ensure the smooth flow of
operation. The position and elevation of equipment need to be designed accordingly to
prevent loss of energy and wastage. Safety margins were allocated between each equipment
for safety purpose. Besides the main process units, space is also allocated for utilities and
storage.
7.1.8.2.3 Convenience of Operation and Maintenance

Sufficient working space must be allocated to provide access to the equipment. Equipment
that requires frequent operator attention such as sample points is located at convenient
positions and heights. For ease of maintenance such as changing tubes of heat exchanger and
replacement of catalyst in reactor, sufficient space is provided. Equipment like pumps and
compressors that may need to be dismantled is placed under cover.
7.1.8.2.4 Future Expansion

20% extra spaces are allocated in each section of plant layout for future expansion purpose.
Equipment is arranged in such a way that it can be conveniently connected in future
expansion. For future requirements, service pipes are over-sized and empty space are
allocated on pipe racks.
355
Full Report
7.1.8.3 Site Constraints
7.1.8.3.1 Wind Direction

The location of Langkap is susceptible to two monsoon regimes, namely, the Southwest
Monsoon from late May to September, and the Northeast Monsoon from November to March.
The Northeast Monsoon brings heavy rainfall whereas the Southwest Monsoon normally
signifies relatively drier weather (Malaysian Meteorological Department, 2013). The location
of flare is heavily affected by the direction of wind. Hence, the flare is constructed at the
northwest edge of the site. This location is situated far away from the office building at the
opposite end of the plant. This helps in increasing the inherent safety of the plant layout. The
location of the flare also prevents the undesirable fumigation of the industrial site due to
smoke from flaring.
7.1.8.3.2 Separation Distances

Despite the notion of using the shortest possible connection pipes between equipments, most
equipment requires safety distance between each other. Minimum distances must be given
between certain equipments to ensure safe operation. As a rule of thumb, the inter- and intra-
units separations are shown in the table below (GE GAP, 2001):
Table 7.1.8.3.2.1: Inter-unit spacing recommendation (in meters)
Pressure
Process Utilities Cooling Control Compressor Fire
Storage Flare
Units * Areas Towers Rooms Rooms Stations
Tanks
Process
16-61
Units*
Utilities
16-61
Areas
Cooling
16-61 31
Towers
Control
16-61 31 31
Rooms
Compressor
16-61 31 31 31
Rooms
Pressure
Storage 16-61 107 107 107 107
Tanks
Flare 16-61 92 92 92 92 122
Fire
16-61 16 16 16 61 107 92
Stations
*Separation distance may vary according to the hazardous level of the processing units.
356
Full Report
Table 7.1.8.3.2.2: Intra-unit spacing recommendation (in meters)
Heat
Reactors* Columns Compressors
Exchangers
Reactors* 5-8
Columns 5-16 5
Compressors 5-16 16 10
Heat
4-8 4 10 2
Exchangers
*Separation distance may vary according to the hazardous level of the reactions.
7.1.8.3.3 Access Road

Roads must be constructed in a way that transportation of materials within the plant can be
done with ease. Besides, the road constructed must provide the accessibility for maintenance
purposes. Hence, extra area must be allocated for each section for maintenance purposes. The
road must be wide enough to allow movements of trucks and machineries in the plant.
7.1.8.3.4 Processing Sequence

As mentioned earlier, to reduce the amount of pipe works required, the plant is arranged in a
way that the subsequent section is located immediate next to each other. Output from each
section can be transferred to the subsequent section in the shortest time. This can ensure the
pressure drop and heat loss across equipments are minimised. This helps in reducing overall
power consumption of the plant.
7.1.8.3.5 Noise Abatement

Some equipment that will produce loud noises are grouped together and placed in a common
room. Compressors will be placed in the same compressor room due to its load noise during
operation. Pumps, however, are only needed in small quantities and thus are not economically
feasible to be grouped together in a common room. It is placed together with the processing
unit. The noise can be compensated by the lower head required for the pump due to shorter
distances.
7.1.8.3.5 Minor Equipment

Since not all minor equipment was designed in detail, the average dimensions of similar
equipment were used as guidance. Therefore, heat exchangers and condensers are taken to be
2.5m width × 5m length while pumps and compressors are taken to be 1m width × 1m length
and 2m width × 3m length respectively.
357
Full Report
7.1.8.4 Size of Equipments

To estimate the area required for each sections, similar equipments are assume to have the
same dimension. The height is not being considered in this section. The estimated dimensions
and area for each unit are shown in the table below:
Table 7.1.8.4.1: Summary of the estimated size of each equipment
Equipments Length (m) Width (m) Diameter (m) Area (m2)

Shredder 6 3 - 18.00
Dryer 4 2.5 - 10.00
Circulating fluidised-
- - 2 3.14
bed gasifier
Blower 1 1 - 1.00
Scrubber 2 - 1.5 1.77
Bag Filter 1.5 - 0.5 0.20
Autothermal Reformer - - 1.96 3.02
High Temperature Shift
- - 1 0.79
Reactor
Low Temperature Shift
- - 1 0.79
Reactor
Flash Vaporiser - - 1.068 0.90
Synthesis Reactor - - 2.05 3.30
Separator - - 1 0.79
Fired Heater - - 1.96 3.02
Storage Tank - - 5 19.63
Absorption Column - - 1.76 2.43
Stripping Column - - 1.76 2.43
Methanator - - 1 0.79
Waste Heat Boiler 6 5 - 30
Heat Exchanger 5 2.5 - 12.5
Compressor 3 2 - 6
Pump 0.5 0.5 - 0.25
Reboiler 5 2.5 - 12.5
Condenser 5 2.5 - 12.5
358
Full Report
7.1.8.5 Site Allocation for Process-Related Site
7.1.8.5.1 Pre-Treatment Section

Pre-treatment section consists of shredder and dryer. This section is placed near to the
secondary entrance and tank farm for the convenience of transporting raw materials and
chemicals. Hence, the area allocated for pre-treatment section is allowing for
safety distances and future expansion as well as after taking into consideration of spaces
required to store palm oil trunk.
7.1.8.5.2 Gasifying Section

The gasifying section consists of following units:
Table 7.1.8.5.2.1: Summary of the area required for gasifying section
Equipment Area (m2) Quantity Total Area (m2)

Circulating fluidized-bed gasifier 3.14 1 3.14
Blower 1.00 1 1.00
Scrubber 1.77 1 1.77
Bag Filter 0.20 1 0.20
Total 6.11
Hence, the area allocated for gasifying section is allowing for safety distances
and future expansion.
7.1.8.5.3 Reforming Section

The reforming section consists of following units:
Table 7.1.8.5.3.1: Summary of the area required for reforming section

Autothermal Reformer 3.02 1 3.02
Fired Heater 3.02 1 3.02
Waste Heat Boiler 30 1 30
Heat Exchanger 12.5 1 12.5
Pump 0.25 1 0.25
Total 48.79
Hence, the area allocated for reforming section is allowing for safety distances
359
Full Report
7.1.8.5.4 Water-Gas Shift Section

The water-gas shift section consists of following units:
Table 7.1.8.5.4.1: Summary of the area required for water-gas shift section

High Temperature Water Gas Shift Reactor 0.79 1 0.79
Low Temperature Water Gas Shift Reactor 0.79 1 0.79
Waste Heat Boiler 30 1 30
Pump 0.25 1 0.25
Total 44.33
Hence, the area allocated for reforming section is allowing for safety distances
7.1.8.5.5 CO2 Removal

The CO2 removal section consists of following units:
Table 7.1.8.5.5.1: Summary of the area required for CO2 removal section

Separator 0.79 1 0.79
Absorption Column 2.43 1 2.43
Stripping Column 2.43 1 2.43
Storage Tank 19.63 1 19.63
Reboiler 12.5 1 12.5
Condenser 12.5 1 12.5
Total 87.78
Hence, the area allocated for CO2 removal section is allowing for safety
distances and future expansion.
360
Full Report
7.1.8.5.6 Methanation Section

The methanation section consists of following units:
Table 7.1.8.5.6.1: Summary of the area required for methanation section

Heat Exchanger 12.5 2 25
Cryogenic Purifier 12.57 1 12.57
Methanator 0.79 1 0.79
Total 39.15
Hence, the area allocated for methanation section is allowing for safety
7.1.8.5.7 Synthesis Section

The synthesis section consists of following units:
Table 7.1.8.5.7.1: Summary of the area required for synthesis section

Synthesis Reactor 3.30 1 3.30
Total 18.80
Hence, the area allocated for synthesis section is allowing for safety distances
7.1.8.5.8 Ammonia Purification Section

The refrigeration section consists of following units:
Table 7.1.8.5.8.1: Summary of the area required for ammonia purification section

Flash Vaporizer 0.90 1 0.90
Storage Tank 19.63 1 19.63
Total 84.61
361
Full Report
Hence, the area allocated for ammonia purification section is allowing for safety
7.1.8.5.9 Compressor Room

The compressor room consists of following units:
Table 7.1.8.5.9.1: Summary of the area required for compressor room

Compressor 6 15 90
Inter-stage Heat Exchanger 12.5 8 100
Total 190
Hence, the area allocated for compressor room is allowing for safety distances
7.1.8.6 Site Allocation for Non- Process-Related Site
7.1.8.6.1 Security House

There are 2 entrances that give access into the plant. The effect may be catastrophic if the
plant was compromised. Hence, security is very important. Guards are required to ensure
only authorized personnel are able to enter and operate the plants. A security house is
constructed on every entrance. The size of each security house is .
7.1.8.6.2 Loading Bay

This is the place where the palm pressed fibre purchased from the palm oil mill will be
unloaded. The area allocated to unloading bay is . The location of
loading bay is constructed close to the silo where the palm pressed fibre will be stored.
7.1.8.6.3 Car Park

The parking for this plant is situated in front of the administration building. 100 parking lots
are available. 1 car park lot is estimated to be . Thus, considering the extra area
required for the movement of vehicles across the car park, the total area allocated for car park
is
362
Full Report
7.1.8.6.4 Control Room

It is estimated that 2 people will be working in the control room for each shift. Considering
all the equipments and workspace, the total area allocated to the control room is
.
7.1.8.6.5 Administration Building

The administration office must be constructed far away from the processing plant for safety
reasons. Besides, some of the processing units generate high amount of heat and noise and
may create an uncomfortable working environment. will be allocated for each staff
in this building. An estimate of 30 staffs will be working in this building. It is a 3-storey
building, hence the overall area occupied by this section is .
7.1.8.6.6 Cafeteria
The cafeteria must also be constructed some distance away from the processing units. This is
to reduce the possibility of contamination of food and water. Consuming industrially
contaminated food or drinks can be fatal. The cafeteria must be able to accommodate all
employees in 1 shift. Assuming 1 person will occupy . is
allocated for the cafeteria.
7.1.8.6.7 First Aid Room

A small area is allocated to build a first aid room to provide immediate treatment for injured
workers. is sufficient to house all the medical equipments and 2 beds to handle
emergency situation.
7.1.8.6.8 Laboratory
The laboratory is used for R&D and quality control. The laboratory is where samples of
biogas would be tested to check if it meets the desired standard. Considering the number of
laboratory equipments, the area allocated for the laboratory is .
7.1.8.6.9 Warehouse
A warehouse is used to store spare equipments and chemical. An area of
is allowed for this area.
7.1.8.6.10 Workshop
The workshop is used for maintenance of equipments. is allowed for this
area to ensure the workshop is spacious enough for the maintenance of equipments.
363
Full Report
7.1.8.6.11 Fire Fighting Station with Emergency Water

Access is very important for fire fighting. Emergency water is available immediate next to
the fire fighting section. Area of are available for
the fire fighting station and emergency water respectively.
7.1.8.6.12 Electrical Room

An electrical room is present in the plant to distribute power throughout the plant. An area of
is allowed for electrical room.
7.1.8.6.13 Emergency Assembly Area

The emergency assembly area is the place to assemble all the employees when emergency
happens. Hence, its location must be close to the place where most personnel are located at.
The area allocated for the assembly area is . The area is large enough
to accommodate the entire workforce during emergency.
7.1.8.6.14 Cooling Tower

Cooling tower is used to cool the cooling water that has been heated up in heat exchanger.
The diameter of the cooling tower is 3m. Hence, the area for cooling tower is 7.07 .
7.1.8.6.15 Tank Farm

Storage tank is used to store products of the plant. There are 7 storage tanks available for this
plant. The area allocated for the tank farm is . The diameter for each
storage tank is . Each storage tank is 1 diameter apart from each other for safety reason.
7.1.8.6.16 Wastewater Pond

The wastewater pond is used to collect the wastewater generated by the plant. It can then be
sent for further treatment before being released into the environment. The size of the
wastewater pond is with diameter of 2m.
7.1.8.6.17 Steam Drum

Steam drum is used to store the superheated steam generated by waste heat boiler. The
superheated steam is collected at steam drum at 30bar and . The size of the steam drum
is .
7.1.8.6.18 Flare
Flare stack is located at the northwest edge of the plant. This location is situated far away
from the office building at the opposite end of the plant. The diameter of the flare stack is 1m.
The total area allowed for the flare stack is 0.79 .
364
Full Report
7.1.8.7 Drawing of Plant Layout

365
Full Report
7.2 Design Limitation
Table 7.2.1 Explanation on the three types of design limitations
Aspects Limitations Description

Redundancy Increase operating cost, leading to longer payback  Auxiliary equipment are
period and lower NPV generated duplicated to increase the
reliability and efficiency of the
plant
 More heat exchangers are
needed to maintain the
maintain the temperature at
required condition
 More pumps and compressors
are needed to fulfil the required
duty of the process
 More separators are needed to
filter out the unwanted
substances in the process
stream, thus, yielding higher
purity
 Instruments of the control
systems are duplicated to
ensure safe operation of the
plant
Technology Internal Circulating dual fluidized bed gasifier  CFB required biomass size of
(CFB): less than 20mm to operate
 Pre-treatment requirement of the biomass feedstock  High moisture content feedstock
 Tolerant to fluctuations and high moisture content of decrease the efficiency of the
feedstock gasifier (recommended range of
 High particulate level in syngas feed moisture contents: 10~15%
of biomass feedstock)
 High amount of particulates
(from the suspended bed
material, ash and soot) due to
unconverted components of
biomass feedstock, small amount
of tar and fly ash also present in
the syngas evolved after the
gasification process
366
Full Report
Autothermal Reformer (ATR):  additional oxygen is required to

 Requires additional oxygen input be fed into the ATR to
 Lower hydrogen yield compared to Steam methane complement the combustion
reformer reaction
 Limited commercial experience  ATR is new technology
operating based on steam
reforming and partial oxidation
principles. The concept and
viability of ATR have been
researched and discussed for
many years
CO2 Absorption Process using a-MDEA solvent:  The amine solvent promoter,
 Piperazine content in the solvent is highly corrosive piperazine is highly corrosive
 Solvent cannot be fully regenerated which might affect the
performance and lifespan of the
absorption column and processes
in the CO2 removal unit
Methanation Process:  Methanation consumes part of

 Hydrogen consumption the H2 generated in the process,
 Production of additional inert gas lowering its percentage recovery
 Operates at high temperature and pressure  Increase the safety risk and
utility cost
Vertical multi-bed radial flow converter:  Increase the safety risk and
 Operates at high pressure utility cost
 Increase compressor duty
Environment Site climate condition:  Yearly maintenance of protective

 The humidity levels of Perak hover around 70% to coating on vessels are required to
90% prevent external corrosion of the
 The average rainfall is 187 days/ year at 4.4in/hr vessel wall
(the driest season is expected in June and July)  Sufficient lightning and extra
 The wind velocity is about 125.5km/h caution in preparing a risk
assessment for hazards related to
working in a raining
environment; Safety seminars/
trainings are organized monthly
to train workers and contractors
to identify hazards in the
chemical plant regardless of
climate conditions
 Water reservoirs are used to store
367
Full Report
water on normal days for use in

dry seasons
 All equipment should be safely
grounded
 The deflection of a column and
wind-induced vibrations were
considered in the design through
extra structural steel work and
supports as premature failure of
the vessel by fatigue is possible if
the vortex shedding from tall thin
column produces vibrations with
frequency that matches that of
the natural frequency of the
column
Regulatory limits  High amount of particulates, ash,

 The point of discharge or mixed gases should be in dust, tar, CH4, C2H4, C3H6 and
compliance with the standards specifying the C2H6 (from the suspended bed
maximum permissible concentrations of any matter material, ash and soot) emitted to
that may be present in or discharge into the the atmosphere after the
atmosphere as shown below: gasification process was
monitored through post
treatment whereby a series of
cyclones are used to remove
particulates while thermal
cracking is used to remove tar
 A water treatment plant which is
out of the system boundary for
this project will be implemented
as tar condensate and water are
released to the environment at
the syngas cleaning stage, plus
additional liquid effluent are
emitted from the methanation
 The discharge of industrial effluent or mixed process
effluent should be in compliance with Standards A
and B stated in the Fifth Schedule of the
Environment Quality Act 1974
Environmental impact assessments  The emissions that may occur

over time due to corrosion of
equipment, erosion of equipment
material by flowing of water or
other chemicals or gases are
unpredictable as corrosion is a
complex phenomenon and is
368
Full Report
normally based on experience

 Fugitive and abnormal emissions
as well as spillage were excluded
in the life cycle analysis as there
is a lack of data
 The impacts during abnormal
operations, start-up and shut-
down, cleaning and maintenance
activities were omitted as there
are no available data on the
emissions during these steady
state operation
 Impacts from the construction of
the plant, manufacturing of the
catalysts, absorbents, and
refrigerants will add
disproportionately to the time of
the study, thus, they are excluded
369
Full Report
CHAPTER 8 | ECONOMIC PERFORMANCE
8.1 Introduction
The purpose of this section is to provide a contemporary estimate of current and
future costs of a wide range of aspects in the plant that is produce anhydrous fertilizer-grade
ammonia with low carbon footprint not exceeding using
oil palm trunks (OPT) as a feed stock of Alternis BioAmmonia. During normal operation and
capacity utilization of that plant, it is assumed that all anhydrous ammonia produced is fully
sold each year. This plant uses about minimum OPT feedstock of approximately 65 kilo
tonnes on annual basis by operating 300 days a years for a total operating life of 20 years.
Although ammonia fertilizer production has been commercialized in large scales worldwide,
the production of anhydrous fertilizer-grade ammonia on a low carbon footprint is currently
the focus of Alternis BioAmmonia. In this section, a market evaluation on the product along
with a detailed cost estimation comprising of capital and operating costs will be conducted, in
which a cash flow analysis will be used to evaluate the sales revenue and total costs of the
plant. The net present value (NPV) and discounted cash-flow rate of return are also included
to determine the profitability and economic viability of this investment. In addition, this
section further considers the robustness of the economics of the plant by conducting
sensitivity analysis on raw materials cost and product selling price cost associated with
production in current and future market scenarios.
8.2 Market Evaluation of Anhydrous Fertilizer Grade Ammonia

8.2.1 Current Global Market Size and Demand of Anhydrous Fertilizer Grade Ammonia
Ammonia is one of the world’s most produced chemicals and the utilization of ammonia
varies globally according to different sectors. Figure below depicts the global usage of
ammonia in 2010-2012, nitrogen fertilizer consumption accounts for more than three-quarters
of the world ammonia market, with more than half of those going to urea production. The
remaining of ammonia produced is used for non-fertilizer applications which include
refrigerant, cleaning and textile finishing.
370
Full Report
Figure 8.2.1.1: World Ammonia Usage
The world consumption of ammonia has been reported to have an annual growth of 2.3%
from 2005 to 2010, where FIGURE(b) shows that the Asia-Pacific region accounted for
approximately 58.7% share of global demand in the year 2010 (Albany, 2013). The main
drivers of growth for fertilizers include biofuels, food and nutrition security, environmental
concerns, and organic production (AAFC, 2008). Due to growing world population and
declining amount of arable land the market for fertilizers are expected to continue grow in the
future. The population growth is more evident in the Asia-Pacific region as India and China
promise substantial consumption potential by leading a trend among emerging countries
seeking to become self-sufficient in terms of food production.
Governments of developing countries are also seeking to provide food security by
increasing crop production in their nations due to their lack of ability to afford extensive
exports, leading to forecasted increase in global ammonia capacity of about 35 million tons in
these regions. In light of this, the demand from the Asia-Pacific region is set to continue to
drive future global demand, where the global demand for ammonia is expected to have an
annual growth of 2.7% and reach about 160 million tons in 2020 (almost twice the demand of
96.5 million in the year 2000), thus bringing about a forecasted revenue of $102billion in the
year 2020 (Albany, 2013; Schulze, 2012). On a local front, Malaysia is expected to expand
the use of palm oils in biodiesel production, indicating that the local demand for fertilizers is
projected to grow as well.
371
Full Report
8.2.2 Selling Price Estimation and Forecasting
Figure 8.2.2.1: Recent trend in selling price of ammonia
Figure 8.2.2.1 shows that the recent trend of world fertilizer prices were the highest in 2008,
with anhydrous ammonia prices hovering at about . In year 2009 when the
recession hit, the price in 2009 dropped down to USD204/t and then picked up slowly and
increased since then. The prices of fertilizer are expected to remain high due to limited
ammonia manufacturing capacity that restricted increases in supply while nitrogen fertilizer
use continue to increase as shown in figure above. The competitions with other similar
industries are uptight therefore it is viable to maintain the quality and purity of ammonia and
also sold at a reasonable price. Alternis BioAmmonia has standardized the price of ammonia
to be sold at USD898/t or RM 2400/t by basing on the more recent market price.
8.2.3 Main Cost Drivers

Production cost includes raw materials, utilities, maintenance, capital cost and labor. For
Ammonia production the cost of raw materials takes up averagely 75 % of the total
production cost. The price of natural gas currently at RM600/t indirectly affects the ammonia
pricing hence production cost fluctuates based on the current market trend. Capital cost will
contribute to the production cost. The labor cost will differ depending on the location and
country of the plant. The transportation of product and raw materials from the production
plant to supplier or vice versa will incur a cost on the production. The cost for storage will
depends mainly on the period of storage.
372
Full Report
8.2.4 Product Quality Requirement

The standard for fertilizer grade ammonia required to be achieved by Alternis BioAmmonia
is at the range of 99.5%-99.8% NH3 wt. The current production manages to achieve a
maximum purity of 99.8%.
8.2.5 Means of Supply

The most common method of delivery is through pipelines. These pipe length range from
meters to several kilometers and can be places on the surface of the ground or underground
depending on the condition of the area. Frequent maintenance and leak detector can be
installed to ensure no leakage occur to avoid any wastage or pollution. This is one of the
cheapest methods of supply if compared to truck delivery. Ammonia can also be transported
in cargo gas tanks. However the design of these tanks must be at adequate design pressure in
order to withstand the high pressure of Ammonia.
8.3 Capital Cost Estimation

Capital cost estimation can basically be divided into four main sections: land, fixed capital,
working capital and start-up capital. However, in the proposed project, capital spent on
land was considered to be negligible compared to other costs.
8.3.1 Key Assumptions and Parameters

In analysis of the project profitability, key assumptions that are being used are listed below:
 The construction period of the plant is 1 year and the operating life of the plant is 25
years
 The plant operates 300 days per annum
 The plant runs at 80% capacity on the first year and operates at full capacity of 100%
on the subsequent year
 The inflation rate is utilized in calculation of purchased cost
 Inflation factor is estimated by using the Chemical Engineering Plant Cost Index
(CEPCI)
 Discount rate of 15% is adopted in the evaluation of net present value (NPV)
 Corporate tax of 25% is applied in the generation of cash flow diagram
373
Full Report
8.3.2 Inside Battery Limits (IBL) Investment

Inside battery limits investments are mainly comprise of the equipment costs and the
installation cost. Purchased costs of the equipment are estimated based on the constant
provided in Table 6.6 Chemical Engineering Design by Sinnott, R.K. and Towler, G (2009).
Another source to determine the equipment cost is by interpolating the graph provided in
Plant Design and Economics for Chemical Engineers by Peters, M. Timmerhaus, K. and
West, R (2003). In addition, installed cost will be taken into consideration in the calculation
of inside battery limit. The typical factors that will be included in the calculations of installed
cost are summarized in Table 8.3.2.1.
Table 8.3.2.1 Installation Factors for Purchased Cost Calculation
Installation Factor Symbol Factor
Equipment erection 0.3
Piping 0.8
Instrumentation and control 0.3
Electrical 0.2
Civil 0.3
Structures and building 0.2
Lagging and paint 0.1
For better estimation, the material factor will be taken into account which will be based on
the type of the material used for the equipment. The material factors used in the calculation
are tabulated in Table 8.3.2.2.
Table 8.3.2.2 Material Factors
Material
Carbon Steel 1.0
304 Stainless Steel 1.3
316 Stainless Steel 1.3
With all these factors, the estimated installed cost of the equipment can be calculated by
using the equation below:
Due to the fact that the equipment cost obtained from literature sources are relative to the
reference year. Thus, inflation ratio of the current year relative to the reference year will be
374
Full Report
calculated to estimate the exact purchased cost of the year 2014. The index for the year 2014
is determined by extrapolation of the current trend.
Furthermore, the location factor taken is depends the base location where the equipment cost
data obtained. As the equipment cost data used are based on US, the location factor of 1.12
will be applied to South East Asia. The predicted exchange rate for USD in the year of 2014
is at 3.07 per USD. However, the exchange rate for USD in the year 2003 is at 2.63 per USD.
Thus, the calculation of location factor in the year 2014 is:
The location factor is determined to be 1.31 in the year of 2014. For detailed calculation of
the location factor, purchased costs as well as the installed cost are shown in Appendix D.
From the calculation, the total installed cost for all the equipment is RM 48395446.56. The
summary of the installed cost for all equipment is shown in Appendix D.
8.3.3 Outside Battery Limit (OBL)

Outside battery limits comprise of infrastructures in the plant except the main processing
plant. It can be categorized into 3 categories as described below:
 Facilities for the storage of raw materials and products
 Minor plant that is generating utilities for plant usage, such as steam boiler and
cooling tower
 Infrastructures and administration facilities such as warehouses, laboratories,
workshops and offices
The outside battery limit can be estimated by taking a percentage of the total inside battery
limit investment. This allowable range of percentage for outside battery limits 10% to 100%
of the inside battery limit. The allocations of the OBL in this plant are summarized in Table
8.3.3.1.
Table 8.3.3.1 Allocations of OBL in the Plant
Categories Percentage of IBL (%) Cost (RM MILLION)
Storage of raw material and product 10 4.84
Utilities generation unit 15 7.26
Infrastructures and administration facilities 10 4.84
375
Full Report
8.3.4 Engineering Costs and Contingency Charges

Engineering costs generally includes all the contractor charges, home office costs, costs
associated with detailed design and other engineering that is essential to carry out the project.
Contingency charges are the extra costs that have not undergone any of the previous
categories, and that should be added to the project budget to allow for variations in cost
estimates. Generally, engineering costs and contingency charges are each taken as a
minimum of 10% of the ISBL (Sinnot & Towler, 2009). Thus, the total engineering costs
and contingency charges were calculated to be $9.68 million.
8.3.5 Total Fixed Capital Cost

Fixed capital investment include outside battery limits (OBL) and inside battery limits (IBL).
‘Battery limits’ is the boundary of the processing plant which converts raw material to
finished product. The table below shows the summary of the capital cost of this plant. The
summary of fixed capital cost can summarized in Table 8.3.5.1
Table 8.3.5.1 Summary of Fixed Capital Cost
Category Unit Value
Plant Capacity ton/year 30000
Capacity Utilization % 100
No. of operating days Day 300
IBL RM million 48.40
OBL RM million 16.94
Total Engineering and 4.84
RM million
Supervision
Total Contingency Charges RM million 4.84
Total Fixed Capital Cost RM million 75.01
8.3.6 Start-Up Capital

Start-up costs include all the non-recurring costs between the completion of plant
construction and the commencement of successful plant operation. Appendix D summarizes
all the start-up capital associated in the plant, where the total start-up cost was found to be
$1.182 million.
376
Full Report
8.4 Operating Cost Estimation

Operating cost can be divided into manufacturing and non-manufacturing costs. This cost
include all costs related to production, distribution and marketing of products, together with
the ongoing costs of developing or purchasing the necessary technology. It also includes all
the management and business incurred indirectly in making and selling products. The amount
of raw material (Oil Palm Trunk) and utilities used were calculated in Chapter 5 Mass and
Energy Balance (Submission B). The unit cost of the materials and utilities were found from
various sources with application of the location factor and inflation ratio. The following
shows the cost used in operating cost estimation:
 The price of OPT is RM250/ton that will be supplied by Benta Plantation Sdn.
Bhd., United Plantation Sdn. Bhd., Southern Perak Plantation Sdn. Bhd., and
FELDA Besout oil palm plantations.
 The price of the process water/demineralized water for the waste heat boiler is
RM 3.30/m3.
 The electricity cost is RM0.29/kWh
Steam from steam drum was not added into the operating cost as it is produced from
the waste heat boiler. Cooling water from cooling tower and refrigerant were excluded from
the operating cost also as it will only be fed once into the equipment. This will be added to
the startup cost in the cash flow analysis.
Operators will be employed to operate and monitor the plant during normal operations,
start up and shut down operations, maintenance and also abnormal operations. There are 4
shifts per day and 16 operators will be on duty per shift. Furthermore, 35% of the wages is
reserved for payroll overheads and 50% of the labour cost will be reserved for plant overhead.
In addition, 5% of the fixed capital is maintenance and 1% of fixed capital cost is reserved for
insurance and tax respectively. Moreover, for non-manufacturing, the cost is calculated by
taking 3% of production cost for corporate administration. The operating cost worksheet is
shown as below:
377
Full Report
Table 8.4.1: Estimation of Operating Cost at 100% capacity utilization
Product Anhydrous Ammonia (aNH3)
Production Hour 7200 hr/yr
Product Route Plant Capacity 30000 t aNH3/yr
Capacity Utilization 100 %
IBL 48.40 RM Million
OBL 16.94 RM Million
Fixed Capital Contingency Charge 4.84 RM Million
Engineering Costs 4.84 RM Million
Total 75.01 RM Million
Product Selling Price 2600 RM/ t aNH3
Cost of Initial Catalyst Charge 7.76 RM Million
PRODUCTION COST
MANUFACTURING COST
Annual Cost Cost per Tonne
Raw Material Unit Usage, unit/yr Unit Cost, RM/unit (RM Million) (RM/t aNH3)
Oil Palm Trunk (OPT) 94937.00 250.00 23.73 791.14
Demin. Water (m3) 882511.20 3.30 2.91 97.08
Total Raw Material Cost 26.65 888.22
Annual Cost (RM Cost per Tonne (RM/t
Utilities Unit Usage, unit/yr Unit Cost, RM/unit Million) aNH3)
Electricity (kWh) 76706897.00 0.29 22.25 741.51
Natural Gas (t) 2973.60 600.00 1.78 59.47
MDEA&P MakeUp(t) 1306.02 3600.00 4.70 156.72
Total Utility Cost 28.73 957.71
Total Variable Cost (Raw Material+Utilities) 55.38 1845.92
Annual Annual Cost (RM Cost per Tonne (RM/t

Process Labour Number sal.,RM/yr/op Million) aNH3)
Operators 16
No of Shift Teams 4
Total Operators 64 30000.00 1.92 64.00
Payroll Overheads 30% of Op. wages 9000.00 0.58 19.20
Total Wages 2.50 83.20
Royalties 5 % of Fixed C 3.75 125.02
Maintenance 5 % of Fixed C 3.75 125.02
Plant Overheads 50 % of Labor 2.16 72.11
Insurance 1 % of Fixed C 0.75 25.00
Property Taxes 1 % of Fixed C 0.75 25.00
Total Fixed Operating Cost 13.66 455.36
Fixed 13.66 455.36
TOTAL PRODUCTION COST Variable 55.38 1845.92
Total 69.04 2301.29
NON-MANUFACTURING COST
Corporate Administration 2 % of Prod C 1.38 46.03
Total Non-manufacturing Cost 1.38 46.03
Total Operating Cost (Production Cost + Non-Manufacturing Cost) 70.42 2347.31
378
Full Report
8.5 Working Capital Estimation

Working capital is defined as the funds over and above fixed capital, start-up capital and land
investment needed to start and maintain a project (Brennan, 1998a). This includes raw
material and finished product inventories, material in progress inventories and the balance
between account receivables (debtors) and account payables (creditors).
The following are the assumptions and considerations were made in obtaining the working
capital:
a) Raw material costs are evaluated at the purchased costs. Feedstock inventory depends
on the source of raw material, transportation mode, process technology and its
reliability of supply.
i. OPT are considered to be a waste from oil palm industry, thus it has a low
selling price in Malaysia. The purchased price of OPT was found to be within
range of RM 250/tonne.
ii. Demineralized water is estimated to be purchased at RM 3.30/m3
b) The inventory period for both OPT was taken as 3 weeks following the bulk
commodities supplied (Brennan, 1998a). This period was taken by assuming a reliable
supply of OPT that reduces storage space and risk of degradation. Even though
demineralized water is supplied through a pipeline, storage of the water is necessary
thus an inventory period is accounted for 1.5 weeks.
c) Finished product stocks include stocks at the production plant. Products produce in
the plant are anhydrous fertilizer grade ammonia and carbon dioxide. Inventory period
for ammonia and carbon dioxide was taken as 2 weeks following the bulk
commodities supplied daily (Brennan, 1998a).
d) Material in progress inventory was assumed to be negligible
e) Both debtors and creditors period was taken as 6 weeks (Brennan, 1998a).
f) Capacity utilization was assumed to be 80% in first year of operation, 90% in second
year, and 100% from third year onwards.
379
Full Report
Table 8.5.1: Working Capital
Period (Week) Cost, RM Million Cost, RM Million
Raw Material
Oil Palm Trunk (OPT) 3 =(3*7/300)*23.73 1.661
Demineralized water 1.5 =(1.5*7/300)*2.912 0.102
Finished Products
Anhydrous Ammonia 2 =(2*7/300)*70.42 3.286
Carbon Dioxide 2 =(2*7/300)*70.42*(3.2/4.2) 2.504
Credit 6 =(6*7/300)*55.38 -7.753
Debtor 6 =(6*7/300)*86.496 12.109
11.910
Table 8.5.2 Working Capital Build-up

Working Capital Capacity Utilities (%) Total Working Increment (RM
Build Up Capital (RM Million) Million)
1st year operation 80.000 9.528 9.528
2nd year Operation 90.000 10.719 1.191
3rd year operation 100.00 11.910 1.191
380
Full Report
8.6 Project Profitability Assessment

It is a fundamental requirement of any company to operate profitably with the company
generating a positive income after a certain operating period. This section evaluates the
economics of the plant and provides information to the investors on the profitability of this
project, where this section includes a cash flow estimation, return on investment (ROI)
estimation, as well as financial assessment based on net present value (NPV), payback time ,
and discounted cash flow return (DCFR).
8.6.1 Cash Flow Estimation

Cash flow is the analysis of incoming and outgoing of money in the whole plant. The
construction of this plant is expected to be completed in two year which will commence in
2015. The design life of the equipment, piping etc. of this plant is 25 years while the
economic life is 20 years of operation. The first year of operation of the plant will run at 80%
capacity and second year with 90% capacity. From third operation year onwards, the plant
will run at 100% capacity, producing 30 kt per year of anhydrous ammonia per year and 96kt
of CO2 annually.
During the construction period, the cost involved are the fixed capital, initial catalyst
charge and start-up cost which include the cost of activated Methyldiethanolamine (aMDEA),
cooling water and refrigerant. The catalyst charge (start-up capital) was taken as 4 times the
cost of a batch of catalyst. The tax depreciation rate of 10% will be taken into account
throughout the 20 years of economic life. The corporate tax rate is 25% of the taxable income
for economic life which generates income. The inflation rate will not be taken into account in
the cash flow analysis of this plant. The cash flow table is attached as Table 8.6.1.1 in this
report.
8.6.2 Net Present Value (NPV) and Payback time

The net present value (NPV) of a project is the net value of the present value of all cash flows
for the project from the commencement of capital expenditure (2015) to the completion of
economic life (2037) (Brennan, 1998b). For the NPV estimation for Alternis BioAmmonia,
the discount rate is taken as 10% per annum. The total NPV obtained for this project is a
positive value of $10.26 million, implying that the project is profitable and a net cash benefit
is obtained as a result of this project. In addition, the payback time which is the time taken for
the project to recover investment costs is estimated as shown in Figure 8.6.2.1 below. It can
381
Full Report
be concluded from Figure 8.6.2.1 that the plant is able to recover and attain the capital
invested within 9.5 years or after the 7.5th year of operation of the plant.
Figure 8.6.2.1: Discounted cumulative cash flow diagram illustrating the payback period of the project
Based on Figure 8.6.2.1, it can be concluded that although the plant is profitable and is able
to achieve a positive NPV, the value of NPV calculated is still relatively small compared to
the amount of capital invested in the plant. This signifies that plant optimization measures
could be carried out to improve the profitability of the project and achieve a higher NPV
value.
382
Full Report
Table 8.6.1.1: Cumulative Cash Flow and Present Value
Planning Con. 1 Con. 2 Op. 1 Op.2 Op.3 Op.4 Op.5 Op.6 Op.7 op.8 op.9 op.10
2014 2015 2016 2017 2018 2019 2020 2021 2022 2023 2024 2025 2026
0 1 2 3 4 5 6 7 8 9 10 11 12
Fixed Capital (RM Million) -37.51 -37.51
Working Capital (RM Million) -9.53 -1.19 -1.19
Start-Up Capital (RM Million) -1.18
Sales Volume (t NH3/yr) 24000 27000 30000 30000 30000 30000 30000 30000 30000 30000
Selling Price (RM/ t NH3) 2883.20
Sales Revenue (RM Million) 69.20 77.85 86.50 86.50 86.50 86.50 86.50 86.50 86.50 86.50
Variable Cost (RM/t NH3) -1845.92
Variable Costs (RM Million) -44.30 -49.84 -55.38 -55.38 -55.38 -55.38 -55.38 -55.38 -55.38 -55.38
Fixed Cost (RM Million) 15.04 -15.04 -15.04 -15.04 -15.04 -15.04 -15.04 -15.04 -15.04 -15.04 -15.04
Cash Flow Before Tax 9.85 12.96 16.08 16.08 16.08 16.08 16.08 16.08 16.08 16.08
Tax Depreciation rate (%) 10
Tax Depreciation Allowance (RM Million) -3.75 -3.75 -3.75 -3.75 -3.75 -3.75 -3.75 -3.75 -3.75 -3.75
Taxable Income (RM Million) 6.10 9.21 12.33 12.33 12.33 12.33 12.33 12.33 12.33 12.33
Tax Rate (%) 25
Tax Payment (RM Million) -1.53 -2.30 -3.08 -3.08 -3.08 -3.08 -3.08 -3.08 -3.08 -3.08
Cash Flow After Tax (RM Million) 0.00 -37.51 -47.03 5.95 9.47 13.00 13.00 13.00 13.00 13.00 13.00 13.00 13.00
Cumulative Cash Flow After Tax (RM Million) 0.00 -37.51 -84.54 -78.59 -69.12 -56.12 -43.13 -30.13 -17.14 -4.14 8.85 21.85 34.85
Present Value Factor 0.91 0.83 0.75 0.68 0.62 0.56 0.51 0.47 0.42 0.39 0.35 0.32
Present Value (RM Million) -34.10 -38.87 4.47 6.47 8.07 7.34 6.67 6.06 5.51 5.01 4.55 4.14
op.11 op.12 op.13 op.14 op.15 op.16 op.17 op.18 op.19 op.20 Term
2027 2028 2029 2030 2031 2032 2033 2034 2035 2036 2037
13 14 15 16 17 18 19 20 21 22 23
Fixed Capital (RM Million)
Working Capital (RM Million) 11.91
Start-Up Capital (RM Million)
Sales Volume (t NH3/yr) 30000 30000 30000 30000 30000 30000 30000 30000 30000 30000
Selling Price (RM/ t NH3) 2883.20
Sales Revenue (RM Million) 86.50 86.50 86.50 86.50 86.50 86.50 86.50 86.50 86.50 86.50
Variable Cost (RM/t NH3) -1845.92
Variable Costs (RM Million) -55.38 -55.38 -55.38 -55.38 -55.38 -55.38 -55.38 -55.38 -55.38 -55.38
Fixed Cost (RM Million) 15.04 -15.04 -15.04 -15.04 -15.04 -15.04 -15.04 -15.04 -15.04 -15.04 -15.04
Cash Flow Before Tax 16.08 16.08 16.08 16.08 16.08 16.08 16.08 16.08 16.08 16.08
Tax Depreciation rate (%) 10
Tax Depreciation Allowance (RM Million) 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00
Taxable Income (RM Million) 16.08 16.08 16.08 16.08 16.08 16.08 16.08 16.08 16.08 16.08
Tax Rate (%) 25
Tax Payment (RM Million) -4.02 -4.02 -4.02 -4.02 -4.02 -4.02 -4.02 -4.02 -4.02 -4.02
Cash Flow After Tax (RM Million) 12.06 12.06 12.06 12.06 12.06 12.06 12.06 12.06 12.06 12.06 11.91
Cumulative Cash Flow After Tax (RM Million) 46.90 58.96 71.02 83.08 95.13 107.19 119.25 131.31 143.36 155.42 167.33
Present Value Factor 0.29 0.26 0.24 0.22 0.20 0.18 0.16 0.15 0.14 0.12 0.11
Present Value (RM Million) 3.49 3.18 2.89 2.62 2.39 2.17 1.97 1.79 1.63 1.48 1.33
383
Full Report
8.6.3 Internal Rate of Return

The internal rate of return (IRR), also known as discounted cash flow return (DCFR), is the
discount rate for which a NPV of zero is obtained. As the NPV varies at different discount
rates, the effect of discount rates on the NPV is analyzed and the figure is shown below,
where the IRR value is calculated to be about 12%.
Figure 8.6.3.1: Internal Rate of return from Graph of Discount Rate per Annum versus NPV
8.7 Sensitivity Analysis

Sensitivity analysis is integral in the risk evaluation of a project as it explores the variation of
several components of cash flow in the economic viability of the project. In this section, the
effect of changes in product (ammonia) selling price, feedstock (oil palm trunk) purchase
price and fixed capital cost on the payback period and NPV are investigated. The sensitivity
analysis is carried out by varying each component of the cash flow in increments of and
while keeping the remaining inputs unchanged. Through the sensitivity analysis
carried out, it can be seen that the profitability of the plant is most sensitive to product selling
price, indicating that a potential increase in product selling price in the future would lead to
increased profitability.
8.7.1 Product Selling Price

As mentioned above, changes of of and in ammonia selling price on the plant
profitability is investigated. From Figure 8.7.1.1(a) and Figure 8.7.1.1(b) below, it can be
seen that there is a significant increase in the NPV and decrease in the payback period of the
plant upon a slight increase in ammonia selling price. However, it should be noted that the
selling price of ammonia has to be kept within the market price due to the presence of other
competitors. In contrast, a decrease in ammonia selling price will have the opposite effect
leading to a longer payback period and negative NPV which is highly unfavorable.
384
Full Report
Figure 8.7.1.1(a) Effect of changes in selling price on Cumulative cash flow diagram and payback period (b) Effect of
changes in selling price on NPV
8.7.2 OPT Feedstock Purchase Price

The effect of and differences in the purchase price of OPT on the payback
period and NPV is illustrated in Figure 8.7.2.1 (a) and (b) below. It can be observed that
changes in OPT purchase price do not highly affect the payback period and NPV compared to
the previous analysis. An increase in the purchase price of OPT will increase the payback
period and at the same time reduce the NPV, even though NPV still remains positive with a
increase in OPT price. Also, a lower OPT purchase price is more profitable to the plant
as this cost will be taken into account throughout 20 years of operation. As, Malaysia is the
world’s second largest producer of crude palm oil (CPO), the purchase price of OPT will be
385
Full Report
affected by changes in CPO selling price, indicating that there is potential changes in the OPT
purchase price in the future.
Figure 8.7.2.1(a) Effect of changes in purchase price on Cumulative cash flow diagram and payback period (b) Effect
of changes in purchase price on NPV
386
Full Report
8.7.3 Fixed Capital Cost

Lastly, effect of and discrepancies in total fixed capital cost on the payback
period and NPV of the plant is also investigated. From Figure 8.7.3.1 (a) and (b) , it was
determined that a higher total fixed capital cost gives a longer payback period and at the same
time lower the NPV to an undesirable negative value. The total fixed capital cost is
minimized by proper equipment sizing, selection of material of construction and selecting the
best available technology. If the total fixed capital costs can be further reduced, the payback
period will be shortened and NPV will increase simultaneously.
Figure 8.7.3.1(a) Effect of changes in Fixed Capital Cost on Cumulative cash flow diagram and payback period (b)
Effect of changes in Fixed Capital Cost on NPV
387
Full Report
8.8 Critical overview on Economic Evaluation

The economic evaluation of the plant is crucial to allow the investors to assess the
profitability of investing in the design proposed by Alternis BioAmmonia to develop an
anhydrous ammonia plant. Analysis shown in the sections above revealed that the plant can
be regarded as a profitable investment even though a large initial amount of capital
investment of RM 75.01 Million is necessary as the investment could be recovered within the
intended operating life of 20 years. The total amount of fixed capital cost has been minimized
during the design phase by optimization of the process, utilization of reliable technologies
reliable and also minimizing the amount of equipment required where practical. Several
integral decisions were made during the design phase after considering the trade-offs between
capital and operating costs which includes the decision to purchase nitrogen as feedstock
instead of building an additional air separation unit in the plant due to the relatively low
amount of nitrogen required.
Based on the OPT feedstock cost of RM 250 per ton OPT and the current market price f the
products, RM 2400/ton anhydrous Ammonia and by-product Carbon Dioxide, RM 151/ton,
the project has been found to obtain a positive NPV of RM 10.26 Million at the end of the
20 years of operation life/economic life. The payback period of the plant discovered to be on
the 9th year of the design life including the 2 years of construction. In other words, investment
done is to be predicted to be received by the 7.5th operating year of the plant. However, the
IRR value calculated is 12% which is slightly higher than the discount rate taken (10%),
indicating that the project is profitable. These profitability values were obtained based on the
assumption that the demand for anhydrous fertilizer-grade ammonia is sufficiently high to
ensure that the entire product produced will be completely sold every year.
Furthermore, sensitivity analysis is done on several cash flow components as part of the risk
assessment on the economic viability of the project. The sensitivity analysis shows that slight
changes in product selling price will lead to significant changes in the NPV and payback
period of the plant. In contrast, the effect of changes in raw material prices and deviations in
total fixed capital cost are found to be comparably less significant. As there is a forecasted
increase in demand for ammonia powered by the increasing global population, it is safe to
assume that ammonia prices might increase in the future, potentially increasing the
profitability of the plant.
388
Full Report
CHAPTER 9 | PROJECT VIABILITY
9.1 Introduction
Alternis BioAmmonia has proposed to construct an Anhydrous Fertilizer-grade Ammonia
plant at a capacity of 30 kilotonnes per annum. In order to determine the viability of the
project, the factors that examined were economic, technical and environmental sustainability
assessment. Based on the economic, technical and environment criteria, a verdict on the
viability of the project can be made. In this section of report, technical, economic and
environmental base for viability in addition to the long term sustainability of the project and
future recommendation for potential improvements of the project will be presented.
9.2 Technical Viability

This particular design aims to maintain a highly reliable energy efficient system and at the
same time maintaining a low capital cost by reducing the inventories of hazardous materials
from the process. Anhydrous ammonia production is a well-established and utilized process
in the global industries. However upon approval of the Alternis BioAmmonia, it would be the
first commercial producer of ammonia from a biomass feedstock such as Oil Palm Trunk
(OPT). This plant incorporates the tested, proven and reliable technologies available in the
patents, literature and worldwide technologies to provide and assure a reliable and safe design.
As previously discussed and elaborated in Submission A (Chapter 2), Technology Evaluation,

the processing technologies were chosen based on the efficiency, reliability, safety aspect,
cost (including capital and operating cost) and environmental sustainability. Besides that, the
gasifier chosen was also found to be compatible with the OPT after evaluating the elemental
composition of OPT. In this instance, the sulphur content from biomass is comparatively
lower. Hence, desulphurization step could be excluded from the process flow of this plant.
Moreover, the use of autothermal reformer also compliments to this process plant as methane
content from gasification was reduced to minimum. Apart from that, Alternis BioAmmonia
has incorporated established technologies in other section of the process plant which have
been widely used in the conventional ammonia plant production plant.
Additionally, the project team has also performed optimization of the critical process
parameters to ensure the plant has increased efficiencies in the overall consumption of raw
materials, utilities and energy. In order to consider this industry is technically feasible, the
technologies used in Alternis BioAmmonia are able to meet the specific process requirement
389
Full Report
and the product specification requirement without compromising the safety and
environmental aspects of the plant. The list below states some of the conditions the operating
plant has to fulfill:
 Production of 30 of fertilizer grade anhydrous ammonia;
 The temperature of the product produced should be between the ranges of -30 to -
25 .
 The pressure of the system has to be fixed at 200 .
 The purity of the anhydrous ammonia produced ranges between 0.995 to 0.998, and
the water content should be between 0.002 and 0.005.
 Density of the product liquid is specified to be 620 with a clear appearance.
Product specification can be attained with the current plant design, however the assumptions
and simplifications made during mass balance and energy balance might result in a lesser
efficient process compared to prediction and yet in depth analyses of varying design cases are
recommended to optimize the plant for varying feedstock composition and conditions. In
addition, Alternis BioAmmonia also aims to minimize the usage of utilities such as electricity
and cooling water, hence, heat and water integration were conducted to maximize the energy
recovery of the plant. At the same time, the characteristics of feedstock have to be
comprehend in order to optimize the plant to increase the efficiency and improving the
environmental sustainability of the process.
The designs of major and minor equipment have also taken safety design margins into
accounts in accordance to the Australian Standard. Furthermore, the control and
instrumentation system of this plant incorporates appropriate control and alarm systems as
well as monitoring sensors to monitor operating units and streams across the whole plant.
Also, extra safety measures such as connecting the sensors to an independent safety circuit to
ensure appropriate response during possible hazardous events were also included. Moreover,
safety measures present in the plant that are incorporated include safety interlock system and
emergency shutdown system which will be used in case of uncontrolled and runaway
reactions.
Safety and risk assessments were carried out over the entire plant for the purpose of
identifying potential hazards that might occur and affect the community. It was deduced that
during the normal operation, plant have a more acceptable safety measure in place, which
390
Full Report
poses no immense risk to the staff, environment and the local community. It is important to
highlight that the start-up and shut down did not consist of a comprehensive design and that it
can be taken into account more thoroughly before the approval for the construction of the
plant.. Safety considerations were discussed in Submission C (Chapter 7.1.8).
In addition, alternis BioAmmonia has design a plant layout in accordance to Guidelines for
Facility Siting and Layout published by the American Institute of Chemical Engineers
(AIChE) which provides recommended spacing distances betweem hazardous equipment to
avoid chain reaction in the occurrence of hazardous events as shown in Submission C
(Chapter 7.1.9). The load bearing of the concrete foundation was not considered during this
design project but it’s one of the aspects that can be looked into for further improvement.
The criteria for selection of location were minimizing the risk to the environment and local
community. The plant layout in Submission C (Chapter 7.1.9) takes into account these
aspects as well as consideration of space for future expansion. The layout of this plant was
designed in such a way it would allow linear start-up process and thus decreasing the
interconnecting pipe work leading to rapid start-up of the system. The location and layout of
the plants considered to be suitable. The site location chosen for anhydrous ammonia
production offers a large flattened land which is surrounded by palm oil tree. The particular
location has access to river water and roads, thus providing an advantage during the
construction stage. Safety was sensibly and wisely considered during the plant layout stage;
subsequently the flare was located downwind of the processing unit as well as the office and
administration buildings.
The current plant design is capable of meeting the required targets and product specification
whilst also providing a safe working environment for employees. Considering the current
location, technical aspects and layout design of Alternis BioAmmonia the plant is considered
to be viable after considering the criteria mentioned.
9.3 Economic Viability

Generally, the economic viability of a project refers to the assessment of the capacity of the
project to meet the defined objective in addition to generating significant economic gains to
the stakeholders and to the economy. Several keys were carried out to determine the
economic viability of the anhydrous ammonia plant. The economic viability of the anhydrous
391
Full Report
ammonia plant is assessed based on cash flow analysis, payback period, net present value
(NPV) and internal return rate (IRR), where a detailed economic evaluation can be found in
Submission C (Chapter 8). Apart from that, market evaluation on anhydrous fertilizer-grade
ammonia was performed where information on the demand and product selling price were
also included. A detailed study on the end –uses of ammonia revealed that up to 82% of the
world ammonia produces are used as nitrogen fertilizers. Not only that, the world
consumption of ammonia has been experiencing an annual steady growth of about 2.3%. This
indicates that the demand for ammonia is high with prospective future market expansion.
Thus, the ammonia produced will be sold at a competitive market price of RM 2400/ton
ammonia which enables the plant to attain reasonable market share and is able to measure up
to other competitors in the market. In addition, the project also aims to generate additional
revenue through sale of its by-product, carbon dioxide, which will undergo dehydration
process prior to being sold to the market.
In addition, the fixed capital investment comprising of equipment costs, construction costs to
physically erect the plant as well as other miscellaneous cost such as contingency and
engineering cost were estimated to be about RM 75 million which is relatively lower than the
capital cost of $89 million required for typical ammonia production plant of similar capacity
(Maung , et al., 2012). However, for operating costs, it comprises of raw materials, utility
requirements, plant overheads and other fixed charges.
Referring to Submission C (Chapter 8), the cash flow table shows that the plant is able to
generate a total of RM 167.33 millions of profit from the entire lifetime of plant. The shape
of the cash flow diagram as shown in Submission C (Chapter 8) indicates a marginal
profitability. The cumulative cash flow is positive for the greater part of the project life. Sales
revenue of RM 86.5 millions per year can be generated.
The calculated payback time for the project is determined to be approximately 9 years. A
positive value of NPV indicating a net cash benefit can be achieved in this project, it was
found to be RM 10.26 millions at a discount rate of 10%. Including the period of 1 year of
construction, recovery of initial investment capital can be achieved after 9 years of operation.
The payback period is considered acceptable as compared to the long operation lifetime of 25
years.
In order to determine the profitability of the plant, several approaches were used such as
Return on Investment (ROI), financial assessment based on Net Present Value (NPV),
392
Full Report
payback time, and Internal Rate of Return (IRR). The design and planning of the plant will be
carried out within the first year and construction and commissioning will commence in the
subsequent two years before the plant beings operation in the fourth year. The plant is
proposed to reach maximum capacity utilization in phases, with the first year of operation
having a capacity of 80% which increases to 90% in the subsequent year and finally running
at 100% capacity until its last year of operation.
The IRR calculated is around 12%. This is the maximum discount rate that can be obtained to
make the project remains economically viable, higher discount rate will results in a negative
NPV. The IRR achieved in this project using oil palm trunk as feedstock is comparable to
typical ammonia plant using coal as feedstock which has 15% IRR (Pivot, 2013).
Furthermore, 12% of IRR is considered high and will be able to attract more investment into
the project.
From the sensitivity analysis done in Submission C (Chapter 8), the selling price, corporate
tax and electricity are more sensitive to changes. The most sensitive factor would be the
selling price of ammonia. Sensitivity analysis for the selling price was done using the current
average market price. The price is more likely to increase due to the increasing demand of
fertilizer. Hence, this causes this project to be more economically viable.
In order to further increase the profitability of this project, several improvements can be made.
First of all, the operating cost can be reduced by decreasing the labour cost. Second,
technology improvement can be done in the future to make the production process more cost
effective. Lastly, extra revenue is possible to be obtained by selling the electricity generated
using the excess steam produced in the plant.
As a conclusion, the ammonia production plant project proposed by Alternis BioAmmonia is

deemed to be economically viable due to the positive NPV, high ROI as well as IRR values
estimated. However, the profitability of the plant will be affected by fluctuations in operating
expenditures and future market conditions, Alternis BioAmmonia has to ensure that the funds
obtained are used to enhance the economic development of the plant to assure that sufficient
profits can be achieved annually. Furthermore, the proposed plant places strong emphasis on
reducing environmental emissions by utilizing waste biomass as feedstock and reducing
greenhouse gas emissions, thus indicating that there is potential for improvements in the
economic viability of the project.
393
Full Report
9.4 Environmental Viability and Sustainability

Walking towards a more environmental friendly century, issue regarding the depletion of
natural resources has become a crucial topic to be discussed. Since, environmental friendly
and sustainability is an inevitably issue. Therefore pragmatic approaches need to be proposed
to reduce the usage and reliance of natural resources. In this case, oil palm trunk biomass
becoming attractive under special environmental concern where it can be used as a
replacement for natural gas which is the main feedstock in the conventional anhydrous
ammonia production.
The anhydrous ammonia plant proposed by Alternis BioAmmonia will be using oil palm
trunk as feedstock. Environmental Aspect and Impact Register was performed to determine
the viability of the project in the environmental perspective. The subjects that are included in
the evaluation are the impact on air quality, water quality, natural resources noise as well as
the land. Enactments, regulations and mitigation steps are suggested and enforced in the plant
to enhance viability of the plant from environmental perspective.
Using oil palm trunk biomass in the production of anhydrous ammonia is very environmental
friendly as the oil palm trunk will be fully utilized in the process plant especially in the
gasification process to produce syngas. Subsequently, go through a series of process to
generate sufficient amount of nitrogen and hydrogen for the production of anhydrous
ammonia. In contrast, if oil palm trunk is used in plywood industry, only selected layer of the
oil palm trunk will be used. Therefore, instead of treating the oil palm trunk as waste, it is
more environmental friendly to use it as feedstock of anhydrous ammonia production.
Furthermore, replacing natural gas as a feedstock of gasifier in ammonia plant is a greener

approach. As mentioned, depletion of natural resources is an inevitable issue. By replacing
natural gas with biomass can be a first step to step out from the conventional production of
anhydrous ammonia which consumes a lot of fossil fuel.
Apart from that, carbon dioxide will produced in the glycol plant where it will be capture and
sell to gain extra sales revenue. Carbon dioxide has a lot of usage and it is widely used in
various industries. Therefore, carbon dioxide generated will not be a waste and giving any
impact to the environment but a useful output from the plant.
394
Full Report
Based on LCA studies, the major environmental impact categories that are largely
contributed by the production of Ammonia based on both feedstock are Fossil Depletion,
Global Warming Potential, Terrestrial Acidification, and slightly contributes to the
Photochemical Oxidant Formation, Particulate Matter Formation Potential and Marine
Eutrophication. It can be clearly seen that, production of Ammonia using the Biomass
Feedstock largely reduced the impact on the Fossil Depletion as the Conventional production
requires mining of the natural gas. Usage of the palm biomass waste does not require fossil
fuels except for the purpose electricity and transport fuel requirement. This goes along with
the current global issue of mitigating the natural resource depletion as substitution of biomass
as feedstock for ammonia production will help to reserve the natural resources better. Besides
that, the Global Warming Potential due to the carbon dioxide emission is also largely reduced
by the usage of Palm Biomass as the feedstock as the emission of CO2 from the plant is being
compensated by the absorption of CO2 by the palm tree at the plantation stage. Besides, the
ammonia plant designed in a way that the CO2 removed at the Carbon Dioxide Removal
Stage is being sold to the nearby Glycol Plant to be used for other application. In addition,
scope 1 emissions within the LCA boundary include air emissions produced during the
transportation of feedstock in and out of the plant. In order to minimize these emissions,
regular diesel fuel was substituted with the biodiesel fuel blend, B20, which has proven to
reduce the content of VOCs and CO produced through fuel combustion. Not only was that,
the scope 2 emissions of the plant also generated through the use of utilities such as heat and
electricity. These were curtailed by integrating the heat generated and consumed by various
processes within the plant for all heating and cooling purposes within the plant. In spite of
that, the Marine Eutrophication Potential is found to be greater for the production ammonia
using biomass feedstock mainly due to the usage of the pesticides during the plantation,
harvesting of Palm Biomass. This can be overcome, by reducing the usage of chemical
pesticides and substituting it with organic chemicals.
Nevertheless, The wastewater leaving the CO2 removal system and the gas purification
system will be sent off to an off-site wastewater treatment plant where then contaminants in
the wastewater will be decreased to the acceptable limit set by Department of Environment
(DOE) to ensure the preservation of marine life and plant species when released back into the
river. In addition, the plant was designed in compliance with the ISO4001 standard which
encourages the use of inherent identified from the Impact Aspect Register will be managed,
monitored and controlled through an Environmental Management System (EMS) which will
395
Full Report
reduce the potential environmental impacts considerably during the construction, operation
and decommissioning phase of the plant. The suggested mitigation methods will have
positive effect on the waste management, drainage systems and impacts related to water air
within the plant.
In conclusion, the production of Ammonia using OPT Palm Biomass is highly sustainable
and environmental friendly compared to that of the Natural Gas. The project is proven to be
environmentally viable as it strives to minimize the use of the non-renewable resources while
promoting cleaner production.
9.5 Strategic aspects affecting the future viability and sustainability of the
project
9.5.1 Future growth and demand of fertilizer grade ammonia
The future growth and demand of ammonia is expected to increase further in the next future.
According to Potash Corp, 2013, the world’s demand for ammonia is predicted to escalate at
an approximate rate of 3% annually for the next five years. From these prediction, it is
estimated that around 85% of the consumption of ammonia is mainly used for fertilizer,
(Appl, M. 2011), which mostly comes from the agricultural sector. The ammonia production
is suggested to develop in proportion with the world’s population growth (Appl m, 2011), due
to the production of agricultural fertilizer that has relatively increased the world’s agricultural
productivity in most area of the world. Therefore, as the yield of the agricultural product
increased, the number of world’s population supported per land utilized by the fertilizer
would also improve. Generally, the main driving source for developing the fertilizer
production, which is the demand of ammonia, is mostly due to the economic growth as well
as the nutrient improvement in the developing countries. In order to yield higher production
of ammonia, more feedstock would therefore be required. Natural gas has commonly been
use as the feedstock due to its plentiful supply and low cost. However, considering the high
carbon emission and non-renewability of the gas, better alternatives feedstock that is more
sustainable and environmentally friendly like biomass are highly demanded. On a local front,
Malaysia is expected to expand the use of palm oils in biodiesel production, indicating that
the local demand for fertilizers is projected to grow as well. In particular, the use of biomass
for the production of hydrogen and biofuels is believed to improve the development of a new
396
Full Report
market. According to Agensi Inovasi Malaysia, 2011, the utilization of 20 million tonnes of
palm oil biomass by 2020 has the ability to take part in the economy of the country.
9.5.2 Future trends in technology

The concept of substituting fossil-derived energy has promoted the use of biomass as a fuel
source due to energy security concerns, potential for economic development and reduction of
environmental consequences of fossil energy use. Alternis BioAmmonia plant is planning to
invest in the research for the development of alternative feedstock for the production of
ammonia using palm oil residue, the palm oil trunks. This research and development is
deemed to be the future trend in the production technology of ammonia. Several suggestions
for the improvement in technologies are presented in the subsequent future recommendation
section.
9.6 Future Recommendations

Based from the evaluations earlier, the design project proposed by Alternis BioAmmonia
plant to design an anhydrous ammonia production plant using palm oil trunk feedstock is
predicted to be viable in terms of technical, economic and environmental viewpoints. Even so,
several recommendations could be taken for the design to improve the overall project
viability and profitability of the plant.
Firstly, it can be achieved through the investigation of alternate technologies and

optimization of current design. The greatest barrier lies in the biomass gasification
technology system which is still currently facing challenges in its commercial development.
Gasifier is the core component of the Alternis BioAmmonia plant, since it sets the primary
requirements of the biomass raw material inputs and determines the product gas composition.
One of the challenges is that it requires high investment due to crucial monitoring devices for
process control and demanding conditions in regards to the materials used under severe
physical and chemical stress (Dutta & Acharya, 2011). Furthermore, gasifier’s poor carbon
conversion of biomass is a major problem knowing oil palm trunk feedstock is only available
during the felling process and is thus a limited feedstock. A future recommendation can be
through using other parts of oil palm waste or other form of biomass for the gasification
process. Additionally, a comprehensive and well-funded R&D program is recommended to
create models to better understand the kinetics and particulate behaviour of fuel inside
397
Full Report
gasifier to increase performance of the gasifier and decrease the cost of operation and
maintenance, therefore substantially optimizing gasifier operation. Moreover, increasing
carbon recovery also means less carbon levels in gasifier ash which in turn increases ash
quality. This can be used to create revenue from the sales for their end-uses such as for
fertilizer on agricultural forest soils or as a raw material in cement and brick industry.
In terms of process optimization, recommendations include generation of electricity

requirements onsite. The plant should incorporate the simultaneous generation of heat and
power (cogeneration) to increase robustness of the energy infrastructure in the plant. This will
thus reduce CO2 emissions as well as cost for utility and thus improves viability in terms of
environmental and economic aspect respectively. Furthermore, rather than discharging
wastewater to off-site treatment facilities, considerable savings in water consumptions can be
achieved by having an on-site treatment plant. Added to that, the demand of Ammonia is
known to be increasing globally as discussed previously, which is due to the increased
demand in agricultural product. As a result, Alternis BioAmmonia plant has accommodated
for future expansion of the plant to coincide with the product demand. Apart from that, rises
in fossil fuel prices, their scarcity and penalties for environmental contamination would
become a convincing reason for the application of government incentives for the plant’s
capital funding. This project is highly appealing since this would also solve disposal
problems of abundant oil palm wastes by converting them into an economically useful
alternative. If a government funding can be approved, this will increase profitability with
earlier payback period and higher NPV.
In conclusion, Alternis BioAmmonia is a potential viable plant proven through various

viability studies evaluated earlier. The recommendation described above listed the possible
improvements that can be explored and implemented in the future to increase the efficiency,
sustainability and profitability of the plant. By incorporating these recommendations to the
design plant, the project is thus recommended for corporate approval.
398
Full Report
Reference
Chapter 1
ANON. 2000. Production Of Ammonia. Best Available Techniques for Pollution Prevention and
Control in the European Fertilizer Industry
[Online], 1. Available: http://www.diquima.upm.es/docencia/tqi/docs/amonia00.pdf
BANK, W. 2012. State and trends of the carbon market 2012. Washington DC: World Bank.
BARTELS, J. R. 2008. A feasibility study of implementing an Ammonia Economy Master of Science,
Iowa State University.
BARTELS, J. R. & PATE, M. B. 2008. A feasibility study of implementing an Ammonia Economy.
United states.
C.W.RITZ, B.D.FARICHILD & M.P. 2004. Implications of Ammonia Production and Emissions
from Commercial Poultry Facilities: A Review. 2. Available:
http://japr.fass.org/content/13/4/684.full.pdf
CFINDUSTRIES 2013. Anhydrous Ammonia Material Safety Data Sheet. 1-15.
CHANDRAWATHANI, P., PREMALATHA, B., ERWANAS, A., ZAINI, C. & RAMLAN, M. 2013.
<Chandra- Malaysia Chandra (2) (1).pdf>.
DERIS, R. R. R., SULAIMAN, M. R., DARUS, F. M., MAHMUD, M. S. & BAKAR, N. A. 2006.
PYROLYSIS OF OIL PALM TRUNK (OPT). 245-250.
DIVISION, E. A. I. D. 2012. Oil Palm Planted Area, Dec 2012. In: DIVISION, E. A. I. D. (ed.).
Malaysia: Malaysian Palm Oil Board.
E&Y 2012. The future of global cabon markets. Ernst & Young.
GILBERT, P., THORNLEY, P., ALEXANDER, S. & BRAMMER, J. 2009. Biomass gasification for
ammonia production. International conference on polygeneration strategies. Vienna:
University of Manchester.
GOH, C. S., TAN, K. T., LEE, K. T. & BHATIA, S. 2010. Bio-ethanol from lignocellulose: Status,
perspectives and challenges in Malaysia. Bioresour Technol, 101, 4834-41.
I., W. A., S., M., H., Z., S., O. & MORI, Y. 2005. Malaysian Oil Palm Biomass. In: (FRIM), F. R. I.
M. (ed.). Osaka, Japan: Global Environment Centre Foundation.
JUNG, Y. H., KIM, I. J., KIM, J. J., OH, K. K., HAN, J. I., CHOI, I. G. & KIM, K. H. 2011. Ethanol
production from oil palm trunks treated with aqueous ammonia and cellulase. Bioresour
Technol, 102, 7307-12.
KELLY-YONG, T. L., LEE, K. T., MOHAMED, A. R. & BHATIA, S. 2007. Potential of hydrogen
from oil palm biomass as a source of renewable energy worldwide. Energy Policy, 35, 5692-
5701.
KOSUGI, A., TANAKA, R., MAGARA, K., MURATA, Y., ARAI, T., SULAIMAN, O., HASHIM,
R., HAMID, Z. A., YAHYA, M. K., YUSOF, M. N., IBRAHIM, W. A. & MORI, Y. 2010.
Ethanol and lactic acid production using sap squeezed from old oil palm trunks felled for
replanting. J Biosci Bioeng, 110, 322-5.
MCKENDRY, P. 2002. Energy production from biomass (part 3): gasification technologies.
Bioresource Technology, 83, 55-63.
MURATA, Y., TANAKA, R., FUJIMOTO, K., KOSUGI, A., ARAI, T., TOGAWA, E., TAKANO,
T., IBRAHIM, W. A., ELHAM, P., SULAIMAN, O., HASHIM, R. & MORI, Y. 2013.
Development of sap compressing systems from oil palm trunk. Biomass and Bioenergy, 51, 8-
16.
NIPATTUMMAKUL, N., AHMED, I. I., KERDSUWAN, S. & GUPTA, A. K. 2012. Steam
gasification of oil palm trunk waste for clean syngas production. Applied Energy, 92, 778-782.
POTASHCORP 2011. World Ammonia Production Profile.
RAFIQUL, I., WEBER, C., LEHMANN, B. & VOSS, A. 2005. Energy efficiency improvements in
ammonia production—perspectives and uncertainties. Energy, 30, 2487-2504.
RICHMOND, S. & KARLIN, A. 2010. Lonely Planet Malaysia, Singapore & Brunei, Lonely Planet
399
Full Report
RUSHING, S. A. 2010. Carbon Dioxide – A friend in industry and an environmental foe [Online].
Advanced Cryogenics. [Accessed 4th September 2013].
S.RAY, M. & W.JOHNSTON, D. 1989. Chemical Engineering Design Project A Case Study
Approach US, Gordon and Breach Science.
SABRI, M. A. 2009. Evolution of Fertilizer Use By Crops in Malaysia: Recent Trends and Prospects.
IFA Crossroads Asia-Pacific 2009. Kota Kinabalu, Malaysia: Fertilizer Industry Association
of Malaysia.
SINGH, K. J. 2013. National Biomass Strategy 2020: New wealth creation for the palm oil industry In:
AGENCY, M. I. (ed.). Malaysia: Malaysia Innovation Agency.
SVENDSEN, G. T. 1998. Towards a CO2 market in the EU: the case of electric utilities. European
Environment, 8, 9.
VEAL, G. & MOUZAS, S. 2012. Market-based responses to climate change: CO2 market design
versus operation. SAGE, 33, 29.
Chapter 2
Aden, A. (2009). Survey and Down-Selection of Acid Gas Removal Systems for the Thermochemical
COnversion of Biomass to Ethanol with a Detailed Analysis of an MDEA system . California:
National Renewable Energy Laboratory .
AdvancedGasTechonologies. (2011). Advanced Gas Technologies. Retrieved 4 September, 2013,
from http://www.adgastech.com/Nitrogen-Costs.aspx
Almqvist, E. (2003). History of Industrial Gases (1st ed.). New York: Springer.
Alvis, R. S., Hatcher, N. A., & Weiland, R. H. (2012).
CO2 Removal from Syngas Using Piperazine‐Activated MDEA and Potassium Dimethyl Glycin
ate. Optimized Gas Treating , 1, 1-10.
Appl, M. (2005). Ammonia, 2. Production Processes. Ullmann's Encyclopedia of Industrial Chemistry
.
Armarego, W., & Chai, C. (2009). Purification of Laboratory Chemicals (6th ed.). Massachusetts:
Elsevier.
AGARWAL, P. 2010. Ammonia: The Next Step. Process and Reactions.
AHMED BENSAFI, B. T. 2007. Transcritical R744 (CO2) heat pumps. Technician's Manual.
Sustainable Heat and Energy Research for Heat Pumps Applications. SHERHPA.
ALIJANI, A. & IRANKHAH, A. 2013. Medium-Temperature Shift Catalysts for Hydrogen
Purification in a Single-Stage Reactor. Chemical Engineering & Technology, 36, 209-219.
AMERICAN SOCIETY OF HEATING, R. & ENGINEERS, A.-C. 2001. 2001 ASHRAE Handbook:
Fundamentals, ASHRAE.
AMOS, W. A. 2003. Biological Water-Gas Shift Conversion of Carbon Monoxide to Hydrogen. 1-22.
ANSBRO, J. 2003. Packaged Ammonia Chillers with Variable Frequency Drives. Johnson Controls.
AXON, S. A. & CASCI, J. L. 2008. Recycling of Spent Catalysts Containing Base Metals. Handbook
of Heterogeneous Catalysis, 1863-1871.
Babicki, M. (2003). Chemical Processing. Retrieved 29 August, 2013, from
http://www.chemicalprocessing.com/articles/2003/322/?start=2
Bhatia, A. (2013). Overview of Chillers Compressors. CED Engineering .
Bridgwatera, A., Meierb, D., & Radleinc, D. (1999). An overview of fast pyrolysis of biomass.
Organic Geochemistry , 30 (12), 1479-1493.
BANK, W. 2012. State and trends of the carbon market 2012. Washington DC: World Bank.
BANKS, C., ZHANG, Y., HEAVEN, S., WATSON, G. & POWRIE, W. 2010. Particle size
requirements for effective bioprocessing of biodegradable municipal waste. Technology
Research and Innovation Fund Project Report University of Southampton.
400
Full Report
BARBAROSSA, V., BASSANO, C., DEIANA, P. & VANGA, G. CO2 Conversion to CH4.
BENAMOR, A. 2012. The removal of CO2 from natural gas-conventional and emerging technologies.
Gas Processing Center, 36.
BFPIC. 2009. Oil Palm Biomass [Online]. Beijing: Beijing Forestry and Parks Department of
International Cooperation. [Accessed 2nd February 2013].
BRONSON, B., PRETO, F. & MEHRANI, P. 2012. Effect of pretreatment on the physical properties
of biomass and its relation to fluidized bed gasification. Environmental Progress & Sustainable
Energy, 31, 335-339.
BUSTAMANTE, F., ENICK, R. M., KILLMEYER, R. P., HOWARD, B. H., ROTHENBERGER, K.
S., CUGINI, A. V., MORREALE, B. D. & CIOCCO, M. V. 2005. Uncatalyzed and wall-
catalyzed forward water-gas shift reaction kinetics. AIChE Journal, 51, 1440-1454.
Callaghan, C. A. (2006). Kinetics and Catalysis of the Water-Gas-Shift Reaction: A Microkinetic and
Graph Theoretic Approach . Worcester: Worcester Polytechnic Institute.
Chaiyaomporn, K., & Chavalparit, O. (2010). Fuel Pellets Production from Biodiesel Waste.
EnvironmentAsia , 3 (1), 103-110.
Chewa, J. W., Haysb, R., Findlayb, J. G., Knowltonb, T. M., Karrib, S. R., Coccob, R. A., et al.
(2012). Cluster characteristics of Geldart Group B particles in a pilot-scale CFB riser. I.
Monodisperse systems. Chemical Engineering Science , 68 (1), 72-81.
CHEN, D. & HE, L. 2011. Towards an Efficient Hydrogen Production from Biomass: A Review of
Processes and Materials. ChemCatChem, 3, 490-511.
CHEN, W.-H., HSIEH, T.-C. & JIANG, T. L. 2008. An experimental study on carbon monoxide
conversion and hydrogen generation from water gas shift reaction. Energy Conversion and
Management, 49, 2801-2808.
CHIANG, K.-Y., CHIEN, K.-L. & LU, C.-H. 2012. Characterization and comparison of biomass
produced from various sources: Suggestions for selection of pretreatment technologies in
biomass-to-energy. Applied Energy, 100, 164-171.
DERIS, R. R. R., SULAIMAN, M. R., DARUS, F. M., MAHMUD, M. S. & BAKAR, N. A. 2006.
PYROLYSIS OF OIL PALM TRUNK (OPT). 245-250.
Derks, P. (2006). Carbon Dioxide Absorption in Piperazine Activated N-Methyldiethanolamine. The
Netherlands: University of Twente.
Dyer, C. K., Moseley, P. T., Ogumi, Z., Rand, D. A., & Scrosati, B. (2013). Encyclopedia of
Electrochemical Power Sources (1st ed.). (J. Garche, Ed.) Amsterdam: Elsevier.
Eigenberger, G. (1992). Fixed-Bed Reactors. Ullmann's Encyclopedia of Industrial Chemistry , 4,
200-237.
E4TECH 2009. Review of Technologies for Gasification of Biomass and Wastes. UK: NNFCC.
E&Y 2012. The future of global cabon markets. Ernst & Young.
EUROPEAN, C. 2007. Integrated Pollution Prevention and Control. 1-446.
Fang, Z. (2013). Pretreatment Techniques for Biofuels and Biorefineries. In Green Energy and
Technology (pp. 94-282). Heidelberg: Springer.
Farrauto, R., Hwang, S., Shore, L., Ruettinger, W., Lampert, J., Giroux, T., et al. (2003). Annu. Rev.
Mater. Res. 33, 1-27.
FORMENT, G. F. 1977. Fixed bed catalytic reactors. Industrial and engineering chemistry, 18-26.
FOSCOLO, P. U. 1997. Production of Hydrogen-rich Gas by Biomass Gasification Application to

Small Scale, Fuel Cell Electricity Generation in Rural Areas. France: University OF L’Aquila.
FRANCESCONI, J. A., MUSSATI, M. C. & AGUIRRE, P. A. 2007. Analysis of design variables for
water-gas-shift reactors by model-based optimization. Journal of Power Sources, 173, 467-477.
401
Full Report
GAO, J., YINGLI WANG, YUAN PING, HU, D., XU, G., GU, F. & SU, F. 2012. A thermodynamic
analysis of methanation reactions of carbon oxides for the production of synthetic natural gas.
RSC Advances, 2, 2358–2368.
GEORGE, C. 2001. Carbon Monoxide. Kirk-Othmer Encyclopedia of Chemical Technology, 5, 1-27.
GÖRANSSON, K., SÖDERLIND, U., HE, J. & ZHANG, W. 2011. Review of syngas production via
biomass DFBGs. Renewable and Sustainable Energy Reviews, 15, 482-492.
GPSA 2004. Engineering Data Book, Gas Processor Supplier Association
GRASA, G., WELLMAN, R. G., KILGALLON, P., SIMMS, N. J. & OAKEY, J. E. 2004. Novel Hot
Gas Cleaning/ Heat Recovery System. Bedfordshire: Cranfield University
Gutiérrez, L. F., Sánchez, Ó. J., & Cardona, C. A. (2009). Process integration possibilities for
biodiesel production from palm oil using ethanol obtained from lignocellulosic residues of oil
palm industry. Bioresource Technology , 100, 1227-1237.
Heyne, S. (2013). Bio-SNG from Thermal Gasification -. CHALMERS UNIVERSITY OF
TECHNOLOGY. Göteborg: Chalmers Reproservice.
Heyne, S., Seemann, M. C., & Harvey, S. (2010). Integration study for alternative methanation
technologies. Heat and Power Technology , 2-4.
Hindrichs, H. (1962). A New Type of Converter for Ammonia Synthesis. Ammonia Synthesis .
Hooi, K. K., Alauddin, Z. A., & Ong, L. K. (2009). Laboratory-Scale Pyrolysis of Oil Palm Pressed
Fruit Fibre. Journal of Oil Palm Research , 21, 577-587.
Hutchings, G., Copperthwaite, R., Gottschalk, F., Hunter, R., Mellor, J., Orchard, S., et al. (1992).
Catalyst. 137 (1), 408-422.
HEIDENREICH, S. 2013. Hot Gas Filtration - A review. Fuel, 104, 83-93.
HEYNE, S., SEEMANN, M. C. & HARVEY, S. Integration study for alternative methanation
technologies for the production of synthetic natural gas from gasified biomass. Heat and power
technology.
INAYAT, A., AHMAD, M. M., HAMID, M. I. A. & YUSUP, S. 2010. Flowsheet Modelling of
Biomass Steam Gasification System with Carbon Dioxide Capture for Hydrogen Production.
Proceedings of International Conference on Advances in Renewable Energy Technologies
ICARET2010, 1-6.
INDIA, P. P. 2013. Gravity Settling Chamber. Available:

http://www.productivity.in/knowledgebase/Environmental%20Management/b.%20Air%20Qualit
y%20Management/Air%20Pollution%20Control/Gravity%20Settling%20Chambers.pdf.
INSTITUTE, E. N. 2007. The Importance of Nickel Compounds. Belgium: The Weinberg Group
LLC.
I.PLASYNSKI, S. 2000. Review of CO2 capture technologies and some improvement opportunities.
Available:
http://web.anl.gov/PCS/acsfuel/preprint%20archive/Files/45_4_WASHINGTON%20DC_08-
00_0644.pdf [Accessed 28 August 2013].
IGNOU. n.d. Refrigeration and Air Conditioning [Online]. IGNOU. Available:

http://www.ignou.ac.in/upload/Unit%203-32.pdf [Accessed 31st August 2013].
IIAR 2005. Ammonia: The Natural Refrigerant. IIAR Green Paper.
J.M., M. (2009). Design of water-gas shift reactors. Chem. Eng. Prog. , 58 (3), 6-33.
J.R., L., & J.P., W. (2003). Handbook of Fuel Cells. In Fuel Cell Technology and Applications (Vol.
3, pp. 190-201). England: John Wiley & Sons.
Jakobsen, H. (2008). Chapter 11: Packed Bed Reactors. In Chemical Reactor Modeling (pp. 953-984).
Berlin : Springer.
402
Full Report
KBR. (2013). Purifier™ (Cryogenic Gas Purification System). Retrieved 19 August, 2013, from
KBR: http://www.kbr.com/Technologies/Proprietary-Equipment/Purifier-Cryogenic-Gas-
Purification-System/
Kubek, D. I., Polla, E., & Wilcher, F. P. (2009). Purification and Recovery Options for Gasification.
University of Phoenix, Illinois.
Kunjunny, A. M., Patel, M. R., & Navin, N. (1999). Revamping if CO2 removal section in ammonia
plant at IFFCO Kalol. Fertiliser News , 44 (8), 53-55.
KANG, W. R. & LEE, K. B. 2013. Effect of operating parameters on methanation reaction for the
production of synthetic natural gas. Korean J. Chem. Eng., 30.
KAUSHAL, P. & TYAGI, R. 2012. Steam assisted biomass gasification-an overview. The Canadian
Journal of Chemical Engineering, 90, 1043-1058.
KEPICS, S., HARDING, D. & HUBBARD, T. 2008. CRYOGENIC SEPARATION OF

HYDROGEN METHANE MIXTURE.
KHORSAND , K. & DEHGHAN, K. 2007. Modeling and Simulation of Reformer Auto-Thermal

Reactor in Ammonia Unit. Petroleum & Coal, 49, 64-71.
KIENBERGER, T., MUEHLBERGER, T., LETNNER, J. & KARL, J. 2009. Methanation with an
allothermal smart-lab scale gasification system. Available: http://www.energetische-
biomassenutzung.de/fileadmin/user_upload/Optmierung_Achental/Dokumente/2009.Kienberger.
Mathanation_with_an_allothermal_smart_lab-scale_gasification_system.pdf [Accessed 23
August 2013].
KIM, G.-Y., MAYOR, J. R. & NI, J. 2005. Parametric study of microreactor design for water gas shift
reactor using an integrated reaction and heat exchange model. Chemical Engineering Journal,
110, 1-10.
KIRNBAUER, F. & HOFBAUER, H. 2011. Investigations on Bed Material Changes in a Dual

Fluidized Bed Steam Gasification Plant in G ssing, Austria. Energy & Fuels, 25, 3793-3798.
KITZLER, H., PFEIFER, C. & HOFBAUER, H. 2012. Gasification of Different Kinds of Non-
Woody Biomass in a 100kW Dual Fluidized Bed Gasifier. 21st International Conference of
Fluidized Bed Combustion. Naples, Italy: EnzoAlbani Editore.
KNOEF, H. E. M. S. A. H. A. M. 2010. Small Scale Gasification Systems. The Netherlands: Biomass

Technology Group BV.
KOHL, L. & NIELSON, R. 1997. Gas Purification- Alkanolamine for Hydrogen Sulfide and Carbon
Dioxide Removal, Gulf Professional Publishing.
Lau, H. L., Choo, Y. M., Ma, A. N., & Chuah, C. H. (2008). Selective extraction of palm carotene and
vitamin E from fresh palm-pressed mesocarp fiber (Elaeis guineensis) using supercritical CO2.
Journal of Food Engineering , 84, 289-296.
Li, J. C. (2005). Radial-Flow Packed-Bed Reactors. Ullman's Encyclopedia of Industrial Chemistry ,
1-14.
Lide, D. (2004). The CRC Handbook of Chemistry and Physics (85 ed.). CRC Press.
Lima, D. F., Zanella, F. A., Lenzi, M. K., & Ndiaye, P. M. ( 2012). Modeling and Simulation of
Water Gas Shift Reactor: An Industrial Case. In D. V. Patel (Ed.), Petrochemicals (pp. 54-74).
Rijeka: InTech.
Loo, S. V., & Koppejan, J. (2008). The Handbook of Biomass Combustion and Co-firing. London:
Earthscan.
403
Full Report
LAURENCE, L. C. & ASHENAFI, D. 2012b. Syngas Treatment Unit for Small Scale Gasifiscation-
Application to IC Enginer Gas Quality Requirement. Journal of Applied Fluid Mechanics, 5, 95-
103.
LEVIN, D. B. 2012. State of the Art and Progress in Production of Biohydrogen. 1-258.
LI, D. H., FINNEY, D. K., SWITHENBANK, J. & SHARIFI, V. N. 2010. EPSRC Thermal
Management of Industrial Processes: A Review of Dring Technologies Sheffield University:
Sheffield University Waste Incineration Centre.
LIMA, D., ZANELLA, F., LENZI, M. & NDIAYE, P. 2012. Modeling and Simulation of Water Gas
Shift Reactor: An Industrial Case. Petrochemicals, 53-74.
LINDE 2011. Linde's Advanced Concept For Ammonia Plants. Chemical Industry Digest, 45-48.
M.A.Fahim, T.A.Al-Sahhaf, & A.S.Elkilani. (2010). Fundamentals of Petroleum Refining (1st ed.).
Amsterdam: Elsevier.
Matar, S., & Hatch, L. F. (2001). Chemistry of Petrochemical Processes (2nd ed.). Massachusetts:
Gulf Professional.
Maxwell, G. R. (2005). Synthetic Nitrogen Products: A Practical Guide to the Products and
Processes (1st ed.). New York: Springer.
McKendry, P. (2001). Energy production from biomass (part 3): gasification technologies.
Bioresource Technology , 83, 55-63.
Mendes, D., Mendes, A., Madeira, L. M., Iulianelli, A., Sousa, J. M., & Basile, A. (2009). The water-
gas shift reaction: from conventional catalytic systems to Pd-based membrane reactors – a
review. Asia-Pacific Journal of Chemical Engineering , 5, 111-137.
Metz, B., Davidson, O., Coninck, H., Loos, M., & Meyer, L. (2005). Carbon Dioxide Capture and
Storage. New York: Cambridge University Press.
Mokhatab, S., & Poe, W. A. (2012). Handbook of Natural Gas Transmission and Processing (2nd
ed.). Oxford: Elsevier.
Mondal, M. K., Balsora, H. K., & Varshney, P. (2012). Progress and trends in CO 2
capture/separation technologies: A review. Energy , 46, 431-441.
M.J.TUINER, ANNALAND, M. V. S., G.J.KRAMER & J.A.M.KUIPERS 2010. Cryogenic CO2
capture using dynamically operated packed beds. Chemical Engineering Science, 65, 6.
M.WANG, A.LAWAL, P.STEPHENSON, J.SIDDERS, C.RAMSHAW & H.YEUNG 2011. Post-

combustion CO2 capture with chemical absorption: A state-of-the-art review. Chemical
Engineering Research and Design, 89, 22.
MANIATIS, K. 2001. Progress in Biomass Gasification: An Overview. Brussels, Belgium.
MODAK, J. M. 2002. Haber Process for ammonia synthesis. 69-77.
MOTTOS, M. D., SOUZA, V. M. & SCHMAL, M. Supported Nickel Catalyst For Steam Reforming
of Methane. Barzil: Enpromer.
ndsu.edu. (2011). Managing The Risk. Retrieved 17 August, 2013, from

http://www.ag.ndsu.edu/pubs/ageng/safety/ae1149.pdf
Newsome, D. S. (1980). The Water-Gas Shift Reaction. Catalysis Reviews: Science and Engineering ,
21 (2), 275-318.
NELES. 2011. Amine plant-Absorption. Available:
http://valveproducts.metso.com/neles/ApplicationReports/2721_Refinery/2721_22_01en.pdf
[Accessed 17 August 2013].
404
Full Report
O.EISA & M.SHUHAIMI 2010. Thermodynamic study of hot potassium carbonate solution using
aspen plus. World Academy of Science, Engineering and Technology, 38, 5.
Olajire, A. A. (2010). CO 2 capture and separation technologies for end-of-pipe applications-A

review. Energy , 35, 2610-2628.
Owlnet. (1997). Ammonia Synthesis and Refrigeration. Retrieved 27 August, 2013, from
http://www.owlnet.rice.edu/~ceng403/nh397heat2.htm
Owlnet. (1979). Reactor Project:Ammonia Synthesis. Retrieved 27 August, 2013, from
http://www.owlnet.rice.edu/~ceng403/nh3syn97.html
Padurean, A., Cormos, C. C., & Agachi, P. S. (2012). Pre-combustion carbon dioxide capture by gas–
liquid absorption for Integrated Gasification Combined Cycle power plants. International
Journal of Greenhouse Gas Control , 7, 1-11.
Phillip, W. M. (2007). Selective Membrane Separation for Ammonia. Retrieved 20 August, 2013,
from http://www.nt.ntnu.no/users/skoge/prost/proceedings/aiche-2008/data/papers/P140228.pdf
Potashcorp. (2013). Potashcorp. Retrieved 30 August, 2013, from
http://www.potashcorp.com/overview/nutrients/nitrogen/overview/world-uses-and-top-producers
Prasertsan, S., & Prasertsan, P. (1996). Biomass Residues from Palm Oil Mills in Thailand: An
Overview on Quantity and Potential Usage. Biomass and Bioenergy , 11 (5), 287-295.
Price, S. (2007). Vapor-Compression Refrigeration. Retrieved 1 September, 2013, from http://ffden-
2.phys.uaf.edu/212_spring2007.web.dir/sedona_price/phys_212_webproj_refrigerators.html
PADBAN, N. & BECHER, V. 2005. Clean Hydrogen-rich Synthesis Gas. Literature and State-of-the-
Art review
PARK, J.-H., KIM, J.-N. & CHO, S.-H. 2010. Performance Analysis of Four-Bed H2 PSA Process
Using Layered Beds. Korea Institute of Energy and Research, 1-13.
PLATON, A. & WANG, Y. 2010. Hydrogen and Syngas Production and Purification Technologies.
American Institue of Chemical Engineers, 311-328.
PORUBOVA, J., BAZBAUERS, G. & MARKOVA, D. 2011. Modeling of the Adiabatic and
Isothermal Methanation Process. Environmental and Climate Technologies, 6.
POTASHCORP 2011. World Ammonia Production Profile.
PROPERTY, L. M. 2007. Ammonia Refrigeration. Rish Management Guide. Liberty Mutual Property.
Rase, H. (1977). Case Studies and Design Data. In Chemical Reactor Design for Process Plants (pp.
325-440). United Kingdom: John Wiley and Sons.
Riegel, E. R. (2010). Kent and Riegel's Handbook of Industrial Chemistry and Biotechnology (11th
ed.). (J. A. Kent, Ed.) New York: Springer.
Roberts, W. (1999). Analysis of Boiling Liquid Expanding Vapor Explosion. Retrieved 2 September,
2013, from
http://www.efcog.org/wg/sa/docs/minutes/archive/2000%20Conference/papers_pdf/roberts.pdf
Rushing, S. (2013). Ethanol Producer Magazine. Retrieved 4 September, 2013, from
http://www.ethanolproducer.com/articles/7674/carbon-dioxide-apps-are-key-in-ethanol-project-
developments
RABIEI, Z. 2012. HYDROGEN MANAGEMENT IN REFINERIES. Petroleum & Coal, 54, 357-368.
RATNASAMY, C. & WAGNER, J. P. 2009. Water Gas Shift Catalysis. Catalysis Reviews, 51, 325-
440.
ROOS, C. J. 2008. Biomass Drying and Dewatering for Clean Heat & Power Washington, USA:
Northwest CHP Application Centre.
405
Full Report
RUSHING, S. A. 2010. Carbon Dioxide – A friend in industry and an environmental foe [Online].
Advanced Cryogenics. [Accessed 4th September 2013].
Safeworks. (2007). Hierachy of Control Measures. Retrieved 27 August, 2013, from

http://www.safework.sa.gov.au/contentPages/EducationAndTraining/HazardManagement/Electri
city/TheAnswer/elecAnswerHierarchy.htm
Schl€ogl, R. (1991). Catalytic Ammonia Synthesis. New York and London: Plenum Press.
Schnitkey, G. (2013). farmdocdaily. Retrieved 30 August , 2013, from
http://farmdocdaily.illinois.edu/pdf/fdd060813.pdf
Seemann, M. (2006). Methanationofbiosyngasin a fluidized bed reactor. evelopment ofa one-step
synthesisprocess,featuring , 124.
Shimekit, B., & Mukhtar, H. (2012). Natural Gas Purification Technologies – Major Advances for
CO2 Separation and Future Directions. Universiti of Teknologi Petronas, Malaysia. Croatia:
InTech.
Smith, B., Loganathan, M., & Shantha, M. S. (2010). A Review of the Water Gas Shift Reaction
Kinetics. International Journal of Chemical Reactor Engineering , 8 (4), 1-31.
Smith, K. H., Anderson, C. J., Tao, W., Endo, K., Mumford, K. A., Kentish, S. E., et al. (2012). Pre-
combustion capture of CO 2 —Results from solvent absorption pilot plant trials. International
Journal of Greenhouse Gas Control , 10, 64-73.
Spinelli, R., Hartsough, B. R., & Magagnotti, N. (2005). International Journal of Forest Engineering.
Testing Mobile Chippers for Chip Size Distribution , 16 (2), 29-35.
Spiraxsarco. (2013). Flash Vessel. Retrieved 28 August, 2013, from www.spiraxsarco.com
Svihusa, B., Kløvstada, K., Pereza, V., Zimonjaa, O., Sahlströmb, S., Schüllerc, R., et al. (2004).
Animal Feed Science and Technology. Physical and nutritional effects of pelleting of broiler
chicken diets made from wheat ground to different coarsenesses by the use of roller mill and
hammer mill , 117 (3-4), 281-293.
S.WONG. 2006. CO2 compression and transportation to storage reservoir. Available:
http://science.uwaterloo.ca/~mauriced/earth691-
duss/CO2_Materials_From_ARC_APEC_Beijing_2006/CarSeq_Module4.pdf [Accessed 3
September 2013].
SALAM, P. A., KUMAR, S. & SIRIWARDHANA, M. 2010. The Status of Biomass Gasification in
Thailand and Cambodia. Mekong Asian Institure of Technology.
SCHMID, J. C., WOLFESBERGER, U., KOPPATZ, S., PFEIFER, C. & HOFBAUER, H. 2012.
Variation of feedstock in a dual fluidized bed steam gasifier-influence on product gas, tar content,
and composition. Environmental Progress & Sustainable Energy, 31, 205-215.
SCOTT, P. 2003. What is an evaporative condenser? [Online]. [Accessed 2nd September 2013].
SMITH, B. R. J., LOGANATHANY, M. & SHANTHA, M. S. 2010. A Review of the Water Gas
Shift Reaction Kinetics. International Journal of Chemical Reactor Engineering, 8, 1-32.
SPATH, P. L., AMOS, W. A. & MANN, M. K. 2002. Process Analysis Work for the DOE Hydrogen
Program - 2001. Hydrogen, Fuel Cells and Infrastructure Technologies, 554-557.
SVENDSEN, G. T. 1998. Towards a CO2 market in the EU: the case of electric utilities. European
Environment, 8, 9.
SWANSON, R. M. 2009. Techno-economic Analysis of Biomass-to-liquids production based on

Gasification. Iowa State: Iowa State University Graduate College.
406
Full Report
Tabi, A. N., Zakil, F. A., Fauzai, W. N., Ali, N., & Hassan, O. (2008). The Usage of Empty Fruit
Bunch (EFB) and Palm Pressed Fibre (PPF) as Substrates for the Cultivation of Pleurotus
Ostreatus. Technology Journal , 49, 189-196.
Tanks, B. (2000). Ammonia Tank, Anhydrous Ammonia Storage Tank. . Retrieved 27 August, 2013,
from www.ammoniatank.com
The Linde Group. (2013). Linde ENgineering Malaysia. Retrieved 28 8, 2013, from http://www.linde-
india.com/userfiles/image/File/Linde%20Isothermal%20Reactor.pdf
T.DONNELLY, S. 2006. The use of MDEA and mixtures of amine for bulk CO2 removal Bryan
Research Engineering, 9.
TECHNOLOGIES, G. R. 2012. R717 vs R404A: Do the advantages outweigh the disadvantages?

[Online]. Ammonia 21. Available: http://www.ammonia21.com/news/view/3717 [Accessed 31st
August 2013].
UN Industrial Development Organization, I. F. (1998). Fertlizer Manual . Dordrecht: Springer.

UNDERWOOD, J. 1997. Design of a CO2 absorption system in an ammonia plant. Available:
http://www.owlnet.rice.edu/~ceng403/co2abs.html [Accessed 26 August 2013].
USHER, M. 2011. CO2 compression report.American electric power mountaineer CCS II project.
American Electric Power, 30.
VMEprocess. (2012). VME Process. Retrieved 29 August, 2013, from

http://www.vmeprocess.com/products.cfm?mode=link&form=58&cat=15
V.ALONSO, T. 2010. High Efficiency on CO2 Removal in Natural Gas with UCARSOL Solvent. Rio
Oil & Gas Expo and Conference.
VASAREVICIUS, S. 2011. PlasTep. Available:

http://www.plastep.eu/fileadmin/dateien/Events/2011/110725_Summer_School/Vasarevicius_Ai
r_Cleaning_PlasTEP__2_.pdf.
VEAL, G. & MOUZAS, S. 2012. Market-based responses to climate change: CO2 market design
versus operation. SAGE, 33, 29.
VICTORIA, S. 2009. Compressed Air System. Energy Efficiency Best Practice Guide [Online].
VREUGDENHIL, B. J. & ZWART, R. W. R. 2009. Tar Formation in Pyrolysis and Gasification.

Netherland: Energy Research Centre of the Netherland.
WARNECKE, R. 2000. Gasification of biomass: comparison of fixed bed and fluidized bed gasifer.
Biomass and Bioenergy, 2000, 489-497.
WORLEY, M. 2011. Biomass Drying Technology Update. Atlanta, GA: Harrie Group Inc.
YU, C.-H., HUANG, C.-H. & TAN, C.-S. 2012. A review of CO2 capture by absorption and
adsorption. Aerosol and Air Quality Research, 12, 25.
Zainac, Z., & Abdullah, Z. (2002). Briquetting of palm fibre and shell from the processing of palm
nuts to palm oil. Biomass and Bioenergy , 22 (6), 505-509.
407
Full Report
Chapter 3
A.CENGEL, Y. & A.BOLES, M. 2007. Thermodynamics, United States, McGraw-Hlli.
AEA Technology, 2000. Thermodynamics and HYSYS, United Kingdom: AEA Technolgy Plc..
Felder, R. M. & Rousseau, R. W., 2005. Elemeatry Principles of Chemical Processes. 3 ed. New
Jersey: John Wiley and Sons, Inc..
FOGLER, H. S. 2005. Elements of Chemical Reaction Engineering. Fourth ed. UK.
GAO, J., YINGLI WANG, YUAN PING, HU, D., XU, G., GU, F. & SU, F. 2012. A thermodynamic
analysis of methanation reactions of carbon oxides for the production of synthetic natural gas.
RSC Advances, 2, 2358–2368.
Kunjunny, A. M., Patel, M. R. & Navin, N., 1999. Revamping if CO2 removal section in ammonia
plant at IFFCO Kalol. Fertiliser News, 44(8), pp. 53-55.
M.FELDER, R. & W.ROUSSEAU, R. 2005. Elementary Principles Of Chemical Process, Wiley.
Proust, P. & Vera, J. H., 2009. PRSV: The stryjek-vera modification of the peng-robinson equation of
state. Parameters for other pure compounds of industrial interest. The Canadian Journal of
Chemical Engineering, 67(1), pp. 170-173.
Sinnot, R. & Towler, G., 2009. Chemical Engineering Design. 5 ed. Oxford: Elsevier Ltd..
Smith, B., Loganathan, M. & Shantha, M. S., 2010. A Review of the Water Gas Shift Reaction
Kinetics. International Journal of Chemical Reactor Engineering, 8(4), pp. 1-31.
Chapter 4
BRENNAN, D. 2012. Sustainable Process Engineering: Concepts, Strategies, Evaluation and

Implementation, Pan Stanford.
GOEDKOOP, M., REINOUT, H., HUIJBREGTS, M., SCHRYVER, A., STRUIJS, J. & ZELM, R.
2008. ReCiPe 2008.
JINENZ-GONZALES, C., KIM, S. & OVERCASH, M. R. 2008. Methodology for Developing Gate-
Gat Life Cycle Inventory information. LCA Methodology, 1.
MAHLIA, T. M. I. 2001. Emissions from electricity generation in Malaysia. Renewable Energy 27,
293-300.
NETL. 2010. Extraction of Natural Gas Life Cycle Inventory.
PACKAGING, G. 2010. The Oil Palm Tree [Online]. Malaysia: GreenTree Packaging Pte. Ltd.
Available: http://www.greentreepackaging.com.sg/info.asp?id=153 [Accessed 25 th October
2013].
SMITH, A. G. 2008. Goodyear Luxembourg Tires. Goodyear Luxembourg Tires.
408
Full Report
Chapter 5
Alvis, R. S., Hatcher, N. A. & Weiland, R. H., 2012.

CO2 Removal from Syngas Using Piperazine‐Activated MDEA and Potassium Dimethyl Glycinate.
Optimized Gas Treating, Volume 1, pp. 1-10.
Callaghan, C. A., 2006. Kinetics and Catalysis of the Water-Gas-Shift Reaction: A Microkinetic and
Graph Theoretic Approach , Worcester: Worcester Polytechnic Institute.
Delmarlearning, 2006. Piping and Vessel. [Online]

Available at: http://webtools.delmarlearning.com/sample_chapters/1418030678_ch03.pdf
[Accessed 29 September 2012].
Eigenberger, G., 1992. Fixed-Bed Reactors. Ullmann's Encyclopedia of Industrial Chemistry, Volume
4, pp. 200-237.
Fujita, S.-I. & Takezawa, N., 1997. Difference in the selectivity of CO and CO2 methanation
reactions. Chemical Engineering Journal, Volume 68, pp. 63-68.
Gandy, D., 2007. Carbon Steel Handbook. 1st ed. USA: Electric PowerResearch Institute.
HaldorTopsøe, 2013. Haldor Topsøe. [Online]

Available at: http://www.topsoe.com/business_areas/ammonia/processes/methanation.aspx
[Accessed 7 October 2013].
Hawkins, G. B., 2009. Methanator Catalyst and Operation, Chicago: GBHEnterprises.
Jakobsen, H., 2008. Chapter 11: Packed Bed Reactors. In: Chemical Reactor Modeling. Berlin :
Springer, pp. 953-984.
Koch-Glitsch, 2010. Intalox Packed Tower Systems: Packed Tower Internals, USA: Koch-Glitsch.
Koch-Glitsch, 2012. Mist Eliminator Liquid-Liquid Coalescing, USA: Koch-Glitsch.
Kohl, A. & Nielsen, R., 1997. Gas Purification. 5th ed. Texas: Gulf Publishing Company.
Lima, D. F. B., Zanella, F. A., Lenzi, M. K. & Ndiaye, P. M., 2012. Modeling and Simulation of
Water Gas Shift Reactor: An Industrial Case. In: D. V. Patel, ed. Petrochemicals. Rijeka: InTech, pp.
54-74.
Morabiya, P. Y. & Shah, P. J. A., 2012. Modeling and SImulation of Water Gas Shift Reaction.
International Journal of Scientific Engineering and Technology, 1(3), pp. 106-110.
Nichols, M. E., 2012. 20 - Paint Weathering Tests. In: Handbook of Environmental Degradation of
Materials. 2nd ed. Oxford: William Andrew Publishing.
Pilkington Insulation & Willoughby, J., 2003. 30 - Insulation. In: M. H. H. H. D. A. S. Eur Ing CEng,
ed. Plant Engineer's Reference Book (Second Edition). Oxford: Butterworth-Heinemann.
409
Full Report
Pilling, M. & Holden, B. S., 2009. Choosing Trays and Packings for Distillation, USA: Sulzer
ChemTech.
Rhodes, M., 2008. Introduction to Particle Technology. 2nd ed. Melbourne: John Wiley.
Sinnot, R. & Towler, G., 2009. Chemical Engineering Design. 5th ed. Burlinton: Elsevier.
Sinnott, R. & Towler, G., 2009. Chemical Engineering Design. 5 ed. Oxford: Elsevier Ltd..
Sinnott, R. & Towler, G., 2009. Chemical Engineering Design. USA: Elsevier l.
Sinnott, R. & Towler, G., 2009. Chemical Engineering Design. 5th ed. USA: Coulson & Richardson's
Chemical Engineering Series.
Smith, B., Loganathan, M. & Shantha, M. S., 2010. A Review of the Water Gas Shift Reaction
Kinetics. International Journal of Chemical Reactor Engineering, 8(4), pp. 1-31.
Zhang, W., Thompson, K. E., Reed, A. H. & Beenken, L., 2006. Relationship between packing
structure and porosity in fixed beds of. Chemical Engineering Science, 61(24), pp. 8060-8074.
Chapter 6
Chapter 7
BAUSBACHER, E. & HUNT, R. 1993. Process Plant Layout and Piping Design. 1-460.
DAUIR, Z. 2010. Tank Farm Design Basics. 1-25.
ELLIOTT. 2013. Multi-Stage Centrifugal Compressors [Online]. Available: http://www.elliott-

turbo.com/Files/Admin/Literature/compressors.pdf [Accessed 30/10/2013.
FILEDER. 2013. SPECTRUM Inox PFH-SPC-52 Round Housing Drawing [Online]. Available:
http://www.fileder.co.uk/SPECTRUM%20Inox%20PFH-SPC-
52%20Round%20Housing%20Drawing.pdf [Accessed 28/10/2013.
GAP 2001. Oil And Chemical Plant Layout And Spacing. GE Global Asset Protection Guidelines, 1-
13.
KLM 2007. KLM Technoogy Group Flare Selection and Sizing. 1-54.
PIONEER. 2013. Standard Centrifugal Pumps [Online]. Available:

http://www.pioneerpump.com/pumps/Standard-Centrifugal.aspx [Accessed 30/10/2013.
TOOLBOX, T. E. 2013. Ceramic Material Properties [Online]. Available:

http://www.engineeringtoolbox.com/ceramics-properties-d_1227.html [Accessed 6 October
2013 2013].
410
Full Report
Chapter 8
AAFC, 2008. Emerging Market Opportunites and Trends, in: Canadian Fertilizer Products
Forum. Canada, pp. 1–45.
Albany, N., 2013. Global ammonia market will reach 160093693 tons in 2020- Food security
concerns drivind demand for ammonia-based fertilizres [WWW Document]. ResearchMoz. URL
http://www.culrav.org/pr/global-ammonia-market-will-reach160093693-tons-in-2020-food security-
concerns-driving-demand-for-ammoniabased-fertilizers.php (accessed10.22.13).
Brennan, D., 1998a. Process Industry ECONOMICS.
Brennan, D., 1998b. Process Industry Economics. IChemE, United Kingdom.
Deloitte, 2012. International Tax and Business Guide- Connecting you to worldwide
information. Malaysia.
Schulze, S., 2012. Ceresana Analyzes the Global Ammonia Market [WWW Document]. Ceresana
Research. URL http://www.chemanager-online.com/en/topics/chemicalsdistribution/ceresana
analyzes-global-ammonia-market (accessed 10.22.13).
Sinnot, R., Towler, G., 2009a. Chemical Engineering Design.
Sinnot, R., Towler, G., 2009b. Chemical Engineering Design, fifth edit. ed. Elsevier Ltd.
Chapter 9
APPL, M., 2000, Ammonia. Ullmann's Encyclopedia of Industrial Chemistry. Wiley-VCH

Verlag GmbH & Co. KGaA.
PIVOT, I. 2013. Louisiana Ammonia Plant. 1-18.

FULL Report - Alternis Monash PDF

Uploaded by

Copyright:

Available Formats

FULL Report - Alternis Monash PDF

Uploaded by

Document Information

Original Title

Copyright

Available Formats

Share this document

Share or Embed Document

Sharing Options

Did you find this document useful?

Is this content inappropriate?

Copyright:

Available Formats

FULL Report - Alternis Monash PDF

Uploaded by

Copyright:

Available Formats

What is the scope of the design project?

What is the scope of the design project?

What are the objectives of the processing plant?

What are the objectives of the processing plant?

MONASH UNIVERSITY

IEM Chemical Engineering

Jenny Yap Wee Li

Lydia Yap Li-Ya

Fatimah Azizah Riyadi

CHAPTER 1 | FEASIBILITY REPORT .......................................................................................................... 1

CHAPTER 1 | FEASIBILITY REPORT

The conventional method of synthesizing ammonia from hydrogen and nitrogen

1.2 Processing Objectives

The processing objectives of Alternis BioAmmonia plant are to optimize sources of

Furthermore, Alternis BioAmmonia aims to maximize the ammonia recovery through

1.3 Feedstock Specification

1.4 Product Specifications

Table 1.4.1: Anhydrous Ammonia product specification

Ammonia wt % 99.5 99.8

Water wt % 0.2 0.5

BL run down temperature (oC) -30 -25

BL run down pressure (kPa g) 200

Boiling point (oC) -33.4

Melting point (oC) -77.7

Flammable limits (by volume in air) 16% to 25%

Autoignition temperature (oC) 651

Solubility in water (per 100g water) 51.0 g at 20oC

1.5 Scope of Design

 Site selection in which the location of industrial site is determined by considering

 Piping and instrumentation diagram is developed through analyzing and decision

 A safety evaluation is conducted describing layers of protection and thus a HAZOP

By implementing the manufacture of ammonia from biomass there can be a reduction in

Figure 1.5.1: Process Block Diagram indicating Inlet and Outlet

1.6 Definition of Terminal Points

1.7 Plant Availability and Capacity

1.8 Feedstock Availability

1.9 Site Characteristic Constraints

1.9.2 Site Characteristics

Moreover, the proposed plant site is at an appropriate location which is located 12 km

1.9.3 Utilities and Storage

Table 1.10.1 Condition Specification and Cost of Supply for Utilities

Liquefied anhydrous ammonia produced is usually stored in 3 different types of tanks

1.10 Feedstock Characteristics

1.11 Market of Product and Byproduct

1.11.1.1 Current Demand

1.11.1.2 Future Trend

1.11.1.3 Selling Price

Figure 2.2.3: Ammonia price trends from 1988-2012

1.11.2 Market of Byproduct

CHAPTER 2 | PROCESS FLOW DIAGRAM AND EQUIPMENT SELECTION

2.2 Equipment Selection

2.2.1 Evaluation of Pre-treatment

2.2.1.1 Technology Evaluation

2.2.1.4 Process Selection

2.2.2 Evaluation of Gasifier

2.2.2.1 Technology Evaluation

Table 2.2.2.1.1: Classification of Gasifiers according to few common categories

Figure 2.2.2.1.2: Gasifier and Combustor Chambers

2.2.3 Evaluation of Post Treatment