Simulation Design Project 2013 PDF
Simulation Design Project 2013 PDF
Simulation Design Project 2013 PDF
Design Project
Cumene Production Plant
A new cumene production plant which produces 100,000 metric tons per year
was designed for the client, Dr Who Chemicals Ltd., is presented in this
project. The variable parameters of the process were reaction size, column
design, heat exchanger design, feed ratio, purity of the propylene feed and
individual stream compositions and conditions. It was required that the plant
be economically viable by determining the optimum values of the variable
parameters.
The major assumptions for the design included steady-state operation and
inert behaviour of the propane impurity. A basis of 100,000 metric tons per
year production of cumene with purity greater than 99.9% was selected as it
was the requirement specified by the client.
HYSYS simulation and hand calculations lead to the determination that the
ideal height and diameter of the benzene column would be 23.77m and
1.707m respectively, the ideal height and diameter of the cumene column
would be 37.17m and 1.286m respectively and the optimal size of the heat
exchanger is 60.32m2.
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Table of Contents
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7.3 Operating Labour........................................................................................ 88
7.4 Raw materials and products ...................................................................... 89
7.5 Economic Appraisal ................................................................................... 90
7.5.2 Discounted Cash Flow Rate of Return (DCF ROR)............................. 90
7.6 Economic conclusions ............................................................................... 92
8 Recommendations......................................................................................... 93
9 References ..................................................................................................... 94
Appendix A – Plant Location Justification .......................................................... 97
Appendix B – Means End Analysis ..................................................................... 99
Appendix C - Key Chemical Information and Associated Hazards .................104
C1. Propylene ..................................................................................................104
C2. Propane .....................................................................................................107
C3. Benzene.....................................................................................................110
C4. Cumene .....................................................................................................114
C5. Di-Isopropyl Benzene ...............................................................................117
Appendix D - Environmental Impacts ................................................................120
D1. Propylene ..................................................................................................120
D2. Propane .....................................................................................................120
D3. Benzene.....................................................................................................120
D4. Cumene .....................................................................................................120
D5. Di-isopropyl Benzene ...............................................................................121
Appendix E – HAZOP..........................................................................................122
Appendix F – Industrial Catalyst Options ..........................................................123
Appendix G - Kinetics Conversions ...................................................................125
Appendix H – Energy balance calculations.......................................................126
Appendix I - Reactor Performance: Feed ratio .................................................128
Appendix J - Reactor Sizing ...............................................................................129
Appendix K – Column Sizing Calculations ........................................................131
K.1 Benzene Column.......................................................................................131
K.2 Cumene Column .......................................................................................138
Appendix M – Economic Evaluation ..................................................................141
M.1 Fixed Capital Investment .........................................................................141
M.2 Spread sheets for the utilities ..................................................................143
M.3 Labour Requirements for the plant .........................................................145
M.4 Raw materials and products ....................................................................146
M.5 Summary of all costs ................................................................................146
M.6 Net Present value calculation tables ......................................................149
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M.7 Discounted Cash Flow Rate of Return Calculation ...............................150
M.8 Payback period .........................................................................................151
M.9 Return on Investment ...............................................................................152
Appendix N – Design Project Meeting Minutes ................................................153
Meeting Minutes (Week 1) ..............................................................................153
Meeting Minutes (Week 2) ..............................................................................154
Meeting Minutes (Week 3) ..............................................................................155
Meeting Minutes (Week 4a) ............................................................................156
Meeting Minutes (Week 4b) ............................................................................158
Meeting Minutes (Week 5a) ............................................................................159
Meeting Minutes (Week 5b) ............................................................................160
Meeting Minutes (Week 6) ..............................................................................161
Meeting Minutes (Week 7) ..............................................................................162
Meeting Minutes (Week 8) ..............................................................................163
Meeting Minutes (Week 9) ..............................................................................164
Meeting Minutes (Week 10) ............................................................................165
Meeting Minutes (Week 11) ............................................................................166
Meeting Minutes (Week 11) ............................................................................167
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Table of Figures
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Figure 37: Case study results for reactor length and diameter sizing ............. 57
Figure 94: Heat Exchanger E-101 ...................................................................... 58
Figure 39: Plate spacing ...................................................................................... 63
Figure 40: Multipass Tray Diagram ..................................................................... 64
Figure 41: Spread sheet 1C ................................................................................ 65
Figure 42: Splitter used to replace depropaniser column ................................. 68
Figure 43: Gibbs reactor and maximum conversion ......................................... 68
Figure 44: HYSYS DataBook case study setup for reactor volume and
conevrsion ............................................................................................................. 69
Figure 45: HYSYS case study results; optimum reactor volume ..................... 69
Figure 46: Reactor volume vs conversion .......................................................... 70
Figure 47: Defining variables bounds in HYSYS reactor sizing case study.... 71
Figure 48: HYSYS DataBook setup for reactor diameter and length case study
................................................................................................................................ 71
Figure 49: HYSYS case study results; reactor diameter and length with
associated conversion.......................................................................................... 72
Figure 50: HYSYS plot of relationship between diameter, length and
conversion ............................................................................................................. 72
Figure 51: PFR setup in HYSYS; defining calculated size ............................... 73
Figure 52: Main and side reaction conversion for pure Propylene feed .......... 73
Figure 53: Mass flow of Cumene produced with pure Propylene feed............ 74
Figure 54: HYSYS reactor sizing input ............................................................... 75
Figure 55: HYSYS DataBook case study setup ................................................ 76
Figure 56: HYSYS adiabatic PFR reactor .......................................................... 76
Figure 57: Adiabatic reactor relationship of inlet and outlet temperatures ...... 77
Figure 58: HYSYS case study setup .................................................................. 78
Figure 59: HYSYS plot of temperature effect on reaction conversion ............. 78
Figure 60: Relationship between main reaction conversion and inlet
temperature ........................................................................................................... 79
Figure 61: Relationship between side reaction conversion and inlet
temperature ........................................................................................................... 79
Figure 62: HYSYS PFR reactor with introduces 'SET' stream ......................... 80
Figure 63: HYSYS isothermal case study setup................................................ 80
Figure 64: HYSYS plot of temperature effect on reaction conversion ............. 81
Figure 65: Relationship between main reaction conversion and inlet
temperature ........................................................................................................... 81
Figure 66: Relationship between side reaction conversion and inlet
temperature ........................................................................................................... 82
Figure 67: Payback period for the impure feed Cumene plant ........................ 91
Figure 68: Structure of Propylene ......................................................................104
Figure 69: Propylene phase diagram from HYSYS ..........................................106
Figure 70: Structure of Propane .........................................................................107
Figure 71: Propane phase diagram from HYSYS ............................................109
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Figure 72: Structure of Benzene ........................................................................110
Figure 73: Benzene phase diagram from HYSYS ............................................113
Figure 74: Structure of Cumene .........................................................................114
Figure 75: Cumene phase diagram from HYSYS ............................................116
Figure 76: Structure of Di-isoprpyl Benzene .....................................................117
Figure 77: DIPB phase diagram from HYSYS ..................................................119
Figure 78: Benzene and Propylene feed ratio relationship with conversion ..128
Figure 79: Benzene and Propylene feed ratio relationship with Cumene to
DIPB selectivity ....................................................................................................128
Figure 80: HYSYS calculate volume and associated conversion ...................129
Figure 81: Relationship of reactor length and diameter with conversion .......130
Figure 82: Spread sheet - 1D .............................................................................134
Figure 83: Spread sheet - 1C .............................................................................135
Figure 84: Spread sheet - 1H .............................................................................137
Figure 85: Spread sheet - 2C .............................................................................138
Figure 86: Spread sheet - 2D .............................................................................139
Figure 87: Spread sheet - 2H .............................................................................140
Figure 82: Fixed Capital for pure feed case ......................................................141
Figure 83: Fixed capital investment for impure feed case ...............................142
Figure 84: Cost of utilities for the pure feed case .............................................143
Figure 91: Utilities for impure feed case ............................................................144
Figure 92: Labour requirements for both feed cases .......................................145
Figure 93: Raw material costs and product revenues for pure case ..............146
Figure 94: Raw material costs and product revenue for impure case ............146
Figure 95: Summary of costs for pure feed material ........................................147
Figure 96: Summary of costs for impure feed case .........................................148
Figure 97: NPV calculations for the pure feed ..................................................149
Figure 98: NPV calculations for impure feed case ...........................................150
Figure 99: Payback period for pure feed case ..................................................151
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List of Tables
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1 Allocation of Work
1.1 Report
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1.2 HYSYS Simulation
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2 Introduction
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3 Literature Review
CONVENTIONAL PROCESS
A typical conventional process for the synthesis of Cumene involves a packed
bed reactor in either the liquid or gas phase. Pathak et al (2011), considered a
gas phase reaction when evaluating the optimal design of a Cumene
synthesis plant. In their design, the feed streams are mixed with a Benzene
recycle stream before being vaporised and entering a heat exchanger. The
heated stream enters a cooled packed bed reactor (PBR) (the PBR consists
of a shell and tube heat exchanger that is packed with catalyst), and then
enters the first of three columns (Pathak et al, 2011). The first column
separates the Propane and Propylene and removes them as a fuel gas
mixture, with the bottoms stream then feeding into the next column. The
second column separates the flow into the Benzene recycle stream and the
Cumene/ Di-isopropyl Benzene stream, the latter of which flows into the final
column, where it is separated into its components.
DOW-KELLOGG PROCESS
The Dow-Kellogg process is a recognised Cumene synthesis procedure and it
has been used industrially since 1993. Initially, a large excess of Benzene is
mixed with Propylene (Dimian and Bildea (2008) report that the molar ratio
should be larger than 5:1). The first reactor, for alkylation is packed with a
catalyst.
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It can be observed from Figure 1: Process schematic for Dow-Kellogg process
that the product stream from the alkylation reactor enters four consecutive
distillation columns, which separate the fuel gases, the Benzene recycle, the
Cumene, and any other side reaction products. The reaction takes place in
liquid phase, and is generally at between 160 0C and 2400C (Dimian and
Bildea, 2008). There are two recycle streams, one for the Benzene, and the
other recycles some of the alkyBenzenes produced in the final column
(Dimian and Bildea, 2008). These are then recycled into a second reactor.
The key features of this process are that it uses minimal energy, and that the
recycle streams are kept to a minimum.
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Figure 3: Improved Reactive Distillation Column (no rectification
section) (Pathek et al., 2011)
Since Propylene is less dense than Benzene, there may be an issue with
separating these two components. To avoid this problem, Benzene can be
used in excess so that Propylene is fully converted (Dimian and Bildea, 2008).
Propylene is fed into the bottom of the column and Benzene is fed above the
reactive zone to generate counter-current reactant flow (Pathak et al, 2011;
Dimian and Bildea, 2008). Benzene then moves down the column, while
Propylene simultaneously travels up the column. Any Propane (which is
unreactive in the system) in the feed steam will be removed from the column
with unreacted Propylene as a vapour. The heavier components, Cumene
and di-isopropyl Benzene are removed at the bottom of the column (Pathak et
al, 2011). Unreacted Benzene then returns to the reactive area.
The catalyst in the reactive distillation process is a zeolite catalyst, which will
react in liquid phase. Selectivity is enhanced as the temperature decreases,
and a purity of 99.9% for Cumene is achievable. Reactive distillation is 47%
less expensive than the conventional process (Pathak et al, 2011). This
process combines the Benzene recycle, the chemical reaction and the
separation of the Propylene/Propane from other components into one step
(Pathak et al, 2011).
CATALYTIC SEPARATION
Catalytic separation incorporates the reaction and a partial separation in one
column (Buelna and Nenoff, 2005). The major advantage of this technology is
that it uses less energy than both catalytic separation and reactive distillation.
This is because catalytic separation can be operated at normal, atmospheric
pressure, rather than the higher pressures that the distillation processes
require (Buelna and Nenoff, 2005). The dissimilarity of the boiling points of
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the reactants and products can be utilised to remove the unreacted Benzene
from the products as they form.
3.2 Reaction
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A secondary, undesired side reaction occurs simultaneously. This reaction
involves Propylene and Cumene reacting to form p-diisopropyl Benzene. This
reaction is represented by
→
Where = 2.75 x 106
The rate law for this reaction is equal to the product of the rate constant and
the concentration of Propylene and Cumene.
3. Selectivity
As well as low reactor temperature and high pressure, selectivity
for Cumene can be increased if Cumene and Propylene have a
low concentration in the reactor. Thus, Benzene should be present
in excess, and this excess should be recovered from the reactor
and then recycled (Luyben, 2010).
3.3 Catalyst
A catalyst is used in the alkylation reaction between Benzene and Propylene
to form Cumene. By providing an alternative reaction pathway which utilises a
lower activation energy, the catalyst is able to accelerate the rate of reaction
and reduce side reactions without being consumed. Dr Who Chemicals has
attained the patent for a new type of catalyst and specifies that the catalyst
has a particle diameter of 1.5 mm and a density of 1600 kg/m3. However, no
other specifications of the catalyst are known.
The type of catalyst used in this reaction has changed over time. Originally,
acidic catalysts were favoured (solid phosphoric acid and sulfuric acid).
Aluminium chloride and nickel/ ϒ-alumina catalysts have also been used.
Presently, zeolite catalysts are gaining popularity and are increasingly being
used. Zeolite catalysts are made of a silica and alumina tetrahedra organised
in molecular sieves (Dimian and Bildea, 2008). There are several types of
zeolitic catalysts and all feature large pore openings which are used to obtain
high selectivity.
The properties of catalysts depend on both the activation process and the
treatment of the surface. Industrially, most catalysts are used in pellet shape.
The rate determining step of the reaction is the internal diffusion (Dimian and
Bildea, 2008). The catalyst used in the reaction has a direct impact on the
selectivity of the reaction.
The four main types of catalyst that have been investigated are acidic
catalysts, the aluminium chloride catalyst, the nickel/ ϒ-alumina catalyst and
zeolitic catalysts. Below is a summary of the advantages and disadvantages
of the catalysts.
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Table 1: Advantages and disadvantages of common catalysts for the
synthesis of Cumene from Benzene and Propylene
A basic hazard and operability (HAZOP) study has been prepared. This
focuses mainly on the columns and reactor, and gives the action required if
the unit operations are not functioning at the desired conditions. The HAZOP
can be found in Appendix E – HAZOP.
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3.5 Environmental impacts
The environmental impacts of the Cumene plant have been investigated.
Many of the chemicals involved have adverse effects on the environment.
These effects have been evaluated so that any potential adverse
environmental impacts can be negated. It will be important for the plant to
avoid any release of chemicals into the air, water or onto land, and always
dispose of chemicals properly. A full analysis of the environmental impacts of
Benzene, Cumene, di-isopropyl Benzene, Propane and Propylene can be
found in Appendix D - Environmental Impacts.
Demand for Cumene has grown in the last ten years to reflect an increase in
the use of chemicals bisphenol-A, polycarbonate and phenyl resins, for which
Cumene is a raw material (ASD, 2013). These chemicals are used in a range
of industries including pharmaceuticals, construction and packaging.
The growth in the demand for Cumene has not been uniform over the last
decade. During the global financial crisis, the demand for phenol (and thus
Cumene) slumped (Nexant, 2013). There was a rapid recovery in the demand
of most Propylene derivatives in 2010, before once again the demand
weakened in 2011/2012. During this time period, China and India restricted
their monetary policies, many of the European Union members faced debt
concerns and the uncertain political climate in parts of Africa and the Middle
East resulted in very little growth in the need for Cumene (Nexant 2013).
On average, over the last decade, China has experienced a large amount of
growth. 85% of the worldwide demand in 2011 was from China, Japan,
Taiwan or South Korea, firmly establishing the Asia Pacific region as the
largest consumers of Cumene (CM, 2013). The growth in the demand for
Cumene in the Asian markets has counteracted the sluggish growth in
America and Europe (ASD, 2013). Growth in the demand for Cumene has
also been strong in developing countries. This is fortuitous for Dr Who
Chemicals Ltd, as the close proximity of Australia to these countries will result
in a steady demand for their product.
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stoke the demand for Cumene. In September 2013, approval was given in
China for a 650,000 ton phenol/acetone plant to be constructed
(ChemicalOnline, 2013). It will be the largest plant for these chemicals in
China and is expected to be complete in 2016. It will require 550,000 tons of
Cumene per annum (ChemicalOnline, 2013). The plant site is in an industrial
park near railways and waterways; export from Dr Who Chemicals in Australia
to the plant should be possible.
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The climate in this region is adequate. There is little humidity, and no extreme
heat (maximum mean temperature is 25.5⁰C in Geelong) or cold (mean
minimum temperature in winter is 5.1⁰C) conditions in the area. The only
potential danger could be bushfires in summer, and some safeguards may
need to be implemented to avoid safety issues. Events like floods, hurricanes
or other natural disasters are unlikely to occur. The supply of water, due to the
proximity of the ocean and rivers is plentiful.
The final major factor is the community and political considerations. The
region around Geelong already has a lot of manufacturing and construction
sites. Additionally, Victoria’s unemployment rate in June 2013 was 5.5%,
above the national average of 5.1% (Australian Bureau of Statistics, 2013).
The proposed Cumene plant will employ workers not only in the daily
operations, but during the construction phase, many labourers, electricians
etc. will be employed.
The Geelong Ring Road Employment Precinct (GREP) would be ideal for the
plant. It is located north of Geelong, has zoning approval for all industrial
work, and is 500 hectares. It is located 6km from the port, and is close to the
Princes Freeway (GREP, 2013).
Figure 5: Proposed location for the Cumene plant. The area marked
'grep' is the industrial park. This max shows the proximity to the rail line,
wharf, port, airport and the Geelong central business district.
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4 Process Synthesis
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Figure 6: Process Flow Diagram (PFD)
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Figure 6A: HYSYS worksheet stream summary
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4.1.2 Process Fluid Package
The selection of an accurate and most appropriate fluid package was
important in the HYSYS simulation because it reflects what is happening in
the process plant. An incorrect fluid package could result in inadequate model
parameters and hypothetical component generation and cause consistency
problems with the plant data. The key requirement is the selection of the fluid
package was to accurately describe the chemical species and their physical
properties. HYSYS provides the option of Equation of State, Activity
Coefficient, vapour pressure, Semi-empirical and specialty models.
The Peng-Robinson fluid package was selected as it was the most enhanced
model in HYSYS. It is an equation of state model with the largest applicability
range in terms of temperatures and pressure. It also had a wide range of
operating conditions and was suitable for single, two and three phase systems
and had the largest binary interaction parameter database.
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4.2 Reactor
Design specifications, reaction stoichiometry and kinetic data were used to
design the reactor and determine reactor performance from conversion and
selectivity. The client, Dr Who Chemicals required the reaction to be
conducted at a temperature between 200 oC and 400oC and pressure between
20 and 30 bar. Since the specified temperature range is above the boiling
point of both reactants, it is reasonable to assume the reaction will be
operated in vapour phase. Therefore, a plug flow reactor is suggested for the
process. Literature states similar parameters to design process, where a
temperature of 358oC and pressure of 25bar were used for the reactor
(Luyben, 2010). As these conditions met the requirements for the given
design they were used as the initial starting point for the reactor specifications.
A Gibbs reactor was constructed to estimate the maximum conversion that
the reaction could achieve in actual process, and thus obtain a limit to
subsequent calculations. Two case studies of reactor sizing, and temperature
impact to the reactor performance relate to conversion were performed. It was
determined that optimum operating conditions were inlet temperature of
358oC, outlet temperature of 493.6oC and reactor pressure of 25 bar. A
Benzene to Propylene feed ratio of 3:1 and reactor volume of 43m3 (length
18.94m and diameter 1.7m) were calculated to achieve maximum conversion.
An adiabatic, single tube, catalyst packed bed plug flow reactor was used in
the process. The exothermic nature of the reaction which produced Cumene
was considered and resulted in the reactor outlet temperature exceeding the
inlet temperature. It was determined that the reactor was operating for vapour
phase components from the specifications outlined by the client. The effluent
of reactor has a higher temperature and was first liquefied in a condenser
before proceeding through a heat exchanger to recycle the thermal energy to
preheat the reactor feed stream. It was assumed that the catalyst can operate
efficiently up to 500oC. However, in reality a cooling system is recommended
for the reactor due to the exothermic nature of the reaction increasing the
reactor temperature and potentially degrading the catalyst and decreasing
reactor efficiency. The reactor design is shown below in Figure 8: Plug Flow
Reactor (below).
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A reactor inlet temperature range of 275oC to 362oC was recommended
(Luyben, 2010). It was determined that optimal reactor operating temperature
was 365 oC. An inlet temperature of 358 oC was calculated to achieve the
maximum reaction conversion and selectivity. The plug flow reactor conditions
specified in HYSYS can be seen in Figure 9: PFR Worksheet (below).
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thermal energy generated by the exothermal reaction in the reactor to keep
the reactor below 500oC in order to protect the catalyst. Consideration of the
economics of the process lead to the conclusion that increasing the feed ratio
results in more excessive benzene being recycled in the system which
reduces the feed cost. A result of increased feed ratio was related unit size
increasing correspondingly. It was determined that a Benzene to Propylene
feed ratio of 3:1 achieved a reasonable selectivity and resulted in appropriate
reactor size and outlet temperature.
The optimal length of the reactor was determined to be 18.94m and diameter
of 1.7m which results in a volume of 43m3. This produced a conversion of
0.9939 which is less than the Gibbs conversion limit of 0.9942. The HYSYS
sizing of the plug flow reactor is shown below in Figure 10: PFR Rating. Sizing
calculations are discussed in report section 5.2.
REACTION SPECIFICATIONS
The desired and side reactions were set up in HYSYS in the simulation basis
manager. The reaction stoichiometric coefficients were defined in the
‘Stoichiometry’ tab in HYSYS. All reactants had negative stoichiometric values
as they were consumed in the reactions; while products were assigned
positive values as they were produced. Stoichiometric coefficients for each
component were added to the stoichiometry tab as shown in Figure 11:
Reaction stoichiometry for Rxn-1 and Rxn-2.
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Figure 11: Reaction stoichiometry for Rxn-1 and Rxn-2
Propylene was the limiting reactant in the production of Cumene and was
specified as the ‘Base Component’ in the ‘Basis’ tab in HYSYS. The reaction
phase was stated as “VapourPhase” as the specified temperature range is
above the boiling point of both reactants (Propylene 60 oC and Benzene 240oC
at 25bar) so it is reasonable to assume the reaction will be conducted in the
vapour phase. An illustration of Basis set up for two reactions are shown in
Figure 12: Reaction basis for Rxn-1 and Rxn-2.
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Figure 12: Reaction basis for Rxn-1 and Rxn-2
Finally in the ‘Parameters’ tap, Arrhenius constant (A) and activation energy
(E) values were defined. Both of these values were provided as part of the
kinetics data supplied by the client. The units of Arrhenius constant (A) and
the activation energy (E) were required to be converted to SI units. The unit
conversion calculations are provided in Appendix G - Kinetics Conversions.
Figure 13: Reaction parameters for Rxn-1 and Rxn-2 shows both reactions
parameter set ups in HYSYS.
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Figure 13: Reaction parameters for Rxn-1 and Rxn-2
REACTOR PERFORMANCE
The conversion of Propylene in the designed plug flow reactor was calculated
to be 0.9939. After the reactor was designed from the specifications above its
performance could be determined by calculating the reaction conversion using
the following equation:
This equation involved setting the desired Cumene reaction and the side
reaction of producing DIPB to operate simultaneously or sequentially. It was
calculated by connecting reactor feed and product stream and defining
Propylene molar flow rate in the HYSYS spreadsheet. From this the
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conversion value of 0.9939 was obtained as shown in Figure 14: PFR spread
sheet.
In addition, the desired reaction (Rxn-1) had 95.74% conversion and the side
reaction (Rxn-2) had a conversion of 3.645% as shown in Figure 15: PFR
reaction conversion results.
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Selectivity was calculated dividing the desired product by undesired product,
as described by the equation below. Figure 14: PFR spread sheetshows that
the selectivity was determined to be 27.56. This reflects the desired
production rate and directly influences economic benefit.
GIBBS REACTOR
The Gibbs Reactor was used to determine the theoretical maximum
conversion possible for the reaction. A Propylene basis with and Propylene
feed 5%wt impurity of Propane was used in the Gibbs reactor. Figure 16:
Gibbs Reactor shows the Gibbs reactor constructed for the HYSYS simulation.
A ‘set’ function was used to set the molar feed into the Gibbs reactor as that
entering the plug flow reactor in the process and the molar composition of the
stream was manually defined. The Gibbs reactor generated a reaction without
defining a reaction set. A spread sheet was used in HYSYS to calculate the
conversion of the Gibbs reactor. The HYSYS spread sheet is shown in Figure
17: Gibbs reactor spread sheet. A maximum reaction conversion of 0.9942
was calculated using the Gibbs reactor (at 358 oC and 25 bar) and it was
ensured that the plug flow reactor conversion limit did not exceed this value
during sizing calculations.
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Figure 17: Gibbs reactor spread sheet
It was determined that the most appropriate material of construction for the
reactor was stainless steel. This choice has been detailed and explained in
7.1.1 Material of Construction.
The design of the heat exchanger (E-101) was determined by the conditions
required by the reactor. The shell and tube heat exchanger had two input and
two output streams and was operated in the liquid phase. The heat exchanger
was an energy efficient process, as it used the heat from the input streams to
heat up the output streams, and required no additional utilities. This was also
advantageous as there was no cost associated with the operation of the heat
exchanger.
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The primary objective of the heat exchanger is to heat the reactants up to the
temperature required by the reactor. Since the heat exchanger did not heat
the ‘preheat’ stream up to the required temperature, a heater (E-100) was
required immediately after the heat exchanger. The product stream from the
reactor was cooled by cooler E-105 in order to reduce the minimum approach.
The minimum approach for the impure feed case was 59.59°C.
Pressure Drop
The pressure drop of the heater was simulated in the ‘steady state’ mode
(every other property of the heat exchanger was simulated in the ‘end point’
mode.
Figure 18: Pressure drop specifications for the heat exchanger, E-101
Using HYSYS, the tube side pressure drop was calculated to be 10 kPa, and
the shell side pressure drop was calculated to be 19.28 kPa. The shell side
experiences more turbulence as the fluid flows over the shells than the tube
side and this is reflected in the larger pressure drop when compared to the
tube side.
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Figure 19: Overall sizing specifications for heat exchanger, E-101
The overall specifications for the heat exchanger can be seen in Figure 19:
Overall sizing specifications for heat exchanger, E-101. The values for the
overall heat transfer coefficient (U) and overall heat transfer coefficient
multiplied by the area (UA) were used to calculate the area of the heat
exchanger.
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Figure 20: Shell side specifications for heat exchanger, E-101shows the
specifications for the shell side of the heat exchanger. The values in blue
(shell diameter, number of tubes, tube pitch, baffle cut and baffle spacing)
have been inputted from the calculations done in Figure 20: Shell side
specifications for heat exchanger, E-101. The tube layout angle which was
desired was already the default, and so left. It was assumed that the shell
baffle type was ‘single’.
The specifications for the tube side of the heat exchanger can be observed in
Figure 21: Tube side specifications for heat exchanger, E-101 the values in
blue (outer diameter, inner diameter and tube length) have been inputted, and
are based on standard pipes, as discussed in Figure 21: Tube side
specifications for heat exchanger, E-101. The specific heat and density of the
wall have been left empty.
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Figure 22: Temperature/Heat Flow graph for heat exchanger, E-101
The relationship between the temperature and the heat flow can be observed
in Figure 22: Temperature/Heat Flow graph for heat exchanger, E-101. Since
the tube side and shell side temperature profiles are parallel, this shows that
there was no phase change in the heat exchanger. The distance between the
two lines is equal to the minimum approach, which was 59.59°C.
Propylene and Propane have similar boiling points that are much less than
that of Benzene, Cumene and DIPB. Therefore, the unreacted Propylene and
Propane separation occurred first. They were separated as top product with
by the depropaniser distillation column; while the Benzene, Cumene and
DIPB continued as the bottom products through further separation processes.
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Due to similarity in boiling points of Propylene and Propane they were not
further separated and sold as impure fuel gas.
The final separation process was achieved in the Cumene distillation column.
Cumene has a lower boiling point than DIPB so was produced as the top
product; while DIPB was separated into the bottom stream.
40 | P a g e
Table 4: Benzene Column Separation
OPERATING PRESSURE
According to the operation pressure graph (Figure 25: Operation pressure
determination graph), the operation pressure is determined by the bubble
point of the mixture at the temperature of the cooling water used in the
condenser. The condenser pressure was first determined with the use of a
splitter. The HYSYS ‘SET’ function was used to set the molar feed rate into
the splitter as that entering the Benzene distillation column. Parameters of
stream ‘b’ exiting the top of the splitter were used to determine the operating
pressure of the condenser. The pressure of the condenser should be such
that the exiting temperature is at least 5 to 10oC above the bubble point of the
condenser cooling water. Maintaining a vapour fraction of 1 ensures that the
stream temperature is above the boiling point. It was observed that
decreasing the pressure of the stream decreased its temperature while
maintaining a vapour fraction of 1; so the lowest possible pressure would be
optimal. The operation pressure graph was used in conjunction with the
assumption that summer cooling water is 49oC and the bubble point of all
components in stream ‘b’ at 49oC which suggested total condensation.
However, pressures below 1 bar are considered uneconomical and too
difficult to maintain so it was decided that the most appropriate condenser
pressure would be 1 bar for the Benzene distillation column. The
determination of this pressure was not part of the iterative process.
41 | P a g e
diameter of the column. Pressure drop hand calculations determined the
actual reboiler pressure which should be then entered as the new value in the
shortcut column. This was the iterative process in the column sizing
calculations and was performed separately for both the Benzene and Cumene
distillation columns.
42 | P a g e
Figure 26: Conditions of splitter stream b
SHORTCUT COLUMN
A shortcut column in HYSYS was used to calculate the number of trays and
reflux ratio of the distillation column. The ‘SET’ function was used to set the
molar flow into the shortcut column as that into the Benzene distillation
column while the composition of the input stream was manually entered. The
condenser pressure of 1 bar and final converged solution of 1.725 bar for the
reboiler pressure were also defined in the shortcut column. HYSYS then
computes the column reflux ratio and actual number of trays. The results of
HYSYS calculations were viewed in the ‘Performance’ and ‘Design’ tabs of
the shortcut column.
The two major components in the composition of stream 3 were Benzene and
Cumeme. Benzene has a lower boiling point than Cumene at the same
pressure; therefore, Benzene can be identified as the ‘light’ constituent and
Cumene as the ‘heavy’ component. As the purity of the Cumene product is at
43 | P a g e
least 95wt% the tolerace for Benzene exiting the column in the bottom stream
is less than a mole fraction of 0.0001. For the Cumene column, Cumene was
identified as the ‘light component and DIPB as the ‘heavy’ constituent. In
order to achieve the Cumene purity requirement and minimise the reflux ratio
the tolerance of DIPB in the overhead stream was mole fraction of 0.0003.
HYSYS then computes the column reflux ratio and actual number of trays.
The results of HYSYS calculations were viewed in the ‘Performance’ and
‘Design’ tabs of the shortcut column. Under the ‘Performance’ tab of the
shortcut column the HYSYS calculated actual number of trays was viewed in
Figure 28: Shortcut column performance. It was observed that the associated
actual number of trays was 20 for a reboiler pressure of 1.725 bar. It was also
computed by HYSYS that the minimum number of trays was 9 and the
optimum feed stage was 4.
The HYSYS calculated minimum reflux ratio was viewed in the ‘Design
Parameters’ tab of the shortcut column. The external reflux ratio was
calculated as 1.4 times the minimum reflux ratio. This ratio is the column
reflux ratio and was calculated to be 0.763 for the final reboiler pressure of
1.725 bar. Figure 29: Shortcut column design parameters shows the design
parameters of the shortcut column; the blue numbers were input by the user
while the black number have been computed by HYSYS.
44 | P a g e
Figure 29: Shortcut column design parameters
RIGOROUS COLUMN
The information calculated by the shortcut column was used to set up the
rigorous column as shown in Figure 30: Rigorous column connection page. In
the “Monitor” section of the rigorous column design “Add Spec” of the “Comp
Fraction” (as shown in Figure 31: Rigorous column Monitor page and Figure
32: Component Fraction Specifications), the “Reflux ratio” and “Comp Frac”
were selected as the active variables. The degree of freedom then became 0
so HYSYS was able to automatically begin the calculation and continue until a
converged solution was reached. The same procedure was applied to the
Cumene column.
45 | P a g e
Figure 31: Rigorous column Monitor page
46 | P a g e
5 Material/Energy Balances and Equipment Sizing
Reactor Balance
Benzene
Benzene
PFR REACTOR
Cumene
Propylene
(5wt% Propane)
Propane
47 | P a g e
assumed that no Cumene is lost between exiting the reactor and leaving the
process as top product from the Cumene distillation column.
Cumene is formed in the main reaction and consumed in the side reaction.
Therefore, the amount of Cumene exiting the reactor is equal to the amount
produced minus the amount consumed.
In the desired main reaction Propylene is the limiting reagent and the
stoichiometric ratio of Propylene and Cumene is one. Therefore the moles of
Cumene produced are equal to the moles of Propylene reacting in the main
reaction.
Propylene fed into the reactor is consumed in both the desired and side
reaction. It was assumed that all Propylene feed into the reactor is consumed;
therefore it can be concluded that the sum of Propylene reacted in the main
and undesired reaction is the total amount entering the reactor.
The selectivity of Propylene to react and form Cumene over DIPB is 31:1
(Kugler, 1995).
The above generated equations allowed the moles of Propylene entering the
reactor to be calculated using trial and error method varying moles of Cumene
consumed in the side reaction in Excel. Figure 33: Excel Material Balance
Calculations is a screenshot of the Excel document used. The highlighted cell
(C6) represents the mole of Cumene consumed in the side reaction and was
varied until the desired selectivity of 31 was reached. Cumene produced (cell
E3) was a function of Cumene consumed from their relationship defined
above. Propylene (cell C3 and B6) were functions of the Cumene produced
48 | P a g e
and consumed respectively; as defined by the appropriate stoichiometric
relationships.
49 | P a g e
Propylene feed is composed of 95wt% Propylene and 5wt% Propane impurity.
This composition was used to calculate the feed rate of Propylene into the
reactor.
Benzene was fed to reactor in excess, with Propylene to Benzene feed ratio
of 3.
50 | P a g e
Di-isopropyl Benzene was produced in the side reaction as the undesired
product and its flow rate out of the rector was calculated using stoichiometric
relationships.
A material balance was also performed around the columns to determine the
flow rate and compositions of streams exiting the process. The assumption of
complete separation in each distillation column was made to simplify
calculations but was not made during HYSYS simulation. From this
assumption it could be included that all Propane that entered the system into
the reactor exited the process in the top product stream of the depropaniser
51 | P a g e
column; while the remaining products leaving the reactor continued through
the process in the bottom product.
The composition of the top stream was 100% Propane (assumption) and the
composition of the bottom stream was calculated (as below).
TOTAL 188.6194 1
The Benzene column separated Benzene from Cumene and DIPB. The
assumption of complete separation resulted in all Benzene entering the
column leaving in the tops product while the DIPB and Cumene continue
through the process in the bottoms product.
52 | P a g e
The composition of the top stream was 100% Benzene (assumption) and the
composition of the bottom stream was calculated (as below).
The composition of the top stream was 100% Cumene (assumption) and the
composition of the bottom stream was 100% DIPB (assumption).
TOTAL 3.463 1
53 | P a g e
5.1.2 Energy Balance
Assumptions
Weight-averaged values for heat capacities were used for each stream
and calculated from reference data (Perry and Green, 2007)
Ignored pressure effects of streams
Heat capacities used were at constant pressure (
No heat is lost across units
Heat capacity of DIPB is negligible
The energy balances for the main unit operations were calculated by hand
using assumptions for simplification purposes so the results varying from
HYSYS simulated values. These calculations are outlined in Appendix H –
Energy balance calculations and the results are concluded below (Table 11:
Energy balance summary).
54 | P a g e
5.2 Equipment Sizing
The plug flow reactor was sized based on reactor volume. HYSYS was used
to calculate the conversion as a function of reactor volume. A Case Study
function in HYSYS was used to investigate the relationship between reactor
volume and conversion. The PFR reactor volume was selected as the
independent variable and the reaction conversion which was calculated in the
HYSYS spread sheet (cell A4) was selected as the dependent variable.
According to the HYSYS calculated results, a reactor volume of 43m3 has
conversion 0.9911 which was closest to the maximum conversion without
exceeding it. Detailed data is provided in Appendix J - Reactor Sizing.
55 | P a g e
Figure 35: Case study results of reactor volume and conversion
After the reactor volume was determined it was necessary to investigate the
diameter and length which resulted in this fixed volume. The reactor length
affected the reactor resistant time and the diameter effected the flow rate. It
was required to ensure that the reactor tube diameter and length (L/D) ratio
was between 10 and 20. The diameter and length combinations provided by
HYSYS simulation at the determined fixed volume of 43m3 were compared. It
was discovered that the optimal reactor length was 18.94m and diameter of
1.7m, which resulted in a length to diameter (L/D) ratio 11.14. Another case
study investigated that conversion of reactor with length and diameter of 8.9m
and 1.7m respectively was 0.9853 and reactor with length and diameter of
19.0m and 1.7m respectively was 0.9862. Both conversions were less than
the maximum reaction conversion as determined using the Gibbs reactor. The
reactor size of 18.94m and diameter 1.7m was determined to be optimal.
Detail data and explanation is provided in
56 | P a g e
Appendix I - Reactor Performance: Feed ratio.
Figure 36: Case study data selection for reactor length and diameter
sizing
Figure 37: Case study results for reactor length and diameter sizing
57 | P a g e
5.2.2 Heat exchanger sizing
Reactor Out
Reactor Cool
The mass flow rates were calculated from HYSYS. For E-101, the mass flow
rate was 3.33x104 kg/h. The enthalpies (mass based) of the streams
‘MixFeed’ and ‘Preheat’ were also taken from HYSYS, with values of 613.4
kJ/kg and 699.4 kJ/kg respectively.
Based on these values,
( ) ( )
( )
( )
58 | P a g e
The friction factor also needs to be calculated, which involved calculating the
R and S values. R is given by
( )
( )
S is given by
( )
( )
Using the friction factor chart (Zhang, 2013), this corresponds with a friction
factor of 0.814.
Using equation 31 the only remaining unknowns are U and A. Once again,
these values have been calculated on HYSYS. The value of U is given as
1438 kJ/h.m2.0C, and UA is given as 8.672x104 kJ/h.0C. Substituting these
value is the equation 31 gives an area for E-101 of 60.32 m2.
59 | P a g e
Therefore, the number of tubes required is 124.
The fluid velocity was then calculated, to confirm that the tube geometry is
acceptable.
The fluid velocity is given by the mass flow rate divided by the product of the
density and cross sectional area.
( )
( )
Since the fluid velocity is less than 2.4 m/s (Zhang, 2013), this shows that the
tube geometry is acceptable.
√ ( )
CL is the tube layout. The tube layout has been assumed to be the triangular
pitch (30o layout). This has been chosen as it gives a better heat transfer and
a larger surface area than the square pitch layout. CL has a value of 0.87.
CTP is the tube constant. The tube constant is a correction factor which is
used to show the ‘incomplete coverage’ between the shell and outer tube
(Zhang, 2013). This has a value of 0.93 for a one tube pass.
60 | P a g e
The tube pitch , is required to be between 1.25 and 1.5. The tube pitch was
assumed to be the average of this range, and a value of 1.375 was used in
calculations.
Substituting values into equation 38, gives the following
√ ( )
From the diameter of the shell, several important conclusions can be drawn.
Since the shell diameter is between 254mm and 508mm, the recommended
maximum number of tube passes is 6 (Zhang, 2013).
The shell-diameter-to –tube-length ratio needs to be between 1/5 and 1/15
(Zhang, 2013). Since 0.0667<0.0696<0.2, the ratio is just within the
acceptable range.
The number of tubes on the tube side can be confirmed using the shell side
geometry. The relationship for the number of tubes is given as
Where A1 is given by
And Pt is
Therefore,
And thus,
The number of tubes calculated by this method is 0.1% more than previously
calculated. Since this shows that there is as insignificant change in the
required number of tubes, 124 is accepted as the amount of tubes.
BAFFLES
The optimal baffle spacing is estimated to be in the range of 40% to 60% of
the shell diameter. Using the previously calculated shell diameter, 0.4247m,
the calculated range of the baffle spacing is between 0.1699m and 0.2548m.
When this value was input into HYSYS, the lower bound was used as this
value will maximise the heat exchanger’s heat transfer. It is recommended to
use a baffle cut of 25-35% (Zhang, 2013).
61 | P a g e
HEATER AND COOLER SIZING
As part of the economic analysis, the heaters and coolers needed to be sized,
as the price per unit is a factor of the surface area. The procedure for
determining the size was very similar to the process used for the heat
exchanger, with a few alterations. The temperature difference is now equal to
the temperature out minus temperature in. This is because each cooler and
heater only has one input and one output stream. The equation for the area
thus becomes
̇
U, the heat transfer coefficient, has been calculated by taking the average of
the values provided in the design brief (Zhang, 2013).
For example, the heat transfer coefficient for heater E-100 was calculated as
follows:
h(boiling organic)= 1000 W/m2.0C
h(flowing liquid)= 600 W/m2.0C
h(flowing gas)= 60 W/m2.0C
Therefore,
The following table presents a summary of the final sizes of the heaters and
coolers for the impure case.
Table 12: Summary of the sizes of the heaters and coolers for impure
feed base case
The sizes for the heat exchanger, coolers and heaters have also been
calculated for the pure feed case. All of the sizes are slightly different.
62 | P a g e
5.2.3 Distillation column sizing
Sizing the distillation columns required an iterative process to determine the
condenser and reboiler pressure, reflux ratio and number of trays. A splitter,
shortcut column, flash separator and spread sheets were used to determine
distillation column dependent variables and diameter and height calculations.
BENZENE COLUMN
The calculations for sizing the Benzene distillation column begun with initially
assuming the reboiler pressure is the same as the condenser pressure. This
gave an initial iteration reboiler pressure guess of 1 bar. The reflux ratio,
actual number of trays, feed stage, column diameter and height were then
calculated using shortcut column, flash column and HYSYS spread sheets for
pressure drop calculations. From these values a new reboiler pressure could
be calculated. The results were then input into the Benzene distillation column
and the calculations repeated by HYSYS until a converged reboiler pressure
of 1.725 bar was reached. The method used is described below using the final
iteration step which resulted in the converged reboiler pressure and column
diameter and height of 1.707m and 23.77m respectively. The column height to
diameter ratio is 13.92 which agreed with heuristic data.
Column diameter
A spread sheet in HYSYS was used to calculate the column diameter. First
the liquid and vapour mass flow, surface tension, mass density, molecular
weight were input from the HYSYS model. The estimated tray space was 24
inch and a CSB of 0.4 was used as FLG was determined to be less than 0.01. It
was important to check the proposed tray spacing after the column diameter
was calculated as 24 inch tray spacing is only applicable if the column
diameter is greater than 4 feet.
63 | P a g e
Then the calculations were performed using the spread sheet (Figure 82:
Spread sheet - 1D) and equations as defined in Appendix K – Column Sizing
Calculations.
According to the spread sheet (Figure 82: Spread sheet - 1D), the column
diameter was calculated as:
[ ]
( )
The number of tray passing was first determined using the Multipass tray
diagram (Figure 40: Multipass Tray Diagram).
According to Spreadsheet-1D:
Dt (ft) =1.72m=5.65ft
Liquid Flow Rate=57.96 kg/hr=57.96*264.172/(8120*60)=0.0314(gal/min)
From the Multipass tray diagram above, the intersection is close to single-
pass curve so it was recommended that the column used the single-pass
method.
64 | P a g e
The detail calculations are provided in Appendix K – Column Sizing
Calculations and the results are shown in the HYSYS spread sheet below.
A third spread sheet (Figure 84: Spread sheet - 1H) was created in HYSYS
and used to calculate the Benzene distillation height and ratio of column
height to diameter. The calculations and equations used are described below.
Temperature, viscosity and molar flow of each component into the column
were imported from the HYSYS model into the spread sheet. These values
were then used in the calculations.
CUMENE COLUMN
The condenser and reboiler pressure, column height and diameter, actual
number of trays and reflux ratio of the Cumene distillation column were
calculated using the same method as the Benzene column. This included the
use of HYSYS and splitter, shortcut column, flash column and numerous
spread sheets which involved many calculations.
65 | P a g e
6 Case Study Summaries
6.1 Base Case: Pure Benzene and Propylene with 5wt% Propane
Impurity Feed
Pure Benzene and Propylene with 5wt% Propane feeds were selected as the
base case because it was assumed that the impure feed was more readily
available. It was also predicted in terms of economics the impure Propylene
feed would be more viable as its purchase cost was less than the pure option,
as specified by the client. The impure feed also required three distillation
columns in the process to separate the Propylene and Propane as fuel gas
from the unreacted Benzene and produced Cumene and DIPB; this lead to
the conclusion that using the impure feed as the base case would simplify the
case study calculations. This simplification when considering the case study
involving the pure Propylene feed would arise from replacement of the
distillation column with a splitter due to the Propane impurity no longer being
present.
The base case was setup in HYSYS as represented by the process flow
diagram in Figure 6: Process Flow Diagram (PFD). The stream temperatures
were initially determined from literature studies but were changed during the
HYSYS simulation to meet the process conditions. Similarly, the final
pressures of the process streams were determined from HYSYS simulation
based initially on researched values.
Flow rates of each stream were determined such that overall process
produced 100,000 metric ton of Cumene per year. Initial material balances
were performed with numerous assumptions made to simplify the
calculations; the results of these balances vary from the final simulated results
due to these assumptions made. A ‘recycle’ function was used in HYSYS
simulation with the convergence method deemed most appropriate. This
function resulted in a ‘tear’ of the Benzene recycle stream which allowed
HYSYS to iteratively calculate the compositions and flow rates of each
process stream.
66 | P a g e
The vapour fraction of each stream was also determined by HYSYS from the
stream temperature, pressure and composition, in combination with the
knowledge of the boiling points of each component at the defined conditions.
67 | P a g e
6.2 Case Study One: Pure Benzene and Pure Propylene Feed
The case study involved replacing the impure Propylene feed with pure
Propylene and comparing the economics associated with the two feed options
to determine the most economical option. As a result of this change the
depropaniser was substituted by a splitter and the reactor size had to be
recalculated as the base case dimensions were no longer applicable. The
remaining unit operations in the Cumene production plant remained constant
between the base case and case study to investigate the effect of the purity of
the Propylene feed.
SPLITTER
As there is no Propane impurity present in the process the depropaniser
distillation column was replaced with a splitter for simplification purposes.
REACTOR SIZING
The maximum possible conversion for the reaction was determined using the
Gibbs reactor to be 0.9941. This conversion was required when sizing the
reactor for this case study.
68 | P a g e
Due to the change in feed composition, which resulted in the removal of the
inert Propane component, it was expected that the reactor size would change.
The optimum reactor size was investigated in HYSYS by creating a
‘DataBook’ case study. The conversion of the reaction associated with
volumes analysed by HYSYS was calculated and compared to the maximum
theoretical value determined from the Gibbs reactor. The reactor size was first
investigated in terms of its volume because it was simpler to observe the
relationship between reactor volume and corresponding reaction conversion.
In the HYSYS case study the reactor volume was set as the independent
variable, while the percentage conversion was specified as the dependent
variable. The method and analysis of this case study is discussed below.
Figure 44: HYSYS DataBook case study setup for reactor volume and
conevrsion
69 | P a g e
The data obtained from the HYSYS calculations was then transferred to Excel
and the plot of the relationship between reactor volume and reaction
conversion was created (Figure 46: Reactor volume vs conversion) so further
analysis could occur. From the graph it was seen that the conversion
increases with increasing reactor volume. It was also observed that a reactor
volume of 37m³ resulted in the highest conversion of 0.9940 without
exceeding the Gibbs calculated conversion.
The reactor diameter and length were then estimated from the determined
optimal reactor volume of 37m³. The possible range of ratio of length to
diameter (10 to 20) restricted the feasible combinations of length and
diameter. It was assessed by analysing HYSYS case study simulation that the
reactor diameter and length was 1.6m and 18.4m, respectively.
70 | P a g e
Figure 47: Defining variables bounds in HYSYS reactor sizing case
study
Another case study was performed in HYSYS, this time however the diameter
and length of the reactor were set as the independent variables and the
conversion was again specified as the dependent variable. A low and high
bound on the size variables was implemented during the case study.
Figure 48: HYSYS DataBook setup for reactor diameter and length case
study
It was determined that reactor diameter and length of 1.6m and 18.4m
respectively resulted in the maximum conversion of 0.9934 without exceeding
the Gibbs conversion. This was found by observing the feasible combinations
of length and diameter, which produced the optimal reactor volume and the
conversions they resulted in. These results were viewed in the ‘Results’
section of the HYSYS case study DataBook. The reactor size was determined
to be the length and diameter which produced the conversion closet to the
calculate Gibbs reactor value without exceeding it.
71 | P a g e
Figure 49: HYSYS case study results; reactor diameter and length with
associated conversion
72 | P a g e
Figure 51: PFR setup in HYSYS; defining calculated size
The optimal reactor sizing determined was then used to define the size of the
Plug flow reactor in the Cumene production process. From this the actual
conversion of the main and side reaction were found to be 95.2% and 4.053%
respectively.
Figure 52: Main and side reaction conversion for pure Propylene feed
73 | P a g e
CUMENE PRODUCTION PROCESS
The production of the Cumene with the pure Propylene feed was 12, 610
kg/hr compared to the impure Propylene feed which resulted in Cumene
production of 12,592 kg/hr. The amount of Cumene produced exceeded the
client’s requirements of the plant so in terms of production the pure Propylene
feed would be better. However, after the economic analysis was performed it
was found that the pure Propylene feed was not economical viable and the
best option was the base case of impure Propylene feed.
Figure 53: Mass flow of Cumene produced with pure Propylene feed
74 | P a g e
6.3 Case Study Two: Reactor Performance – Temperature
The effect of temperature on reaction conversion and overall process
performance on an isobar (25 bar) reactor was investigated for adiabatic and
isothermal conditions. An optimal reactor temperature of 365oC and an inlet
temperature ranging between 275oC and 362 oC was obtained from the case
study.
DataBook in HYSYS was used to set up a case study to investigate the effect
of feed temperature on conversion. The independent variable of the case
study was set as feed temperature and the dependent variables as the
conversions of each reaction. In Figure 55: HYSYS DataBook case study
setup (below) “Feed 25bar” refers to the reactor input stream, “Act.%Cnv.
(Act. %Cnv._1)” represents the conversion of the main reaction and similarly
“Act.%Cnv. (Act. %Cnv._2)”represents the conversion of the side reaction.
75 | P a g e
Figure 55: HYSYS DataBook case study setup
Adiabatic reactor
An adiabatic reactor was used in the Base Case and Case Study 1 and has
no energy stream connection. Heat is generated in the reactor by the
exothermic reaction, causing an increase in reactor temperature. As a result
the outlet stream temperature is greater than the inlet stream temperature.
76 | P a g e
Figure 57: Adiabatic reactor relationship of inlet and outlet temperatures
A case study for the effect of inlet temperatures from 260 oC to 380oC on
conversion was investigated by plotting the main and side reaction
conversions against inlet temperature. In the case study a step size of 5oC
was used. Figure 58: HYSYS case study setup shows the HYSYS Case
Study setup.
It was observed from Figure 60: Relationship between main reaction
conversion and inlet temperature and Figure 61: Relationship between side
reaction conversion and inlet temperature that the main desired reaction
reached a maximum conversion of 97.80% with an inlet temperature of 275
o
C. While the relationship between inlet temperature and side reaction
conversion does not reach a maximum value and can be approximated by a
linear relation of as demonstrated by
Figure 61: Relationship between side reaction conversion and inlet
temperature. As the conversion of the side reaction increases with increasing
inlet temperature, a low inlet temperature is desired. At 275 oC the side
reaction has a relatively low conversion of 2.13% which is considered
acceptable.
77 | P a g e
Figure 58: HYSYS case study setup
78 | P a g e
Main Conversion
100
95
%Conversion
90
85
80
75
70
260 270 280 290 300 310 320 330 340 350 360 370 380
Temperature
8
%Conversion
0
260 270 280 290 300 310 320 330 340 350 360 370 380
Temperature
Isothermal reactor
An isothermal reactor was also used to investigate the effect of reactor
temperature on reactor performance. As the reaction is exothermic a cooling
stream (represented by ‘CS’ in Figure 62: HYSYS PFR reactor with introduces
'SET' stream) was introduced to remove heat generated by the reaction and
maintain isothermal reactor conditions. A ‘SET’ function was also used to
ensure the reactor inlet and outlet temperatures remained equal.
79 | P a g e
Figure 62: HYSYS PFR reactor with introduces 'SET' stream
The same method used in the adiabatic reactor section was used to observe
the effect of reactor temperature on reaction conversions. It was decided that
the temperature range of 340oC to 400 oC was of interest in the case study
investigation with a step size of 1oC.
80 | P a g e
conversion. Similarly to the adiabatic reactor case, the main reaction reached
a maximum conversion and the side reaction conversion can be approximated
by a linear relationship. The maximum conversion of 99.12% occurred at
365oC for the main reaction which is a higher conversion and temperature
compared to the adiabatic reactor investigation. The side reaction conversion
is closer to a linear approximation than that obserseved in the adiabatic case
and the conversion of the side reaction increases as reactor temperature
increases. At the optimal isothermal reactor temperature of 365oC the side
reaction conversion is 1.74%; which is less than the adiabatic reactor side
reaction conversion at the maximum desired conversion. It was concluded
that an isothermal reactor optimal temperature would be 365oC
Main conversion
98.5
98
97.5
%Conversion
97
96.5
96
95.5
95
94.5
340 345 350 355 360 365 370 375 380 385 390 395 400
Temperature
81 | P a g e
Side conversion y = 0.0265x - 7.8955
R² = 0.9943
3
2.5
%Conversion
2
1.5
1
0.5
0
340 345 350 355 360 365 370 375 380 385 390 395 400
Temperture
Group 4 82 | P a g e
7 Economical Evaluation
All of the costs given in the design brief (Zhang, 2013) have been given as the
values in 2011. To increase the accuracy of the cost estimations, the chemical
engineering plant cost index (CEPCI) has been used to index the prices. In
2011, the CEPCI was 585.7 (Chemical Engineering Magazine, 2011), and in
May 2013, the CEPCI was 566.5 (Chemical Engineering Magazine, Sept
2013). Converting the prices to 2013 prices was done using the following
formula
The following sections of the report give details in the calculations involved
and the outcomes.
Group 4 83 | P a g e
Table 14: Summary of the equipment costs for the pure feed material
Group 4 84 | P a g e
Table 15: Summary of the equipment costs for the impure feed material
It can be observed from the tables above that the capital cost for the impure
feed is more than the cost for the pure case. This is because of the previously
discussed additional equipment required for the extra separation. It can also
be noted that most of the equipment is very similarly priced from the pure feed
to the impure feed. This is because in general, the flow rates, temperatures
and pressures were similar for both cases.
One of the main costs is for the storage tanks. Peters and Timmerhaus (1991)
suggest that storage tanks have the capacity for one month’s worth of feed
material/product. The pressure of the tanks is a factor of the cost; most of the
Group 4 85 | P a g e
tanks are above 1 bar, and so incur additional cost. If Dr Who Chemicals Ltd
are concerned about the cost of these tanks, they could explore only storing
enough chemicals for a fortnight or a week, although this is likely to incur
additional costs for increased transportation.
Carbon Steel
Carbon steel has many advantages. It is easily formed into new shapes, and
is available in a wide range of sizes. It is widely used for organic chemicals
(Peters and Timmehaus, 1991), and all of the chemicals used in the system
are organic. Compared to stainless steel, it is much cheaper. However, there
are a few disadvantages of carbon steel. It is not resistant to corrosion, and it
is only suitable up to temperatures up to 480⁰C.
Stainless Steel
Stainless steel is often used for corrosive chemicals, and has a much higher
maximum temperature of between 600-1160⁰C, depending on the particular
alloy of stainless steel. Stainless steel has a minimum chromium content of
12%, and increasing the quantity of chromium increases the resistance to
oxidising agents. The main disadvantage of stainless steel is the cost. It is
much more expensive than carbon steel.
7.2 Utilities
As part of the economic evaluation for the plant, the annual utility costs have
been calculated. The type of utilities available, as well as the price were given
in the design brief (Zhang, 2013). The main utilities required were power for
the pump, cooling water for coolers and condensers, and steam for heaters
and reboilers.
There were three types of steam available for use; low pressure, medium
pressure and high pressure. The low pressure steam was available saturated
at 446 kPa. From steam tables, this corresponds with a maximum
temperature of 147.58⁰C. Medium pressure (saturated) steam at 1135 kPa
results in a maximum temperature of 185.47⁰C. The high pressure (saturated)
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steam at 4237 kPa results in a maximum temperature of 253.765⁰C. A lower
pressured steam corresponded with a lower cost for the steam per 1000 kg
required. Thus, to keep utility costs at a minimum, the lowest steam pressure
possible was used for heat transfer.
Similarly, for cooling there were two different utilities available- cooling and
refrigerated water. Refrigerated water costed ten times more than the cooling
water, so the quantity of refrigerated water required was kept to a minimum.
Disposal of the di-isopropyl benzene produced in the side reaction was also
required, and incurred a cost.
A summary of the annual costs of the utilities for the pure and impure feeds
can be found in Table 16: Annual utility costs for the impure feed plant and
Table 17: Annual utility costs for the pure feed plant.
Table 16: Annual utility costs for the impure feed plant
Item Quantity per Price per unit Cost (2011) Cost (2013)
year
Power 2.47E+05 kwh 0.061/ kWh $ 15,087.33 $ 14,592.75
LP Steam 1.68E+08 kg $3/1000 kg $ 50,436.00 $ 48,782.64
MP Steam 1.56E+07 kg $6.50/1000 kg $ 101,296.00 $ 97,975.39
HP Steam 1.86E+08 kg $8/1000 kg $ 1,486,592.00 $ 1,437,859.60
Cooling Water 2.12E+10 kg $20/1000 m3 $ 425,199.23 $ 411,260.65
Refrigerated Water 1.17E+10 kg $200/1000 m3 $ 23,938.40 $ 23,153.67
DIPB disposal 5.36E+06 kg $1/kg $ 5,362,400.00 $ 5,186,613.62
Total - - $ 7,464,948.97 $ 7,220,238.32
Table 17: Annual utility costs for the pure feed plant
Item Quantity per Price per unit Cost (2011) Cost (2013)
year
Power 2.36E+05 kwh 0.061/ kWh $14,368.89 $13,897.86
LP Steam 4.23E+07 kg $3/1000 kg $ 126,799.20 $ 122,642.56
MP Steam 1.55E+07 kg $6.50/1000 kg $100,516.00 $ 97,220.96
HP Steam 1.42E+08 kg $8/1000 kg $ 1,133,248.00 $ 1,096,098.67
Cooling Water 2.03E+10 kg $20/1000 m3 $ 407,663.88 $ 394,300.13
Refrigerated Water 1.20E+09 kg $200/1000 m3 $ 23,922.40 $ 23,138.19
DIPB disposal 6.01E+06 kg $1/kg $6,012,000.00 $5,814,918.90
Total - - $ 7,818,518.37 $ 7,562,217.27
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Overall, it can be observed from tables Table 16: Annual utility costs for the
impure feed plant and Table 17: Annual utility costs for the pure feed plant
that the plant with a pure feed has a higher annual utility cost. This is due to
the cost of disposal of di-isopropyl benzene. The pure feed plant produces
more DIPB than the impure feed, and each addition kilogram of DIPB costs
one dollar to dispose. The cost of disposal is the most expensive utility.
It can be also be observed the plant with an impure feed has a higher annual
utility cost for steam and water. This is because this plant requires an extra
column for separating the fuel gas from the heavier components. This column
requires additional steam and cooling water for the reboiler and condenser. In
the pure feed, there is no fuel gas produced, and the propylene is converted
fully to Cumene. A component splitter is necessary; however, this requires no
additional utilities.
Based on the equipment, it was determined that five workers were required
per shift. These employees needed to monitor and operate the heat
exchangers and reactor in particular, and it is assumed that they would also
manage the vessels, tanks and pumps.
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7.4 Raw materials and products
The raw materials and products have a significant impact on the economic
feasibility of the Cumene plant.
Pure Case
In the pure feed case, the raw materials are Benzene (8564 kg/h) and pure
Propylene (4808 kg/h). To get the quantity required for a year, the mass flow
rates were multiplied by 8000 (the assumed hours of operation per year). This
resulted in the Benzene costing $76,733,440.00 and Propylene costing
$60,388,480.00 per year. The catalyst was assumed to be replaced annually.
This has been assumed because the reactor temperature is above the optimal
temperature for the catalyst, and so it might have been thermally degraded,
and also because some catalysts for this reaction are not regenerable. The
quantity of catalyst required was calculated as the product of the catalyst’s
density, void fraction and the reactor volume. Thus, 29944 kg of catalyst was
required, at an annual cost of $75,458.88. The total cost of raw materials was
determined to be $137,197,378.88 per annum.
For the pure feed, the only desired product was Cumene. 12640 kg/h of
Cumene will be produced annually, with revenue of $145,006,080.00 p.a.
DIPB is also produced at a rate of 751.5 kg/h, however, this is then treated
and disposed of. The cost of the disposal has been included in the utility cost.
There is overall product revenue of $7,808,701.12 per year for the pure case.
Impure Case
The raw materials for the impure case are benzene (8520 kg/h) and
Propylene with a 5% Propane impurity (5043 kg/h). The process for
calculating the annual cost was the same as for the pure case. This flow rates
result in annual costs of $76,339,200.00 for Benzene and $35,502,720.00 for
the impure Propylene. Once again the catalyst was assumed to be replaced
annually, at cost of $86,688.00. This cost is slightly higher than the pure case,
as the volume of the reactor has increased. The total cost for the raw
materials was calculated to be $111,928,608.00. This is significantly less than
the raw material cost for the pure case, as the Propylene feed is almost half of
the pure Propylene cost.
For the impure feed, there are two products that can be sold. The main
product is Cumene, with a mass flow rate of 12610 kg/h, corresponding to
revenue of $144,661,920.00 per year. The unreacted Propane and Propylene
can be sold as a fuel gas. 280.6 kg/h of fuel gas is produced, with annual
revenue of $1,414,224.00. DIPB is also produced, and the cost of disposal
has been included in the utilities cost.
The overall product revenue per annum is $34,147,536.00 for the impure
case.
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7.5 Economic Appraisal
The following section gives insight into the economic feasibility of constructing
and operating a new cumene plant. All calculations for this section can be
found in Appendix M – Economic Evaluation.
The depreciation schedule was linear, with 10% annually (and no residual
value). The internal hurdle rate was 9% (Zhang, 2013), and the construction
period was 1 year). A marginal tax rate of 35% was used. In the case of the
pure feed, it was assumed that there was no tax being paid as the plant was
losing money each year. This assumes the Dr Who Chemicals Ltd has no
other assets that are generating profits. If the client has other assets that are
returning a profit, Dr Who Chemicals Ltd will have to pay tax on this plant, as
the marginal tax rate is applied to the overall profits of a company, rather than
on individual assets.
Due to the negative magnitude of the NPV for the pure case, no DCF ROR
calculation was possible.
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Economic Evaluation. It can be estimated that the payback period is between
one and two years.
For a more accurate calculation, a graph of the relationship between the time
and the cumulative cash flows has been plotted.
$35,000,000.00
y = 3.1876E+06x - 5.745E+06
$30,000,000.00
$25,000,000.00
Cumulative Cash Flow
$20,000,000.00
$15,000,000.00
$10,000,000.00
$5,000,000.00
$-
0 2 4 6 8 10 12 14
$(5,000,000.00)
$(10,000,000.00)
Years
Figure 67: Payback period for the impure feed Cumene plant
Using the equation Figure 67: Payback period for the impure feed Cumene
plant:
There is no payback period for the pure feed, as the plant will never make
money, and hence, never break even.
Using the above formula, the ROI for the impure case is 205%.
For the pure case, the ROI is -2493%, an unsurprising result given the
magnitude of money that the plant loses each year.
Group 4 91 | P a g e
7.6 Economic conclusions
The net present value, discounted cash flow rate of return, payback period
and return on investment have been calculated for both the pure and impure
cases have been calculated. Each of these economic measures concludes
that the impure feed plant is an economically feasible project. Conversely, the
economic indicators all determine that the pure feed plant is not economically
viable, and the plant will lose a significant amount of money during its life.
Group 4 92 | P a g e
8 Recommendations
It was determined that the reactor, in which the alkylation of Benzene and
Propylene to Cumene occurs with the presence of a catalyst, be a packed bed
reactor of volume of 43m3. This reactor size was calculated to result in a
conversion which was considered to be closest to the ideal maximum. The
Cumene production process proposed has three distillation columns to
separate the unreacted Propylene and Benzene, inert Propane and undesired
DIPB product from the Cumene which is produced with 99.94% purity,
meeting the client’s requirements.
Group 4 93 | P a g e
9 References
Buelna G and Nenoff, TM, 2005, ‘A one-step catalytic separation process for
the production of cumene’, Catalysis Letters, vol 102, no. 3-4 pp
Dai, C, Lei, Z, Zhang, J, Li, Y, Chen B, 2013, ‘Monolith catalysts for the
alkylation of benzene with propylene’, Chemical Engineering Science, 100,
342-251
Zhang, 2013, Design Project Specification, Chem Eng 3030: Simulation and
conceptual design and Chem Eng 3025: Pharmaceutical plant design and
process engineering
Group 4 94 | P a g e
Luyben, W, 2010, Design and Control on the Cumene Process, Ind. Eng.
Chem. Res. 49, 719-734
Matheson Tri Gas, 2008, Material Safety Data Sheet; Cumene, Available from
http://www.mathesongas.com/pdfs/msds/MAT05570.pdf, Date accessed: 2
September 2013
Matheson Tri Gas, 2008, Material Safety Data Sheet; Propylene, Available
from http://www.mathesongas.com/pdfs/msds/MAT19830.pdf, Date accessed:
2 September 2013
Group 4 95 | P a g e
Sigma Aldrich Co LLC, 2013. MSDS Searchand Product Safety Center.
Available from http://www.sigmaaldrich.com/safety-center.html. Date
accessed 15 August 2013
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Appendices
Appendix A – Plant Location Justification
There are numerous factors to consider when determining a potential location
for the Cumene production plant.
Benzene
Produced at some of BHP’s coking plants, as BTX, however would
need to be separated from other components prior to use therefore not
feasible
Huntsman Chemical Co., a chemical manufacturing plant (shut down in
2005), imported 80% of benzene when it was operational, which
suggests that Benzene cannot be supplied from within Australia so will
need to be imported.
Producers of benzene close to Australia
a. Shell
China (CSPCL Nanhai)
Japan (Kawasaki, Yokkaichi, Yamaguchi)
Singapore (Jurong Island, Palau Bukom)
All of the above also produce propylene if the domestic
supply is deemed incapable (Shell Australia, n/a)
2. Storage of raw materials and Cumene product will be in a high
pressure vessel on site.
3. Transportation facilities: Plant should be located close to at least two
major forms of transportation including road, railway, waterway or
ocean for international shipping.
4. Weather/Climate: Extreme weather can increase plant facilities cost
due to the requirement of additional insulation or heating. Ideally plant
location is in and area which is not exposed to regular extreme weather
(heat/cold/humidity).
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5. Safety and Environmental measures: Minimal risk of natural events
including flood, hurricane or fire.
6. Water supply: Availability of large quantities of water which can be
used for steam generation, cooling of the plant, washing. Viable water
sources include large lakes, rivers or artesian wells.
7. Local community considerations: The plant needs to be accepted by
the local community and not impose any health or safety risks on the
surrounding community. There will be a need for satisfactory living
conditions for plant personnel and regional culture needs to be
considered and positively affected by the process plant. New
investments are also preferred by communities to be located in areas
with high unemployment.
8. Client location: Ease of transport and communication between the
company head office and process plant.
Based on the location of raw materials, it was proposed that the Cumene
production plant be located in Victoria, west of Melbourne. This location is
within close proximity to local raw material supply and is located near the
coast so Benzene can be readily shipped from Asian Shell supplies. There is
a river and extensive railway yard located nearby for easy transportation of
raw materials and Cumene product nationally. Extreme weather does not
regularly affect this area (Bureau of Meteorology, 1996) and water required for
process operation can be sourced from numerous nearby rivers and coastline.
There are already a lot of manufacturing and construction sites in regional
Victoria and an industrial area west of Melbourne so concern from the
community would be minimal. Victoria also have an unemployment rate of
5.5% in June 2013 (up from national average of 5.1%) (Australian Bureau of
Statistics, 2012) which could potentially be decreased by the prospect of a
new plant as employment opportunities would be created in the construction,
operating and maintenance of the Cumene production process plant. There is
existing railway infrastructure to transport the product from production site to
client located in Adelaide. Road and air transportation is also readily available
for personnel transfer between Adelaide office and west of Melbourne plant.
Numerous factors were considered in the decision of potential plant location
and it was proposed that a site west of Melbourne, Victoria would be most
appropriate and economically beneficial.
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Appendix B – Means End Analysis
Block Flow Diagram (BFD):
Benzene Recycle
Benzene
Feed Fuel Gas
REACTOR 1 (Propylene & Propane)
Reaction 1: Cumene
Propylene production SEPARATOR
Feed Reaction 2: side
reaction, undesired
DISTILLATION
COLUMN 2
Cumene
VACCUM
DISTILLATION
COLUMN
PDIB
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ALKYLATION OF BENZENE TO CUMENE
SYNTHESIS STEP
1. Eliminate differences in molecular types:
-Chemical reactions of process
Desired Reaction – Alkylation of Benzene
-Benzene reacts with Propylene to produce the Cumene with the presence of a
catalyst
-exothermic, kinetic reaction
- catalyst: -particle diameter=1.5mm
-particle density= 1600 kg/m^3
-void fraction =0.5
-h=75W/m^2C (heat coefficient)
-An undesired product DIPB is formed by the reaction of Cumene and Propylene
2. Distribution of chemicals:
-Sources [Raw materials feed]:
-Benzene >99.9% purity
-Propylene >99.9wt% purity or Propylene with 5wt% propane impurity
-Sinks [Products]:
-Cumene (desired)
-Di-isopropyl benzene (undesired side reaction) Benzene
(unreacted)
RECYCLE
Benzene
In excess,
recycle Cumene (product)
25°C, 1bar 100,000ton/year
104.17kmol/hr
Reactor
Propylene
95%wt 358°C, 25bar
25°C, 1bar DIPB
(p-diisopropyl benzene)
Exothermic Reaction (product)
Propane
5%wt impurity
25°C, 1bar Propylene (unreacted)
Propane (unreacted)
Group 4 100 | P a g e
Table 19: Basic physical properties of components
Benzene Cumene
Overall Separation
Propylene Separate Cumene from
Propane unreacted raw material feed Benzene
Cumene and undesired products Propylene
DIPB Propane
DIPB
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4. Eliminate differences in temperature, pressure and phase:
-pumps used to increase pressure
-heater exchangers used to change temperature Fuel Gas
-valves used to decrease pressure
increase increase
Boiling Points at 25bar
Propylene Propane Benzene Cumene DIPB
24⁰C 34⁰C 237⁰C 306⁰C 419⁰C
Distillation column conditions need to be changed in order to separate Boiling Points at 5bar
Propane and Propylene as Fuel Gas mixture
-17⁰C -10⁰C 144⁰C 217⁰C 296⁰C
Fuel Gas
100°C, 5bar Benzene Recycle
Group 4 102 | P a g e
Benzene
Fuel Gas Recycle Cumene
100°C, 5bar 150°C, 4bar 150°C, 4bar
100⁰ 100°C Pressure 100°C Temperature 150⁰ C Dist Col 2 Dist Col 3
Dist Col 1 150°C DIPB
C 5bar 5bar Change 4bar Change 4bar Benzene Cumene
Depropaniser 4bar 150°C, 4bar
decrease decrease
5. Integration
-The process will be continuous
-Recycle the unreacted excess Benzene to reduce the required feed and associated cost
-Purge propylene and propane mixture for fuel gas
-REQUIREMENT: Cumene Production 100,000 metric ton/year
-Base case
-Economic analysis to compare Propylene feed with 5% impurity and Propylene feed >99.9%
-Recycle/Purge ratio should be determined for propylene, benzene and DIPB
Group 4 103 | P a g e
Appendix C - Key Chemical Information and Associated
Hazards
C1. Propylene
General Properties
Propylene (C3H6) is an unsaturated hydrocarbon. Propylene (at atmospheric
pressure) has a melting point of -185.2⁰C, a flash point of -72⁰C, a boiling
point of -47.4⁰C and an auto-ignition point of 499⁰C. Propylene is immiscible
in water. Propylene is considered stable.
Group 4
Table 21: Potential Health effects and first aid measures associated with
Propylene
Safety
Table 22: Hazard Ratings for Propylene
Flammability 4
Toxicity 2
Body Contact 2
Reactivity 2
Chronic 2
Group 4
propylene is inhaled, patient’s breathing and pulse must be monitored
continuously. If propylene is ingested, milk, oil and alcohol should not be
administered.
Firefighting Measures
Propylene will easily ignite if it comes into contact with heat or flames. A fire
should not be extinguished unless the leak can be stopped; if it cannot be put
out, the gas should be left to burn. If the fire is small, dry chemical, carbon
dioxide or water can be used. Contamination with oxidising agents should be
avoided. Firefighters must be called.
Transport
Propylene is in transport hazard class 2.1. Hazard class 2.1 is for flammable
gases which ignite when they come into contact with an ignition source.
Group 4
C2. Propane
General Properties
Propane (C3H8) is a hydrocarbon. Propane (at atmospheric pressure) has a
melting point of -189.7⁰C, a flash point of -104.44⁰C, a boiling point of -42.1⁰C
and an auto-ignition point of 468⁰C. Propane is slightly soluble in water.
Propane is considered stable.
Group 4
Table 24: Potential health effects and first aid measures associated with
Propane
Safety
Table 25: Hazard ratings for Propane
Flammability 4
Toxicity 2
Body Contact 2
Reactivity 2
Chronic 0
Group 4
may also occur, and if this occurs, the patient should be moved into a warm
area, and affected body parts should be bathed in lukewarm water. Medical
attention should be sought. If propane is inhaled, patient’s breathing and
pulse must be monitored continuously.
Firefighting Measures
A fire should not be extinguished unless the leak can be stopped; if it cannot
be put out, the gas should be left to burn. If the fire is small, dry chemical,
carbon dioxide or water can be used. Contamination with oxidising agents
should be avoided. Firefighters must be called.
Transport
Propane is in transport hazard class 2.1. Hazard class 2.1 is for flammable
gases which ignite when they come into contact with an ignition source.
Group 4
C3. Benzene
General Properties
Benzene (C6H6) is an unsaturated hydrocarbon, with the carbons joined in a
ring shape. Benzene has a characteristic odour. Benzene (at atmospheric
pressure) has a melting point of 5.5⁰C, a flash point of -11⁰C, a boiling point of
80⁰C and an auto-ignition point of 562⁰C. It evaporates quickly and is
immiscible in water. Benzene is considered stable.
Group 4
Table 27: Potential health effects and first aid measures associated with
Benzene
Group 4
Safety
Table 28: Hazard ratings for Benzene
Flammability 3
Toxicity 2
Body Contact 2
Reactivity 1
Chronic 4
The maximum rating is 4.
Benzene has an overall rating for hazards of extreme. Benzene can cause
serious health damage if exposure is prolonged, or if it is inhaled. It is irritating
to both the skin and eyes and can cause cancer and birth defects. Benzene
is also highly flammable. Appropriate clothing, including glasses, closed
footwear should be worn when handling benzene.
Firefighting Measures
If a fire breaks out, foam, BCF, carbon dioxide and dry chemical powder can
be used to extinguish it. Oxidising agents must be avoided. Both liquid and
gas phase benzene are highly explosive, and should be kept away from heat.
Transport
Benzene is in transport hazard class 3 and packing group II. Class 3 is for
flammable liquids and packing group II involves chemicals with a boiling point
above 35⁰C and a flash point less than 23⁰C.
Group 4
Figure 73: Benzene phase diagram from HYSYS
Group 4
C4. Cumene
General Properties
Cumene (C9H12) is an unsaturated hydrocarbon. Cumene has a sharp odour.
Cumene (at atmospheric pressure) has a melting point of 96⁰C, a flash point
of 44⁰C, a boiling point of 152.4⁰C and an auto-ignition point of 422⁰C. It
evaporates slowly and is immiscible in water. Cumene is considered stable.
Group 4
Table 30: Potential health effects and first aid measures associated with
Cumene
Safety
Table 31: Hazard ratings for Cumene
Flammability 2
Toxicity 2
Body Contact 2
Reactivity 2
Chronic 2
Group 4
First Aid Measures
If cumene comes into contract with eyes, it must be washed out instantly, and
totally irrigated. Medical attention should be sought. If skin contact arises,
contaminated clothing should be discarded and skin should be flushed with
water. If cumene is inhaled, the patient should be taken to hospital
immediately. Never give a patient milk, water or alcohol, and seek medical
attention.
Firefighting Measures
If a fire breaks out, foam, BCF, carbon dioxide and dry chemical powder can
be used to extinguish it. Oxidising agents must be avoided. Cumene vapour
reacts explosively with air. Firefighters should be called.
Transport
Benzene is in transport hazard class 3 and packing group III. Class 3 is for
flammable liquids and packing group II involves chemicals which do not fit the
characteristics for packing groups I or II.
Group 4
C5. Di-Isopropyl Benzene
General Properties
Di-isopropyl benzene (C12H18) is an unsaturated hydrocarbon. Di-isopropyl
benzene (at atmospheric pressure) has a melting point of -63⁰C, a flash point
of 91.11⁰C, a boiling point of 203⁰C and an auto-ignition point of 448.89⁰C. Di-
isopropyl benzene is immiscible in water. Di-isopropyl benzene is considered
stable.
Group 4
Table 33: Potential health effects and first aid measures associated with
Di-isopropyl Benzene
Safety
Table 34: Hazard ratings for Di-isopropyl Benzene
Flammability 1
Toxicity 2
Body Contact 2
Reactivity 1
Chronic 2
Group 4
Firefighting Measures
If a fire breaks out, foam, BCF, carbon dioxide and dry chemical powder can
be used to extinguish it. Contamination with oxidising agents should be
avoided. Di-isopropyl benzene is combustible and if combustion occurs,
carbon monoxide may be released. Firefighters should be called.
Transport
Di-isopropyl benzene is in transport hazard class 9 and packing group III.
Hazard class 9 is for miscellaneous chemicals.
Group 4
Appendix D - Environmental Impacts
D1. Propylene
Propylene is a main feedstock chemical for the reaction to produce Cumene.
When propylene is released, it principally enters the air, since it is a gas (The
Dow Chemical Company, 2013). There has been little testing into its effects
on aquatic life since propylene is expected to disperse in the air. Similarly,
propylene is not expected to have a significant effect on soil or wastewater. It
is not thought to bioaccumulate and has a short half-life (The Dow Chemical
Company, 2013).
D2. Propane
Propane is an impurity that is present in the propylene feedstock. Propane is
recognised as an alternative clean fuel source (1990 Clean Air Act, USA), and
has no major environmental impacts. Since it has a low amount of carbon, it
does not react to form carbon dioxide when released into the air. At
atmospheric conditions, it is in the gas phase. Thus, when released, it mainly
enters the air, and it will evaporate from soil or water (The Dow Chemical
Company, 2013). Propane has ‘moderate motility’ in ground soil (ChemWatch,
2013), and low bioaccumulation, although it has moderate toxicity to aquatic
animals.
D3. Benzene
Benzene is one of the feedstocks for the production of Cumene. Benzene has
a large impact on the environment. Benzene reacts with atmospheric air to
produce smog. While it is possible that this smog will disintegrate naturally,
contact with rain can cause it to travel to ground soil and contaminate soil
(National Pollutant Inventory, 2013). The benzene from the contaminated soil
leached into ground water. Benzene will only take a few days to break down
in air- in water and soil, it breaks down over a longer period of time (US
Environmental Protection Agency, 2009). Aquatic animals are poisoned by
benzene, causing reproductive issues and curtailing their lifespans. Similarly,
plant life is damaged by benzene contamination, causing growth problems
and death (National Pollutant Inventory, 2013). Bioaccumulation does not
occur in plants or animals (US Environmental Protection Agency, 2009).
D4. Cumene
Cumene is the desired product of the process. Cumene reacts in air with other
chemicals to form isopropyl phenols; in water or soil, it is broken down by
bacteria (National Pollutant Inventory, 2013). Since it is broken down quickly
in air, it is not expected to leave the region in which it was emitted. Like
benzene, it can cause smog. In water, it is degraded in a matter of days. It
has toxic effects on birds, and to a lesser extent, on aquatic animals (National
Group 4
Pollutant Inventory, 2013). It bioaccumulates to a small extent in animals and
plants.
Group 4
Appendix E – HAZOP
Stream Guide Word Deviation Possible Causes Consequences / Concern Action Required
Mix feed NONE No flow No supply Operation ceases Stop plant operation
MORE OF Excess flow Overflow of feed Overflow provide a safe overflow stream
LESS OF Insufficient feed flow Insufficient pump power Pump failure provide a recycle stream that can supply enough feed
OTHER THAN Wrong feed Operator mistake Wrong product produced Stop plant operation
Reactor inlet NONE No flow No supply Operation ceases Stop plant operation
MORE OF Excess flow Overflow of feed Overflow provide a safe overflow stream
LESS OF Insufficient feed flo Insufficient pump power Pump failure provide a recycle stream that can supply enough feed
Reactor outlet NONE No outlet product Reactor outlet pathway blocked Flow buildup in reactor Stop plant operation
Clear the blockage pathway
MORE OF Excess feed to reactor Overflow of feed High pressure in pipe Install pressure relief valve
LESS OF Insufficient feed to reactor Insufficient of feed Low pump power provide a recycle stream that can supply enough feed
MORE THAN High temperature Heat Exchanger wrongly configured Too much energy in reactor feed Adjust the heat exchanger
LESS THAN Low temperature Heat Exchanger wrongly configured Too little energy in reactor feed Adjust the heat exchanger
PART OF Incorrect productcomposition Wrong reaction Wrong product purity Adjust the heat exchanger
Benzene column NONE No top/bottom stream Column malfunction No recycle/bottom stream Adjust the flow appropriately
Stop plant operation
MORE OF Too much top/bottom stream Incorrect reflux ratio Excess recycle/bottom stream Regulate flow valve
LESS OF Low top/bottom stream Incorrect reflux ratio Insufficient recycle/bottom stream Regulate flow valve
PART OF Wrong flow composition Incorrect column configuration Wrong product purity produced Adjust column configuration
Cumene column NONE No top/bottom stream Column malfunction No cumene/DIPB stream Adjust the flow appropriately
Stop plant operation
MORE OF Too much top/bottom stream Incorrect reflux ratio Excess cumene/DIPB stream Regulate flow valve
LESS OF Low top/bottom stream Incorrect reflux ratio Insufficient cumene/DIPB stream Regulate flow valve
PART OF Wrong flow composition Incorrect column configuration Wrong product purity produced Adjust column configuration
Group 4 122 | P a g e
Appendix F – Industrial Catalyst Options
The technology surrounding the type of catalyst used for this reaction is
constantly changing. One of the first established methods for producing
Cumene involved using a packed bed reactor and an acid catalyst. Then,
catalysts for both the gas and liquid phase reaction were developed. Whilst
historically for a gas phase reaction, a phosphoric acid catalyst has been
used, and an aluminium chloride catalyst has been used for liquid phase,
industry is now leaning towards using new zeolite catalysts (Gera et al, 2010).
There are several different catalysts that can be used for the alkylation
reaction. A summary of these catalysts can be found below.
Acidic Catalysts
Acidic catalysts have been used for the alkylation reaction. For the gas phase
reaction, phosphoric acid coupled with a kieseguhr catalyst has been used,
while for the liquid phase reaction, sulphuric acid has been used (Panming et
al, 1992). In the early 1930s, it was discovered that by combining silica and
phosphoric acid, an effective solid catalyst was formed. As the demand for
Cumene developed, this catalyst became very popular. 99% of Cumene
manufacturers used the solid phosphoric acid catalyst in the 1990s, however,
nowadays the zeolite catalyst is more popular (Rase, 2000). The solid
phosphoric acid is toxic and corrosive (Dai, 2012). As this catalyst is
corrosive, consideration needs to be taken when choosing the material of
construction for the reactor, as it must be able to withstand the corrosion.
Aluminium Chloride
Aluminium chloride is a Friedel-Crafts catalyst which can be used for
alkylation reactions (Panming et al, 1992). Aluminium Chloride historically was
a popular catalyst due to its high selectivity for Cumene production and little
undesired by-product production. It was often used in conjunction with
Group 4 123 | P a g e
hydrogen chloride. However, several issues arise through the use of this
catalyst. Disposal of this catalyst is difficult, and is dangerous to transport or
handle (Dai, 2012). Torres-Rodriguez et al (2012) report that this catalyst is
unsustainable. Similarly to the acid catalysts, the material of construction
needs to survive the corrosive conditions of the catalyst.
Zeolite catalyst
Zeolite catalysts are currently being used for the Cumene alkylation reaction
and it appears they will continue to be used in the future. There are several
different types of zeolite catalyst currently being used industrially. Zeolite
catalysts are non-corrosive, re-generative, produce a high yield and have high
product purity. Dai et al (2013) assert that these zeolite catalysts are efficient
and environmentally friendly, which is an advantage over the previous acid
and aluminium chloride catalysts. These catalysts also reduce the formation
of side-products. One type of zeolite, the β-zeolite has good selectivity and
catalytic activity due to the pore size and acidity of the catalyst (Torres-
Rodriguez, 2012). One major advantage of this catalyst is that it is not
corrosive. This means that the construction material of the reactor does not
need to be as corrosive-resistant, which decreases the capital cost of the
reactor.
Group 4 124 | P a g e
Appendix G - Kinetics Conversions
For Rxn-1:
( )
Where:
( )
For Rxn-2:
( )
Where:
( )
Group 4 125 | P a g e
Appendix H – Energy balance calculations
̇ ̇
Group 4 126 | P a g e
Energy Balance across product cooler 1 E-105
̇
Group 4 127 | P a g e
Appendix I - Reactor Performance: Feed ratio
The figures below were determined using case study function in HYSYS. A
reactor volume with a volume 10 times greater than the determined optimal
volume was used so only the effect of feed ratio on conversion was observed.
The reactor temperature was fixed at 358oC and pressure at 25bar. A similar
method was used in Case Study 2.
Conversion
1
0.9
0.8
0.7
0.6
0.5
0.4 Conversion
0.3
0.2
0.1
0
0 2 4 6 8 10
Ratio
Selectivity
160
140
120
100
80
60 Selectivity
40
20
0
0 2 4 6 8 10
Ratio
Figure 79: Benzene and Propylene feed ratio relationship with Cumene
to DIPB selectivity
Group 4 128 | P a g e
Appendix J - Reactor Sizing
Steps:
Determined the maximum volume for reactor
Determined diameter and length base on L/D>10
Checked the conversion for 18.49m and 1.7m2
Group 4 129 | P a g e
Figure 81: Relationship of reactor length and diameter with conversion
Group 4 130 | P a g e
Appendix K – Column Sizing Calculations
K.1 Benzene Column
The calculations used to determine the column height and diameter are listed
below:
[D2]
[D3]
( ) [D4]
[D5]
( ) [D6]
[D7]
( ) [D8]
[D9]
[D10]
[D11]
Column diameter:
[ ] [D12]
( )
[D13]
Height of weir:
[D15]
[D16]
[D17]
Group 4 131 | P a g e
( ( )) [D18]
[D20]
[D23]
Reboiler Pressure:
[D28]
[B12]
( )
[B13]
Hole velocity:
[B14]
[B15]
Group 4 132 | P a g e
C0 depends on percent hole area and ratio of tray thickness to hole diameter.
Typical value is 0.73.
[B16]
( )( ) [B17]
Capacity Parameter:
( ) [B19]
Weir length:
[B21]
[B22]
Group 4 133 | P a g e
Figure 82: Spread sheet - 1D
Another HYSYS spread sheet (Figure 83: Spread sheet - 1C) was used to
calculate active area and downcomer area of the Benzene distillation column.
The calculations and results were:
Single pass:
Column diameter:
[B2]
Weir length:
Group 4 134 | P a g e
[B3]
Flow path length:
[B4]
Active area:
[B5]
Pressure drop:
[B6]
Maximum downcomer backup %:
[B7]
Downcomer area/tower area:
[B8]
Column height
Average temperature:
[A3]
Average viscosity:
[A6]
Group 4 135 | P a g e
Benzene volatility:
[A9]
Cumene volatility:
[C2]
Relative volatility:
[C3]
Actual stage:
[C5]
Column height:
[C5]
[C6]
Group 4 136 | P a g e
Figure 84: Spread sheet - 1H
Group 4 137 | P a g e
K.2 Cumene Column
Group 4 138 | P a g e
Figure 86: Spread sheet - 2D
Group 4 139 | P a g e
Figure 87: Spread sheet - 2H
Group 4 140 | P a g e
Appendix M – Economic Evaluation
Group 4 142 | P a g e
M.2 Spread sheets for the utilities
The following spread sheets show the utilities required for the equipment. The values of the mass flows and heat flows
were calculated on HYSYS. The costs of utilities were given in the design brief.
Energy Stream T in (⁰C) T out (⁰C) Heat flow (kJ/h) Utility Quantity (kg/h) hours per year quantity per year Cost
E-100 Heat2 145 358 2.45E+07 HP steam 1.44E+04 8000 115200000 8 $ / 1000 kg $ 921,600.00
E-102 Heat BR 130.2 140 1.04E+07 LP steam 4739 8000 37912000 3 $ / 1000 kg $ 113,736.00
E-103 Cool1 159.4 100 4.00E+06 Cooling water (30) 1.91E+05 8000 1530400000 20 $/1000m^3 $ 30,735.86
E-104 CoolerSale 153.1 25 3127000 Refrigerated water 149500 8000 1196000000 200 $/1000m^3 $ 23,922.40
E-105 1111 494.8 200 2.99E+07 Cooling water (30) 1.43E+06 8000 11448000000 20 $/1000m^3 $ 229,916.45
E-108 BF1 25 79 7.59E+05 LP steam 345.6 8000 2764800 3 $ / 1000 kg $ 8,294.40
E-109 PP1 25 55 4.36E+05 LP steam 198.7 8000 1589600 3 $ / 1000 kg $ 4,768.80
BenzeneColCon1 140 79.64 1.30E+07 Cooling water 6.20E+05 8000 4957600000 20 $/1000m^3 $ 99,566.20
BenzeneColRebo1 140 176.5 3.83E+06 MP steam 1933 8000 15464000 6.5 $ / 1000 kg $ 100,516.00
CumeneColnCon2 176.5 153.1 6.18E+06 Cooling water 2.95E+05 8000 2362400000 20 $/1000m^3 $ 47,445.37
CumeneColnRebo2 176.5 232.7 5.63E+06 HP steam 3307 8000 26456000 8 $ / 1000 kg $ 211,648.00
Group 4 143 | P a g e
Utilities for impure feed case
Equipment Energy Stream Heat flow kJ/h Utility kW hours per year kWh per yearprice per kWh Cost
P-101 Pump1 111300 Power 30.91666667 8000 247333.3333 0.061 $ 15,087.33
Energy Stream T in (⁰C) T out (⁰C)Heat flow (kJ/h) Utility Quantities kg/h hours per year quantity per year Cost
E-100 Heat2 145 358 25700000 HP steam 15090 8000 120720000 8 $ / 1000 kg $ 965,760.00
E-102 Heat BR 130.2 140 3402000 LP steam 1549 8000 12392000 3 $ / 1000 kg $ 37,176.00
E-103 Cool1 159.4 100 4338000 Cooling water (30) 207400 8000 1659200000 20 $/1000m^3 $ 33,322.62
E-104 CoolerSale 153.1 25 3129000 Refrigerated water 149600 8000 1196800000 200 $/1000m^3 $ 23,938.40
E-105 1111 494.8 200 30710000 Cooling water (30) 1.47E+06 8000 11744000000 20 $/1000m^3 $ 235,861.18
E-108 HO 25 79 7.55E+05 LP steam 343.9 8000 2751200 3 $ / 1000 kg $ 8,253.60
E-109 H1 25 55 4.58E+05 LP steam 208.6 8000 1668800 3 $ / 1000 kg $ 5,006.40
Depropnizer g 100 33.59 3.67E+05 Cooling water 1.75E+04 8000 140320000 20 $/1000m^3 $ 2,818.12
Depropnizer k 100 209.5 8.28E+06 HP steam 4859 8000 38872000 8 $ / 1000 kg $ 310,976.00
BRC Con1 140 79.64 13770000 Cooling water 6.59E+05 8000 5268000000 20 $/1000m^3 $ 105,800.13
BRC Rebo1 140 176.5 3.86E+06 MP steam 1948 8000 15584000 6.5 $ / 1000 kg $ 101,296.00
CumeneColn Con2 176.5 153.1 6171000 Cooling water 295000 8000 2360000000 20 $/1000m^3 $ 47,397.17
CumeneColn Rebo2 176.5 232.7 5584000 HP steam 3279 8000 26232000 8 $ / 1000 kg $ 209,856.00
Total Utilities per year $ 2,102,548.97
Figure 91: Utilities for impure feed case
Group 4 144 | P a g e
M.3 Labour Requirements for the plant
The labour requirements for the plant are the same for both cases. The labour data was
calculated using the provided ‘economic analysis’ spread sheets
Equipment Number of equipment Operators per Shift total operators per shift
Auxiliary Facilities
Air Plants 0 1.0 0
Boilers 0 1.0 0
Chimneys and Stacks 0 0.0 0
Cooling Towers 0 1.0 0
Water Demineralizers 0 0.5 0
Electric Generation Plants 0 0.5 0
Portable Generation Plants 0 3.0 0
Electric Substations 0 0.0 0
Incinerators 0 2.0 0
Mechanical Refrigeration Units 0 0.5 0
Waste Water Treatment Plants 0 2.0 0
Water Treatment Plants 0 2.0 0
Process Equipment
Evaporators 0 0.30 0
Vaporizers 0 0.05 0
Furnaces 0 0.50 0
Fans 0 0.05 0
Blowers and Compressors 0 0.15 0
Heat Exchangers 8 0.10 0.8
Towers 0 0.35 0
Vessels 3 0.00 0
Pumps 1 0.00 0
Reactors 1 0.50 0.5
A single operator works on average 49 weeks per year * (3 weeks’ time off for vacation and sick leave)
8-hour shifts per week: 5
total shift per year 245
process plant normally 3 shifts per day
plant operates 365 days
total operators required for the operation in a given shift 4.469387755
Group 4 145 | P a g e
M.4 Raw materials and products
The following spread sheets show the costs of the raw materials and the revenue from
the products. Note that while DIPB is a product, it must be disposed of, and the disposal
incurs a cost. The values of the mass flows were calculated on Hysys. The costs of the
chemicals were given in the design brief.
Raw Materials for pure feed case
IN Quantity (kg/h)h Quantityper year
Cost Cost per year
Benzene 8564 8000 68512000 1120 $/1000kg $ 76,733,440.00
Pure Propylene 4808 8000 38464000 1570 $/1000kg $ 60,388,480.00
Catalyst 29944 2.52 $/kg $ 75,458.88
OUT
Cumene 12640 8000 101120000 1434 $/1000kg $ 145,006,080.00
DIPB 751.5 8000 6012000 1 $/kg $ 6,012,000.00
OUT
Fuel Gas 280.6 8000 2244800 630 $/1000kg $ 1,414,224.00
Cumene 12610 8000 100880000 1434 $/1000kg $ 144,661,920.00
DIPB 670.3 8000 5362400 1 $/kg $ (5,362,400.00)
Figure 94: Raw material costs and product revenue for impure case
Group 4 146 | P a g e
Summary of Costs for Pure Case
Fixed Capital
Item % of FCI Cost (2011) Cost (2013)
Fixed Capital Investment 100% $ 6,526,032.33 $ 6,312,100.59
Maintenance 7% $ 456,822.26 $ 441,847.04
Total $ 6,439,344.33 $ 6,228,254.34
Utilities
Item Cost (2011) Cost (2013)
Power $ 14,368.89 $ 13,897.86
LP Steam $ 126,799.20 $ 122,642.56
MP Steam $ 100,516.00 $ 97,220.96
HP Steam $ 1,133,248.00 $ 1,096,098.67
Cooling Water $ 407,663.88 $ 394,300.13
Refrigerated Water $ 23,922.40 $ 23,138.19
DIPB (Disposal) $ 6,012,000.00 $ 5,814,918.90
Total $ 7,818,518.37 $ 7,562,217.28
Labour Costs
` % of labour Cost (2011) Cost (2013)
Labour 100% $ 309,364.90 $ 299,223.52
Employee oncost 20% $ 61,872.98 $ 59,844.70
Supervision 115% $ 355,769.63 $ 344,107.05
Laboratory Labour 115% $ 355,769.63 $ 344,107.05
Total $ 1,082,777.14 $ 1,047,282.31
Product Revenue
Item Cost (2011) Cost (2013)
Cumene $ 145,006,080.00 $ 140,252,594.02
Total $ 145,006,080.00 $ 140,252,594.02
Miscellaneous Costs
Item % of TPC Cost (2011) Cost (2013)
Contingencies 15% $ 21,914,801.16 $ 21,196,405.77
Total $ 21,914,801.16 $ 21,196,405.77
Figure 95: Summary of costs for pure feed material
Group 4 147 | P a g e
Summary of Costs for the Impure Case
Fixed Capital
Item % of FCI Cost (2011) Cost (2013)
Fixed Capital Investment 100% $ 7,671,398.63 $ 7,419,920.31
Maintenance 7% $ 536,997.90 $ 519,394.42
Total $ 8,208,396.53 $ 7,939,314.73
Utilities
Item Cost (2011) Cost (2013)
Power $ 15,087.33 $ 14,592.75
LP Steam $ 50,436.00 $ 48,782.64
MP Steam $ 101,296.00 $ 97,975.39
HP Steam $ 1,486,592.00 $ 1,437,859.60
Cooling Water $ 425,199.23 $ 411,260.65
Refrigerated Water $ 23,938.40 $ 23,153.67
DIPB (Disposal) $ 5,362,400.00 $ 5,186,613.62
Total $ 7,464,948.97 $ 7,220,238.33
Labour Costs
Item % of labour Cost (2011) Cost (2013)
Labour 100% $ 309,364.90 $ 299,223.52
Employee oncost 20% $ 61,872.98 $ 59,844.70
Supervision 115% $ 355,769.63 $ 344,107.05
Laboratory Labour 115% $ 355,769.63 $ 344,107.05
Total $ 1,082,777.14 $ 1,047,282.31
Product Revenue
Item Cost (2011) Cost (2013)
Cumene $ 144,661,920.00 $ 139,919,716.03
Fuel Gas $ 1,414,224.00 $ 1,367,863.92
Total $ 146,076,144.00 $ 141,287,579.95
Miscellaneous Costs
Item % of TPC Cost (2011) Cost (2013)
Contingencies 15% $ 18,071,450.12 $ 17,479,044.72
Total $ 18,071,450.12 $ 17,479,044.72
Figure 96: Summary of costs for impure feed case
Group 4 148 | P a g e
M.6 Net Present value calculation tables
The capital is equal to the fixed capital investment (equipment) in the project. The savings column is equal to the product
revenue less the costs of raw materials, utilities, labour and contingencies. Depreciation is equal to 10% of the capital
investment. It is claimed on a linear basis, with no assumed residual value. Depreciation is a tax deduction, so the taxable
savings is equal to the pre-tax cash flow less depreciation. The tax rate is 35%, and the tax paid is the taxable savings
multiplied by the tax rate. The after tax cash flows are equal to the pre-tax cash flows less the tax paid. The discounted
cash flows are used are the after tax cash flows, discounted by the hurdle rate (9%) to bring the values forward to the
present value. The overall NPV is the sum of the discounted cash flows. The cumulative discounted cash flows is the
aggregate of present value of the cash flows. These are used to calculate the payback period.
NPV
$ (149,041,568.86)
Figure 97: NPV calculations for the pure feed
Group 4 149 | P a g e
NPV for the Impure Feed
Year Capital Savings Pre-Tax CF Depreciation Taxable Income Tax Paid After Tax CF Discounted CF Cumulative CF
0 $ 7,939,314.73 $ - $ (7,939,314.73) $ - $ - $ - $ (7,939,314.73) $ (7,939,314.73) $ (7,939,314.73)
1 $ - $ 7,281,570.45 $ 7,281,570.45 $ 793,931.47 $ 6,487,638.98 $ 2,270,673.64 $ 5,010,896.81 $ 4,597,153.04 $ (3,342,161.69)
2 $ - $ 7,281,570.45 $ 7,281,570.45 $ 793,931.47 $ 6,487,638.98 $ 2,270,673.64 $ 5,010,896.81 $ 4,217,571.59 $ 875,409.90
3 $ - $ 7,281,570.45 $ 7,281,570.45 $ 793,931.47 $ 6,487,638.98 $ 2,270,673.64 $ 5,010,896.81 $ 3,869,331.74 $ 4,744,741.63
4 $ - $ 7,281,570.45 $ 7,281,570.45 $ 793,931.47 $ 6,487,638.98 $ 2,270,673.64 $ 5,010,896.81 $ 3,549,845.63 $ 8,294,587.26
5 $ - $ 7,281,570.45 $ 7,281,570.45 $ 793,931.47 $ 6,487,638.98 $ 2,270,673.64 $ 5,010,896.81 $ 3,256,739.11 $11,551,326.37
6 $ - $ 7,281,570.45 $ 7,281,570.45 $ 793,931.47 $ 6,487,638.98 $ 2,270,673.64 $ 5,010,896.81 $ 2,987,834.05 $14,539,160.42
7 $ - $ 7,281,570.45 $ 7,281,570.45 $ 793,931.47 $ 6,487,638.98 $ 2,270,673.64 $ 5,010,896.81 $ 2,741,132.15 $17,280,292.57
8 $ - $ 7,281,570.45 $ 7,281,570.45 $ 793,931.47 $ 6,487,638.98 $ 2,270,673.64 $ 5,010,896.81 $ 2,514,800.14 $19,795,092.71
9 $ - $ 7,281,570.45 $ 7,281,570.45 $ 793,931.47 $ 6,487,638.98 $ 2,270,673.64 $ 5,010,896.81 $ 2,307,156.09 $22,102,248.80
10 $ - $ 7,281,570.45 $ 7,281,570.45 $ 793,931.47 $ 6,487,638.98 $ 2,270,673.64 $ 5,010,896.81 $ 2,116,656.96 $24,218,905.76
NPV
$ 24,218,905.76
Figure 98: NPV calculations for impure feed case
Because the pure feed never has a positive cash flow, there is no practical DCF ROR
Group 4 150 | P a g e
M.8 Payback period
The payback period can be estimated by the cumulative cash flow column on
the NPV calculations. They can also be calculated by graphing the
relationship between time and the cumulative cash flow. If all of the cash flows
are the same (as in this project) this line is linear. Using the equation of the
line, let Y (cumulative cash flow)=0. The value of X (years) will be equal to the
time it takes to payback the initial investment.
Sample calculation:
For Figure 67: Payback period for the impure feed Cumene plant, the slope of
the line is
Rearranging,
Pure Feed
Using the data from the calculations of the pure feed NPV, the following
relationship is ascertained.
$-
$(20,000,000.00) 0 2 4 6 8 10 12 14
$(40,000,000.00)
Cumulative Cash Flow
$(60,000,000.00)
$(80,000,000.00)
$(100,000,000.00)
$(120,000,000.00)
$(140,000,000.00)
$(160,000,000.00) y = -1E+07x - 2E+07
$(180,000,000.00)
$(200,000,000.00)
Years
It can be observed from Figure 99: Payback period for pure feed case that
there will be no payback period. This is because the plant never has a positive
cash flow.
Group 4 151 | P a g e
M.9 Return on Investment
The return on investment is given by
Group 4 152 | P a g e
Appendix N – Design Project Meeting Minutes
1. Communication
a. Use group wiki and blog as main form of communication
2. Project
a. Discussion of scope of project
b. Agreed that computer simulation should be done as a group
c. Caroline keen to do part 4- the finance section
d. Sophia keen to the hazop and safety section
3. Meetings
a. Meetings on Wednesdays at 1pm during the allotted time.
b. Possibility of meeting starting at 11 on Wednesdays and finishing at 3,
as everyone has free periods during those times
Group 4 153 | P a g e
Meeting Minutes (Week 2)
CHEMENG3030: Simulation and Conceptual Design / CHEMENG 3025:
Pharmaceutical Plant Design and Process Engineering
Date of Meeting 31 July 2013
Location Level 1 Innova
Start time and end 1:15-2pm
Confirmation of previous meeting Confirmed
minutes are a true and accurate
record
Attendance Caroline Jackson
Sophie Nicholls
Luoshan Wang
Ahmad Nazmi Ramlan
Minutes prepared by Caroline Jackson
1. Project
a. Need to start looking for literature articles for background theory
b. PFD needs to be started
i. Look for diagrams for assistance, and change them based on
literature articles
c. Update myuni group wiki with important information regarding
research findings
2. Meetings
a. Potentially book a room in the hub for meetings at 1pm on
Wednesdays, but need to check in with Dr Zhang before leaving class
Group 4 154 | P a g e
Meeting Minutes (Week 3)
1. Project
a. Discussed where what had been done in the previous week
i. Caroline and Sophie had updated the wiki with information
from the project manual, and also from literature reviews
ii. The literature reviews gave details on the kinetics of the
reactor and also some reaction details
iii. Discussed some different PFDs from textbooks and literature
studies
iv. Decided that each new wiki needs to include a ‘fun fact’ about
the process in order to encourage teamwork and create a
stimulating environment
2. Meetings
a. Next meeting will be on Monday in hub room 336 from 11-1
i. The BFD and PFD needs to be done by this stage
ii. Every group member needs to do some research so that they
have some idea of what the diagram will look like.
Group 4 155 | P a g e
Meeting Minutes (Week 4a)
CHEMENG3030: Simulation and Conceptual Design / CHEMENG 3025:
Pharmaceutical Plant Design and Process Engineering
Date of Meeting 19 August 2013
Location Hub room 336
Start time and end 11-1pm
Confirmation of previous meeting Confirmed
minutes are a true and accurate
record
Attendance Caroline Jackson
Sophie Nicholls
Luoshan Wang
Ahmad Nazmi Ramlan
Minutes prepared by Caroline Jackson
1. Project
a. Discussed where what had been done in the previous week
i. PFD was worked on
ii. A few alternatives were proposed, which were very similar, but
with certain unit operations differing
iii. BFD and PFD will be drawn on HYSYS before the next
meeting
2. Task allocation
a. On Wednesday, there is a check on the progress of the group
i. Project outline/summary: Ahmad
ii. Chemical key properties: Caroline
iii. Means-end analysis step by step: Luoshan
iv. BFD: Sophie (after group collaboration on process)
v. Draft PFD: Sophie
vi. Process conditions: Group Collaboration
vii. Project Scope: Group discussion
viii. HYSYS: initial PFD: Sophie
Simulation: Luoshan, Sophie & Group collaboration
ix. Hazop and Safety: Luoshan
x. Economic Analysis: Caroline
xi.
3. Meetings
a. Next meeting will be on Wednesday
i. There is a progress check during this time
Group 4 156 | P a g e
Update wiki Everyone 14/8/13 In progress
BFD and PFD Sophie 21/8/13 Complete
Project Ahmad 21/8/13 Complete
outline/summary
Chemical key Caroline 21/8/13 Complete
properties
Means-end Luoshan 21/8/13 Complete
analysis step by
step
Process everyone 21/8/13 Complete
conditions
Project Scope everyone 21/8/13 Complete
Initial HYSYS Sophie 21/8/13 Complete
Group 4 157 | P a g e
Meeting Minutes (Week 4b)
CHEMENG3030: Simulation and Conceptual Design / CHEMENG 3025:
Pharmaceutical Plant Design and Process Engineering
Date of Meeting 21 August 2013
Location LG Napier
Start time and end 1:10-2pm
Confirmation of previous meeting Confirmed
minutes are a true and accurate
record
Attendance Caroline Jackson
Sophie Nicholls
Luoshan Wang
Ahmad Nazmi Ramlan
Minutes prepared by Sophie Nicholls
1. Project
a. PROGRESS CHECK with Dr Zhang
i. Project Outline
ii. Task Allocation (subject to change)
iii. Key Properties of Chemicals
iv. Means End Analysis step by step
v. BFD options
vi. Draft PFD HYSYS
vii. Detailed process conditions
2. Task allocation
a. Take into account Dr Zhang’s suggestions for BFD and PFD design
b. Read literature surveys to determine temperatures and pressures of
streams (will decide on best option and then adjust with HYSYS
simulation)
c. Luoshan to add diagram to means-end
d. Ahmad to begin material balances once final PFD base case has been
confirmed
3. Meetings
a. Next meeting will be on Wednesday
b. Meeting Monday 26/8 11-1 in Hub
Group 4 158 | P a g e
Meeting Minutes (Week 5a)
CHEMENG3030: Simulation and Conceptual Design / CHEMENG 3025:
Pharmaceutical Plant Design and Process Engineering
Date of Meeting 2 September 2013
Location Hub Room
Start time and end 11-1pm
Confirmation of previous meeting Confirmed
minutes are a true and accurate
record
Attendance Caroline Jackson
Sophie Nicholls
Luoshan Wang
Ahmad Nazmi Ramlan
Minutes prepared by Sophie Nicholls
1. Project
a. Discussed Dr Zhang’s suggestions
b. Drew BFD on whiteboard and added numerous literature values for
temperature and pressure of each stream
c. Came to a decision of the stream properties (temperature and
pressure)
d. Decided base case would be Pure Benzene and Propylene with 5%
Propane impurity Feed because impure Propylene is cheaper and
impurity can be sold as fuel gas
e. Decided to not include a distillation column to separate fuel gas into
individual components due to cost
2. Task allocation
a. Caroline to put unit operations into HYSYS
b. Begin setup of reactions and base case in HYSYS
3. Meetings
a. Next meeting will be on Wednesday
Group 4 159 | P a g e
Meeting Minutes (Week 5b)
CHEMENG3030: Simulation and Conceptual Design / CHEMENG 3025:
Pharmaceutical Plant Design and Process Engineering
Date of Meeting 2 September 2013
Location Hub Room
Start time and end 11-1pm
Confirmation of previous meeting Confirmed
minutes are a true and accurate
record
Attendance Caroline Jackson
Sophie Nicholls
Luoshan Wang
Ahmad Nazmi Ramlan
Minutes prepared by Sophie Nicholls
1. Project
a. Discussed Dr Zhang’s suggestions
b. Drew BFD on whiteboard and added numerous literature values for
temperature and pressure of each stream
c. Came to a decision of the stream properties (temperature and
pressure)
d. Decided base case would be Pure Benzene and Propylene with 5%
Propane impurity Feed because impure Propylene is cheaper and
impurity can be sold as fuel gas
e. Decided to not include a distillation column to separate fuel gas into
individual components due to cost
2. Task allocation
a. Caroline to put unit operations into HYSYS
b. Begin setup of reactions and base case in HYSYS
3. Meetings
a. Next meeting will be on Wednesday
Group 4 160 | P a g e
Meeting Minutes (Week 6)
CHEMENG3030: Simulation and Conceptual Design / CHEMENG 3025:
Pharmaceutical Plant Design and Process Engineering
Date of Meeting 11 September 2013
Location LG Napier
Start time and end 1-2pm
Confirmation of previous meeting Confirmed
minutes are a true and accurate
record
Attendance Caroline Jackson
Sophie Nicholls
Luoshan Wang
Ahmad Nazmi Ramlan
Minutes prepared by Sophie Nicholls
1. Project
a. HYSYS base case simulation has encountered convergence errors,
email Dr Zhang for assistance
b. Hand material balances were discussed
c. Utilities were further discussed and decided on initial requirements
d. Prepare required documents for progress review meeting next week
2. Task allocation
a. Fully specified PFD
b. Specified utilities
c. Plant location
3. Meetings
a. Next meeting will be Progress review with Dr Zhang on Wednesday
11/9
Group 4 161 | P a g e
Meeting Minutes (Week 7)
CHEMENG3030: Simulation and Conceptual Design / CHEMENG 3025:
Pharmaceutical Plant Design and Process Engineering
Date of Meeting 18 September 2013
Location LG Napier
Start time and end 1-2pm
Confirmation of previous meeting Confirmed
minutes are a true and accurate
record
Attendance Caroline Jackson
Sophie Nicholls
Luoshan Wang
Ahmad Nazmi Ramlan
Minutes prepared by Sophie Nicholls
1. Project
a. Progress review meeting with Dr Zhang
i. PFD with stream T, P, vapour fraction and components
specified
ii. Specified unit operations
iii. Specified utilities
iv. Raw material and product specifications
v. Plant Location
2. Task allocation
a. Continue base case on HYSYS
b. Think about report structure and appendices
3. Meetings
a. Next meeting will be on Wednesday 25/9
Group 4 162 | P a g e
Meeting Minutes (Week 8)
CHEMENG3030: Simulation and Conceptual Design / CHEMENG 3025:
Pharmaceutical Plant Design and Process Engineering
Date of Meeting 2 October 2013
Location LG Napier
Start time and end 1-2pm
Confirmation of previous meeting Confirmed
minutes are a true and accurate
record
Attendance Caroline Jackson
Sophie Nicholls
Luoshan Wang
Ahmad Nazmi Ramlan
Minutes prepared by Sophie Nicholls
1. Project
a. Discussed Dr Zhang’s suggestion from progress review
i. PFD needed to be changed – another groups was similar
ii. Base case requires sizes/T/P of reactor, columns and heat
exchanger
b. Need to discuss project less with other groups to avoid similarities with
PFD
2. Task allocation
a. Continue base case on HYSYS
b. Begin thinking about designing reactor, columns, heat exchanger
3. Meetings
a. Next meeting will be on Wednesday 2/10
Group 4 163 | P a g e
Meeting Minutes (Week 9)
CHEMENG3030: Simulation and Conceptual Design / CHEMENG 3025:
Pharmaceutical Plant Design and Process Engineering
Date of Meeting 9 October 2013
Location LG Napier
Start time and end 1-2pm
Confirmation of previous meeting Confirmed
minutes are a true and accurate
record
Attendance Caroline Jackson
Sophie Nicholls
Luoshan Wang
Ahmad Nazmi Ramlan
Minutes prepared by Sophie Nicholls
1. Project
a. Need to begin designing reactor, column and HEX
b. Base case is currently converged with conversion >90%
c. Ensure that overall conversion is less than Gibbs
d. Ensure that Cumene produced is to specifications of client
2. Task allocation
a. Begin column, reactor, HEX design
b. Next meeting will discuss choice for second case study
3. Meetings
a. Next meeting will be on Wednesday 16/10 Progress Review
Group 4 164 | P a g e
Meeting Minutes (Week 10)
CHEMENG3030: Simulation and Conceptual Design / CHEMENG 3025:
Pharmaceutical Plant Design and Process Engineering
Date of Meeting 9 October 2013
Location EngNorth
Start time and end 10-11am
Confirmation of previous meeting Confirmed
minutes are a true and accurate
record
Attendance Caroline Jackson
Sophie Nicholls
Luoshan Wang
Ahmad Nazmi Ramlan
Minutes prepared by Sophie Nicholls
1. Project
a. Progress Review meeting with Dr Zhang
i. Reactor Design – good
ii. Column design – temperature of benzene recycle leaving
column should be 70-80 not ~2.3
2. Task allocation
a. Adjust column design for suitable benzene recycle temperature
b. Report writing
3. Meetings
a. Next meeting will be on Wednesday 23/10
Group 4 165 | P a g e
Meeting Minutes (Week 11)
CHEMENG3030: Simulation and Conceptual Design / CHEMENG 3025:
Pharmaceutical Plant Design and Process Engineering
Date of Meeting 23 October 2013
Location LG Napier
Start time and end 1-2pm
Confirmation of previous meeting Confirmed
minutes are a true and accurate
record
Attendance Caroline Jackson
Sophie Nicholls
Luoshan Wang
Ahmad Nazmi Ramlan
Minutes prepared by Sophie Nicholls
1. Project
a. HYSYS simulation almost finalised and completed
b. Report is almost finished
i. Appendices have been started
ii. Literature survey underway
iii. Mass/energy Balances almost completed
iv. HYSYS summaries almost complete
2. Task allocation
a. Finalise HYSYS simulation by Friday
b. Continue writing report due Friday 1 November
3. Meetings
a. Next meeting will be on Wednesday 30/10
Group 4 166 | P a g e
Meeting Minutes (Week 11)
CHEMENG3030: Simulation and Conceptual Design / CHEMENG 3025:
Pharmaceutical Plant Design and Process Engineering
Date of Meeting 30 October 2013
Location LG Napier
Start time and end 1-2pm
Confirmation of previous meeting Confirmed
minutes are a true and accurate
record
Attendance Caroline Jackson
Sophie Nicholls
Luoshan Wang
Ahmad Nazmi Ramlan
Minutes prepared by Sophie Nicholls
1. Project
a. HYSYS simulation has been finalised
b. Report is almost finished
i. Editing and formatting to be done
2. Task allocation
a. Final edit and format
3. Meetings
a. FINAL MEETING TODAY
Group 4 167 | P a g e