Kanhaiya Iocl
Kanhaiya Iocl
Kanhaiya Iocl
INTRODUCTION
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Indian Oil has been ranked 2nd among 14 national companies in India. Indian
Oil is only oil company in the country with ISO 9001/9002 accreditation for
over 50 units which includes Refineries, Pipelines, Aviation, Fuel station,
Grease plants, Quality control laboratories, LPG bottling plants and Indian
Oil Institute of Petroleum Management. By the end of 1997 all the eight
refineries of Indian oil were accredited ISO 14001 certification for
environmental management system. Indian Oil natures the vision of
becoming an integrated & diversified global energy corporation. It is
augmenting infrastructure and expanding into explanation & production of
crude oil, petrochemicals, power generating LNG & fuel management. It is
globalizing its R&D, training & consultancy services and marketing
lubricants.
Indian Oil is Public Sector undertaken (PSU) company registered under the
companies act 1956 and is managed by the Board of Directors who are
appointed by the President of India. Indian Oil owns and operates seven of
the countries’ fourteen refineries which are as follows :
UNIT CAPACITY
1. Barauni Refinery 06 .00 MMTPA.
2. Panipat Refinery 12.00 MMTPA.
3. Mathura Refinery 08.00 MMTPA.
4. Koyali Refinery 13.70 MMTPA
5. Guwahati Refinery 01.00 MMTPA.
6. Haldia Refinery 9.60 MMTPA.
7. Digboi Refinery 00.65 MMTPA
*MMTPA = Million Metric Tons Per Annum
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All the seven refineries attained a record crude throughput 33.1 million tones,
The previous best of 32.4 million tones achieved in 1999-2000. All the tree
awards in the refining sector instituted by Minister of Power under National
Energy Conservation Award 2000 were bagged by IOCL Refineries which
was the third time in succession.
BIHAR
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the river Ganga at Begusarai District town of Bihar state. The refinery is
strategically located on the crossroads of two important national highways,
NH-30 & NH-31 and two important railways, Eastern railways & North
Eastern railways. The river Ganga flows around 8 km away from the refinery.
The barauni refinery takes its crude oil from foreign countries through
Barauni-Haldia crude pipeline (BHCPL). Barauni refinery is one of the
biggest size oil refinery owned and managed by IOCL. The refinery is
located about 8 km from the town Begusarai and is surrounded by villages.
The construction activity of the refinery commenced in 1962 and it went on
stream in the year 1964 facing insurmountable hurdles, heavy equipment,
men and machinery was moved into a predominantly agrarian district of
Begusarai.
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Barauni refinery was formally inaugurated by Prof. Humayun Kabir, the then
union minister of petroleum & chemicals, govt. of India on January 15, 1965.
It started with a refining capacity of processing 2 Million Metric Tons Per
Annum of Assam crude through the Nahar-Katiya-Barauni pipeline. The
capacity was subsequently enhanced to 3 Million Metric Tons Per Annum.
The refinery consists of three crude oil distillation units, two coker units,
CRU, LRU and BXP. The oil movement and storage section of refinery does
the storage and dispatch of all the products. An LPG bottling plant has also
been provided which is able to fill 3500 to 4000 cylinders per day. A captive
Power plant has been provided to meet the steam and power requirements of
the refinery. Under the expansion program following process units &
facilities are put up :
The Resid Fluidized Bed Catalytic Cracking Unit (RFCCU)
yielding LPG, Diesel and Petrol.
Diesel Hydro Treater (DHDT) for improving Cetone no. and to
meet the Euro emission norms.
Primarily, the refining technology was sourced from the eastern countries like
Russia. Later as the refinery grew over the years, it drew upon technologies
from rest of the world. In February 16 1999the 498 km long Haldia-Barauni
crude oil pipeline commenced its crude supply position of the refinery, which
was dependent on Assam crude alone. At the beginning of the new
millennium, Barauni refinery is poised to touch stellar heights in modernizing
its refining technology. Barauni refinery is among the few refineries in the
world to have scored the coveted ISO 9002 certification.
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The refinery processes imports the low sulphur and high sulphur crude oil to
produce the following:
Barauni refinery, the lifeline of Bihar meets the demand of vital petroleum
products not only of the state but also for sustenance growth of industries all
around. It has been acting as a great synthesizer of a traditionally agrarian
economy with industrial development ushering in prosperity. So, the refinery
is often called as a luminous jewel, reflecting the development of Bihar.
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Production Department
This Department basically takes care of all the production related issues in
the refinery.
Internal Audit
The function of the Internal Audit may be expressed as an independent
appraisal activity within the organization for the review of operations as a
service to management. It is a managerial control which functions by
measuring and evaluating the effectiveness of other Internal Audit.
Inspection Department
Inspection Department is a service Department. The function of the
department is to provide technical backup to Production & Maintenance
department in terms of unit operation, monitoring & inspection of static
equipments such as furnace, vessels, lines, column to prevent failures and
recommends necessary repairs. The department also inspects new jobs under
project and modification.
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Technical Service
Technical Service department is a service department of Barauni refinery. The
main function of the department is to provide technical backup to production
department for monitoring unit operations, fuel & loss, utilities and chemical.
The department monitors refinery emissions/effluent to meet environmental
regulation and safety performance. Accordingly, the department is divided
into different cells to perform various functions.
The various cells are as follows :
Planning & Co-ordination and Cost & Economic Cell (P&C and CEC)
Process Cell
Technical Audit Cell
Off-Site Cell
Safety Audit Cell
Quality Control Department
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This department monitors the quality of initial to final product i.e. from crude
to each extracted product (MS, Diesel etc.)
Table-1
PRODUCT WT%
Gas 3.41
IFO 4.97
LPG 5.07
SRN 3.71
MS 10.8
SK 16.1
HSD 49.77
LDO 1.17
LSHS 0.97
CBFS 0.51 10
RPC 3.37
Sulphur 0.15
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ABBREVIATIONS
Abbreviations used have been explained below;
UNITS
AVU : Atmospheric Vacuum Unit
CDU : Crude Distillation Unit
VDU : Vacuum Distillation Unit
NSU : Naphtha Splitter Unit
HTU : Naphtha Hydrotreater Unit
CRU : Catalytic Reforming Unit
RFCCU: Resid Fluid Catalytic Cracking Unit
DHDT: Diesel Hyerotreater Unit
HGU : Hydrogen Generation Unit
SWSU: Sour Water stripper unit
ARU : Amine Regeneration Unit
SRU : Sulphur Recovery Unit
PRODUCT STREAMS
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PROCESS UNITS
SUMMARY OF FEED AND YIELD PATTERN OF UNITS
The capacity, feed source along with product usage has been tabulated below;
Table-2
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COKER -A,B
Table-3
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Table-4
Table-5
Table-6
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Table-7
Table-8
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HCN,LCO RFCCU
Hydrogen HGU
AVU-I,II,III
Barauni Refinery has , three AVUs, the capacity and yield pattern of these
units have been tabulated in table-1. In AVU-I &II only low sulphur crude
can be processed while the AVU-III has been revamped to process High
sulphur Crude also. Due to the limitations of secondary processing units only
15% High sulphur Crude can be processed.
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In main fractionator column , the top product is naphtha and gas side draws
are Heavy naphtha, Kerosene , LGO,HGO and bottom product is RCO.
The Kerosene, LGO and HGO Circulating refluxes Remove the heat from the
column and maintain vapor liquid traffic in side the column.
Above 3600C the, heavy hydrocarbons starts cracking, to avoid cracking and
separate the middle distillate and feed for RFCCU the RCO is heated to
3870C and is futher distilled in vacuum distillation column.
In vacuum distillation column , the top cut is light vacuum gas , oil next6
comes the Heavy Vacuum gas oil and bottom product is vacuum resid
Table-9
PRETOPPING COLUMN
Pressure, kg/cm² g 1.5
Top temperature °C. 122
Feed temp. °C. 233
Bottom temp. °C. 250
Reflux ratio .56
Hy. Nap. draw off temp. °C. 164
Hy. Nap. Vapor return temp. °C. 155
MAIN FRACTIONATING COLUMN
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STABILIZER COLUMN
Pressure, kg/cm²g ( top) 8.2
Top temperature, oC 61
Feed temp. °C. 120
Liquid temperature ex. reboiler 164.5
Reflux ratio 4.0
VACUUM COLUMN
COKER-A/B
Coker-A with 0.6 MMTPA capacity was commissioned in 1964, in 1986
Additional Coker unit Coker-B of 0.5 MMTPA capacity , along with LPG
recovery unit was commissioned.
The unit can process vacuum residues from a wide variety of crude oils
including Bonny Light and Arab Mix crudes. It can also process a number of
low value streams such as extracts and de-waxed oils from Lube plants and
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decant oil from the bottom of FCC units in addition to the vacuum residues.
The unit is designed to produce unstabilised Naphtha, LPG rich off gas,
Reduced Petroleum Coke (RPC) and components for HSD, LDO and Fuel
Oil pools. The unit upgrades heavy residual oil converting it to above
products.
Table-10
OPERATION DURATION
Switch over 1 hr.
Steam cooling 1 ½ hrs.
Water cooling 4 hrs.
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Water draining (unbolting of the covers and flanges is done during 2 hrs.
draining)
Covers removal, placement of fixture & drill bit etc. 2 hrs.
Drilling 1 ½ hrs.
Tools position change ½ hrs.
Coke cutting 1 ½ hrs.
Boxing up 4 hrs.
Pressure testing 1 hr.
Vapor heating 5 hrs.
The Gases and unstablised naphtha & gases from main fractionator over
head are sent to LRU where these are separated into Gas ,LPG and
MRN. The side cuts are Coker Kero, Coker Gas Oil and Coker Fuel Oil.
Coker Kero and Coker Gas Oil are DHDT feed components while Coker
fuel oil is a good RFCCU feed component. The RFO is used as Refinery
fuel
The Kero, Gas Oil and CFO circulating refluxes controls the thermal balance
as well as
Vapour and liquid traffic inside main fractionators.
The typical operating conditions of Main fractionators and coke chamber are
tabulated in table-11
Table-11
O
LDO CR Return to Main C 269
fractionator
O
Furnace Coil outlet Temperature C 498-500
O
Kero draw temp. C 213-220
O
CGO draw temp. C 329-333
O
CFO draw tempo. C 397-406
O
Coke Chamber Overhead C 450
Temperature (before quench)
O
Coke Chamber Overhead C 440
Temperature (after quench)
K- Main fractionator 1 Flash zone Kg/cm2g 2.83
Pressure
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This unit has been liciensed by M/s. IFP (Institute Francais Du Petrole),
France and IIP (Indian Institute of Petroleum) Dehradun In this unit high
Octane Reformate is produced for the purpose of blending in MS. Heavy
naphtha from NSU is first Hydrotreated to produce desulphurised naphtha.
DSN is free of nitrogen and oxygen also.The hydrotreating reactions are
carried out at pressure of 16.0 Kg/ Cm 2 and 3300C temperature.
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A) DEHYDROCYCLISATION OF PARAFFINS
B) DEHYDROGENATION OF NAPHTHENE
C) ISOMERIZATION OF PARAFFIN
A) DEHYDROCYCLISATION OF PARAFFINS:
this reaction than for the other ones thus leading to coke formation.
B) DEHYDROGENATION OF NAPHTHENE :
B) ISOMERISATION OF PARAFFINS :
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C6H14 CH3-CH-CH2-CH2-CH3
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CH3
N-HEXANE ISO-HEXANE
Along with Reformate , the CRU Off gases with 95 vol% Hydrogen
purity are produced, before 6th May’05 these gases were used as fuel,
PSA of HGU has already been revamped for purifying Hydrogen from
these gases, a Chloride guard, designed in-house has already been put to
remove chlorides from these gases to protect the heat exchangers of
DHDT from deposition of Ammonium Chloride.
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OBJECTIVES OF RFCC :
The Resid Fluidized Catalytic Cracking Units are designed and operated to
crack heavier, low value feed stocks (VGO, unconverted hydrocracker
bottoms, RCO, resid streams etc.) to high valued products. The process can
adapt very well to the changing refining scene as well as market demands.
The RFCC unit utilises micro spheres of catalyst particles which are fluidised
by proper aeration. Due to catalytic cracking, coke deposition takes place on
the catalyst particles making it less active. The catalyst is transferred in
fluidized state to regenerator where the coke is burnt off and the catalyst gets
regenerated and resent to reactor for reaction. The heat liberated by burning
coke in the regenerator is utilised as heat of cracking reactions in the reactor,
thus making the process heat balanced. The Hydrocarbon vapours are
separated into following products in the Fractionation and Gas concentration
section.
Fuel Gas
L.P.G.
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CATALYTIC REACTIONS
The catalytic reactions can be classified into two broad categories
In a RFCC unit, when the feed contacts the regenerated catalyst, the first step
is the vapourisation of the feed by the catalyst. Subsequently formation of
positively charged carbon atoms called carbocations takes place.
Carbocation is a generic term for a positively charged carbon ion. Carocation
can be further divided into carbenium and carbonium ions.
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2. Isomerisation reactions
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6. Dehydrogenation reactions
7. Dealkylation
8. Condensation
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Oil feed to the riser is preheated before entering the reaction system. Preheat
temperature along with regenerated catalyst temperature is controlled to
result in an optimum catalyst to oil ratio. Injection of a passivator into the
fresh feed just ahead of the feed injectors acts to inhibit the undesirable
effects of nickel in the feed. Shutdown valves will stop the feed flow to the
riser and divert it through an auxiliary line to the main fractionator in case of
certain emergencies. A flow controller regulates the total feed to the unit and
hand-controlled globe valves adjust flow to each feed injector. The feed
should be split evenly between the feed injectors as observed by individual
flow indicators. Pressure on each feed injector should be monitored as a
verification of flow and an indication of nozzle condition.
Dispersion steam is supplied to each fresh feed injector to promote fresh feed
atomization and vaporization. The total dispersion steam is flow controlled
with flow to each feed injector adjusted by hand-controlled globe valves.
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Upon emergency shutdown, the steam flow control valve automatically opens
to clear the riser of oil and catalyst with steam
Optional, heavy cracked naphtha may be recycled to the fresh feed injectors
to reduce the viscosity of the material flowing through the injectors. This
optional feature should be used when the fresh feed CCR is above 4.0 wt%.
Reduction of the oil viscosity will promote the atomization of the injectors.
During turndown operations, the oil feed can be taken out of two opposing
feed injectors and operation continued with the two remaining injectors. This
will ensure adequate atomization at lower change rates. Steam flow is
required in the two idle injectors (775 kg/hr/injector) to keep them clear.
CO Incinerator;
The flue gas from the first stage regenerator passes through a double disc
slide valve for controlling the differential pressure between the first and
second stage regenerators. Immediately downstream of the first stage
regenerator slide valve is an orifice chamber designed to reduce the flue gas
pressure to 0.09 kg/cm2g. This orifice chamber is refractory lined to allow a
carbon steel shell construction. As the flue gas passes through each plate,
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the pressure is gradually reduced to the desired outlet pressure. Since the
orifice chamber's pressure reduction is directly affected by the flue gas
volumetric flow rate, the slide valve must have wide pressure reduction
ability.
This pressure reduction is a direct function of the flue gas volumetric flow
rate. To compensate for the varying flow rates through the orifice chamber,
the slide valve provides the flexibility needed in processing varying fresh
feed rates.
The carbon monoxide rich flue gas exits the orifice chamber and enters the
CO Incinerator to convert the carbon monoxide to carbon dioxide to comply
with environmental emission requirements. This CO Incinerator burns
auxiliary fuel gas required to heat the incoming CO rich flue gas and
supplemental combustion air to 982°C. At this temperature the CO reacts
with the oxygen from the auxiliary air and converts to CO2.
Flue gases from the second stage regenerator pass through a double disc
slide valve controlling this regenerator's dilute phase pressure. Immediately
downstream of the second stage regenerator slide valve is an orifice
chamber designed to reduce the flue gas pressure to 0.08 kg/cm2g.
Hot CO Incinerator effluent combines with the second stage regenerator flue
gas. This combined flue gas presses through a cooler where the flue gas
thermal energy is recovered by generating high pressure superheated steam
and superheating high and medium pressure steam generated in the gas
plant. The flue gas is cooled to approximately 28°C above the sulfur dew
point. The cooled flue gases flow to flue gas scrubber, which further
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reduces the catalyst fines and SOX in the flue gases. The off-gases from the
flue gas scrubber flow to a stack.
The fractionators consist of 30 valve trays and 8 rows of shed decks. The
reactor effluent, comprised of cracked hydrocarbon vapors, steam and inert
gases, enters the fractionators at the bottom of the quench section. In this
section of the fractionators the superheated cracked vapors and inert are
cooled, and the bottoms product is condensed.
The small amount of entrained catalyst in the cracked vapors is scrubbed out
and drops to the bottom with the condensed product. The slurry pump
around and the decanted oil product are withdrawn from the bottom of the
fractionators and pumped through the fresh feed preheat exchangers, the
high and medium pressure steam generators and the boiler feed water
preheaters. Depending on the fresh feed preheat requirements, the duties of
these exchangers will vary. Slurry pump around return temperature is
controlled at the boiler feed water preheaters by temperature control on the
water bypass.
The cooled slurry pump around stream is returned on flow control to the top
of the shed decks in the quench section. The decanted oil product is drawn
from the slurry pump around stream after the fresh feed preheat on flow
control reset by fractionators bottoms level control. Entrained catalyst is
removed from the decant oil in the slurry oil filter. The decanted oil then
preheats the fresh feed. The decanted oil is pumped before being cooled to
the required battery limit temperature by an air cooler.
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The slurry filter is back washed with heavy cycle oil material pulled from
the HCO pumparound (vendor to verify back wash material). Backwash
from the slurry filter carries catalyst fines removed from the decant oil back
to the riser. Vent gas from the filter backwash operation can be vented back
to either the fractionator or the flare. The filter vendor will decide the final
destination during detail engineering.
Heavy cycle oil (HCO) pumparound, recycle and reflux are withdrawn
from a total draw chimney tray. The reflux is pumped back to the wash
trays below the HCO chimney tray on chimney tray level control. The HCO
pumparound circuit is utilized to preheat boiler feed water and reboils the
debutanizer tower. The HCO pumparound stream is then returned two trays
above its drawoff chimney tray. The total pumparound flow control is
maintained by adjusting the bypass stream around the debutanizer reboiler.
The pumparound return temperature is controlled by varying the amount of
boiler feed water, which bypasses the HCO PA/BFW exchanger.
In the fractionator the HCO pumparound is used to further cool the cracked
vapors from the slurry section, condense the wash oil for the wash oil trays
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above the quench section, condense the HCO recycle and control the
internal reflux above the HCO section.
The HCO recycle flows to the HCO stripper on stripper level control where
it is stripped on light components by the use of steam. Good stripping
ensures a minimum of light components that can crack into undesirable
products once the HCO is reintroduced into the riser. The stripped HCO
recycle is pumped on flow control and cooled against fresh feed. The HCO
recycle temperature to the riser is controlled by manually bypassing fresh
feed around the HCO Recycle/Feed exchangers. The cooled HCO recycle is
sent to the riser.
The Light Cycle Oil (LCO) pumparound and product are withdrawn from a
partial drawoff chimney tray. The LCO PA is cooled down against the
stripper reboiler, fresh feed, boiler feed water, demineralized water and
finally an air cooled before returning to the fractionator. Most of the LCO
PA flow passes through the stripper reboiler. The total pumparound flow is
maintained with a flow control at the bypass of the stripper reboiler. The
return temperature is controlled by regulating the flow through the LCO PA
air cooler. To maintain the total LCO PA flow requirement, flow is
bypassed around the air cooler on differential pressure control. The LCO
PA is used to control the reflux above the LCO section and to condense the
LCO product.
The LCO product flows to the LCO stripper on stripper level control where
it is stripped of light components by using steam. Good stripping ensures a
satisfactory TCO flash point. The stripped LCO is pumped on flow control
and cooled against fresh feed and air. The cooled net LCO product is
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blended with HCN to produce the TCO product. A portion of the cooled
LCO stream is used for gland and flushing oil.
Sponge absorber lean oil is drawn out of the fractionator from a partial
drawoff chimney tray. The lean oil is pumped on flow control and cooled
against fresh feed and cooling water before being used as lean oil in the
sponge absorber. The rich oil from the bottom of the absorber is then
returned to the fractionator to recover the light ends absorbed in the sponge
absorber.
Part of the lean oil can be recycled back to the reactor riser for Mixed
Temperature Control (MTC).
The net hydrocarbon liquid plus the top reflux and most of the steam are
condensed in the fractionator overhead condensers and separated from the
gas in the overhead receiver. The condensed steam with impurities (sour
water) is also separated from the liquid hydrocarbons in this receiver. A
portion of the condensed hydrocarbons is returned to the top tray of the
fractionator as reflux on flow control reset by fractionator overhead
temperature. The gas from the receiver flows to the wet gas compressor
knockout drum in the recovery section. The net hydrocarbon liquid is
pumped to the recovery section on flow control reset by receiver level
control as a part of the lean oil for the primary absorber tower.
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The sour water is pumped on boot level control to the wet gas compressor
first stage discharge as wash water. Corrosion control chemicals injection
into the wash water is provided to prevent corrosion.
The absorber gas from the sponge absorber contains the majority of the
hydrogen sulfide (H2S) resulting from the cracking reaction plus all the
carbon dioxide (CO2) entrained in the regenerated catalyst as inert. These
two acid gases are removed from the absorber gas before it is sent to the
refinery fuel gas pool.
Acid gas removal from the absorber gas is accomplished by contacting the
sour gas with a 25 wt% solution of diethanol-amine (DEA) in a 21 tray
absorption tower.
The sponge absorber sour gas enters the sour gas K.O. drum to separate any
entrained liquid. Condensed/entrained hydrocarbon liquid is routed back to
the fractionator on drum level control. The sour gas then flows into the
amine absorption column where it is contacted with the DEA solution for
H2S and CO2 removal. The lean DEA solution is the feed to top tray of the
absorption column on flow control. The sweet gas leaves the top of the
column, flows through the sweet gas K.O. drum and then is routed to battery
limits on back pressure control. This pressure controller also controls the
operating pressure of the sponge absorber and the absorber/stripper system.
The rich DEA from the amine absorption column is routed to battery limits
on bottom level control, and the condensed liquid from the sweet gas KO
drum is sent to the rich amine flash drum.
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Desulphurisation Section
Reformer Section
Shift Section Gas Purification.
Desulphurization Section :
The catalysts in the reforming and shift sections are extremely sensitive to
sulphur compounds since these will cause deactivation or poisoning. The
MT shift catalyst in the CO-conversion section is sensitive to sulphur and
chlorine compounds. Therefore, the sulphur and chlorine has to be removed
from the feedstock to a very low level before it is sent to the reforming and
shift sections.
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Hydrogen is added to the mix of FCCU and Naphtha, and the total mixture
is preheated to 266C – 380C (dependant of the amount of olefins in the
mix) in Feed Vaporizer, and Feed Preheater.
Hydrogenation
RCl + H2 RH + HCl
RSH + H2 RH + H2S
COS + H2 CO + H2S
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R1 = R2 + H2 HR1 - R2H
This reaction is strongly exothermic, and the exit temperature from the
hydrogenator must be controlled by the outlet temperature from, 703-E-02,
feed preheater.
In order to control the outlet temperature, the inlet temperature of 703-R-01,
is controlled by recycling part of the hydrogenated feed by means of an
ejector, 703-J-01, thus reducing the content of olefin’s in the inlet gas.
For the given feedstock, a hydrogen flow corresponding to 0.2 Nm 3/kg of
naphtha plus 0.05 Nm3/kg of hydrocarbon in the FCCU Off-gas plus
hydrogen required for saturation of olefins is added. This is sufficient for the
hydrogenation and to keep the top layer of the downstream prereformer
catalyst RKNGR in its reduced state.
CO2 + H2 CO + H2O
The hydrogenation catalyst must not get into contact with hydrocarbons,
which have not been mixed with a proper amount of hydrogen. The result
will be a gradual carbon lay-down and a deactivation of the catalyst leading
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The TK-250 catalyst is in its oxidized state when delivered and obtains its
activity when it is reduced. The maximum activity is obtained when it is has
been sulphided. During the initial start-up, the catalyst is pre-sulphided.
CO + 3 H2 CH4 + H2O
In order to maintain the catalyst in its sulphided state, the sulphur content in
the feedstock should preferably not be below 2 ppm for an extended period.
At lower sulphur content the sulphur on the catalyst will start to be stripped
off, and the risk of the above-mentioned hydro-cracking increases.
In the sulphided state after the catalyst has been in operation, the catalyst is
pyrophoric and must not be exposed to air at temperatures above 70oC.
Absorption of Chloride
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The Chlorine Guard Vessel, 703-R-02, is loaded with 6.8 m 3 in one bed of
the Topsøe HTG-1 catalyst in 5 mm rings. The bulk density is around 0.7
kg/l.
It is essential for the absorption of chloride that the organic chlorides are
hydrogenated by the hydrogenation catalyst TK-250 before entering R-01
because HTG-1 is not active towards organic chlorine compounds. The
adsorption curve is very steep, ensuring an extremely low content of
hydrogen chloride in the exit stream.
The catalyst will react with hydrogen chloride according to the following
reactions:
No special attention should be paid to the HTG-1 during start-up, and the
start-up of
HTG-1 will follow the start-up procedure for the HTZ-3 catalyst.
Absorption of Sulphide :
The hydrogen sulphide is absorbed in the third and fourth reactors in this
section, 703-R-03 A and B.
The two reactors, 703-R-03 A/B are located in series and are identical. 703-
R-03 B is acting as guard vessel in case of breakthrough of sulphur from
703-R-03 A or in case 703-R-03 A is taken out of service for replacing the
catalyst.
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Both reactors have two identical beds, each loaded with 12.95 m 3 of Topsøe
HTZ-3 catalyst which consists of activated zinc oxide. The zinc oxide is
delivered as 4 mm diameter extrusions. The bulk density is around 1.3 kg/l.
The zinc oxide reacts with H2S according to the following equation:
During normal operation the sulphur content of the feedstock in contact with
the zinc oxide catalyst is reduced according to the equilibrium constant:
H 2S
K 2.5 x 10 6 at 380 C
H 2O
Reforming Section :
The gas from the desulphurization section is mixed with steam and sent to
the reforming section – consisting of an adiabatic prereformer and a tubular
reformer - where hydrocarbons are reacted over nickel catalysts with steam.
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Reactions (1) and (2) are endothermic while reaction (3) is exothermic, but
the heat required for reactions (1) and (2) will dominate the picture.
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After mixing with steam, the feed + H 2 from the desulphurization section
are preheated to 490oC by heat exchange with the hot flue gas from the
reformer in the prereformer feed preheat coil, the feed plus steam mixture
continues to the prereformer, 703-R-04, where all the higher hydrocarbons
are reformed into hydrogen, carbon monoxide, carbon dioxide and methane
corresponding to the reaction schemes (1), (2) and (3).
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Tubular Reforming :
The tubular reformer is constructed as one radiant chamber divided into five
sections. The furnace contains a single row of 110 centrifugally cast tubes
fabricated of high alloy Cr-Ni steel (25Cr35NiNbTi).
The 110 tubes contain a total of 13.2 m3 R-67-7H reforming catalyst, size 20
× 18 mm with seven holes.
The process gas flows downwards with the gas entering the top of the
vertically mounted tubes from a header through "hairpins" at a temperature
of about 625oC. The gas leaves the tubes at about 930oC and enters directly
into a refractory lined collector.
The tubes in the furnace are heated by a total of 180 dual nozzle burners
arranged in 6 rows on each side of the furnace to provide easy control of a
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uniform temperature profile along the length of the catalyst tubes. The
burner arrangement also allows for equalization of temperature differences
between tubes, should this become necessary.
The tube arrangement permits easy visual inspection through peep holes of
all tubes during operation.
Flue gas flow is upward with outlet near the top of the radiant chamber. The
flue gas outlet temperature is about 1040oC.
Off-gas from the PSA units and vaporised naphtha are used as fuel for the
reformer.
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Reactions
Hydrotreating Chemistry
Sulfur Removal
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a. Mercaptan
b. Sulfide
c. Disulfide
d. Cyclic Sulfide
e. Thiophenic
Nitrogen Removal
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a. Pyridine
b. Quinoline
c. Pyrrole
Oxygen Removal
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a. Phenols
Olefin Saturation
Olefin saturation reactions proceed very rapidly and have a high heat of
reaction.
a. Linear Olefin
b.Cyclic Olefin
Aromatic Saturation
Aromatic saturation reactions are the most difficult. The reactions are
influenced by process conditions and are often equilibrium limited. Unit
design parameters would consider the desired degree of saturation for each
specific unit. The saturation reaction is very exothermic.
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Metals Removal
The useful life of the catalyst may be determined by the amount of metals
that are accumulated on it during the course of operation. Most Unionfining
units are able to go through several operating cycles without exceeding the
ability of the catalyst for removing metals. Metal removal is essentially
complete above temperatures of 600℉ to a metals loading of 2-3 wt% of
the total catalyst. Above this level, the catalyst begins to approach
equilibrium saturation and metals breakthrough is likely.
The total metals retention of the catalyst system can be increased by using a
guard reactor or a guard bed of catalyst specifically designed for
demetallization. Some demetallization catalysts may retain as much as 100
wt% metals based on fresh catalyst weight. Such catalysts typically have a
lower activity for desulfurization and denitrogenation.
Halides Removal
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Reaction Rates
Desulfurization 1
Olefin Saturation 2
Denitrification 1
Aromatics Saturation 1
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Hydrocracking Reactions
The products of the Unionfining reactions are of a lower density than the
feedstock. Therefore, the total liquid yield will be greater than the feed and
may be as high as 102 lv%. Some hydrocracking may take place in the
Unionfining process. This is especially evident toward the end of an
operating cycle when reactor temperatures are raised to compensate for
lower catalyst activity. Total liquid yield and hydrogen consumption will
increase as hydrocracking retractions proceed. Most of the increase in liquid
volume yield will come from more net stripper overheads at the expense of
lower stripper bottoms product. The Unionfining unit is designed for
maximum bottom production. Economic considerations will determine the
amount of bottoms product that can be lost before the unit is shutdown for
regeneration or catalyst change. Under normal operation, the net stripper
overhead liquid produced in a Unionfining unit should not exceed 2 lv% of
the unit feed.
Hydrogen Consumption
Catalyst
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Catalyst Details
The exact arrangement of lines, vessels, etc. may vary from unit to unit, but
basically all units will consist of a reaction section and a fractionation
section. These two sections are described below in general terms.
REACTION SECTION
The feed to the unit can be either cold (40deg C) and hot (100deg C). Feed
obtained from offsite storage tanks at 40deg C is pumped into a Feed
coalescer (702-V-01) for the removal of potential free water. The feed is
further heated to 100deg c in a preheater exchanger (702-E-01). Hot feed is
sent from offsite and pumped by Hot feed pumps (702-P-08A/B). The
combined feed is sent through a Feed Filter (702-G-01) for removal of
suspended solids to the feed surge drum (702-V-02) which is blanked with
nitrogen to prevent gum formation resulting in possible equipment fouling.
The feed pumps (702-P02A/B) take suction from the feed surge drum (702-
V-02) and pumps the raw oil to reactor loops and the feed is preheated via
process exchangers (702-E-06 and 702-E-03) with reactor effluent.
The feed pumps are high head machines capable of pumping large volumes
of oil at pressures of over 120 kg/cm2. The manufacturer’s instructions must
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be consulted before operating the charge pumps since special care must be
taken to avoid damage due to low flow, high temperatures, vibration, etc.
Proper lubrication and cooling must be assured at all times both for the pump
and its driver if serious damage is to be avoided. This type of pump should
never be operated against a blocked discharge, nor at flow rates below the
minimum recommended by the manufacturer. A spillback to the surge drum
may be added to maintain minimum flow at reduce throughput.
Austenitic stainless steel materials are normally used in the hottest heat
exchangers.
The austenitic stainless steel in DHDT are
1) E08 A/B
2) E07
3) E06 A1/B 2
4) E05
5) E04 A/B
6) E03 A1/2/3 & B1/2/3
7) E02
8) F01
These materials provide the best resistance to the corrosive atmosphere and
severe operating conditions. However, they are subject to stress corrosion
when exposed to air and moisture. This type of corrosion can be avoided by
neutralizing the sulfide scale on the tube walls and by avoiding the
condensation of moisture in the tubes. Protection of austenitic steel
equipment is described in detail in the Special Procedures section.
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After separation of the gas and liquid phases in the high pressure separator
(702-V-04), the gas leaves from the top of the high pressure separator and
flows to the suction of the recycle gas compressor via Recycle gas K.O drum
(702-C-01). The recycle gas will be sent to an amine scrubber to remove H2S
in the future.
Lube oil circulation should also be maintained when the compressor is down,
and only shut off when required for maintenance.
After the recycle compressor discharge, some recycle gas will be split off the
main stream for use as quench gas between catalyst beds in the reactor.
Separate quench gas streams are used to reduce reactor interbed temperatures
before each catalyst bed. Quench flow is regulated by a flow controller
cascaded from a temperature controller at the top of the catalyst bed below
the quench zone.
The makeup gas joins the recycle gas before the recycle gas compressor in
order for a makeup stage of compression to be saved. The combined makeup
and recycle gas is divided into passes which are normally allow controlled
into the combined feed passes going to the combined feed exchanger. The
object is to maintain equal gas flow to each reactor charge heater pass at a
sufficiently high rate to avoid overheating the tubes. From this point until it
returns to the high pressure separator, the gas flows along with the liquid
through the reactor circuit in the same manner previously described.
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Reactor
Once the feed and recycle gas have been heated to the desired temperature,
the reactants enter the top of the reactor. As the reactants flow downward
through the catalyst bed, various exothermic chemical reactions occur and the
temperature increases. Each bed will contain a 3 element radial thermocouple
assembly at the top and at periodic levels down through the bed dependent
upon bed length.
Multiple catalyst beds are provided depending upon the heat of reaction and
unit capacity. Reactor skin thermocouples will be provided at the bottom of
each bed and on the bottom reactor head, for monitoring the reactor wall
temperature.
Specific reactor designs will depend upon several variables. Reactor diameter
is typically set by the cross-sectional liquid flux. As the unit capacity
increases, the reactor diameter increases to the point where two parallel trains
would be considered. Reactor height is a function of the amount of catalyst
and number of beds required. Other local factors may also influence the
reactor design including seismic activity and weight limitations. Crane,
bridges and road capacities are also factors.
The reactors are typically divided into individual catalyst beds supported on a
beam and grid support system. The support system is separated from the next
bed of catalyst by a quench gas distributor, reactant mixing internals and a
vapor/liquid re-distributor tray. The reasons for separating a reactor into
separate beds are the following:
a. If the gas and liquid flows become poorly distributed part way through
the reactor, the catalyst will not be effectively utilized, By separating a
reactor into multiple beds with redistribution trays in between, the reactants
existing one bed are redistributed evenly across the cross-sectional area of the
next catalyst bed. In this way, should there be a problem with distribution in a
bed, the catalyst in the lower beds will still be effectively utilized.
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c. In certain situations the heat of reaction will be large enough that the
temperature increase across a reactor will be greater than the design. If this
were allowed to happen, a reaction could become unstable and result in a
temperature runaway. Therefore, cold recycle gas about 170-150deg C (40-
65deg C) is brought into the reactor at the interbed quench points in order to
cool the reactants and thus control the reaction rate.
Good distribution of reactants at the reactor inlet and at the top of each
subsequent catalyst bed is essential for optimum catalyst performance. UOP's
proprietary reactor internals are used to accomplish this distribution.
Due to the exothermic nature of the reactions taking place in the reactor, the
temperature of the material leaving will be greater than the reactor inlet
temperature. The heat of reaction as well as a large portion of the heat
contained in the reactor feed are recovered in a series of heat exchangers. The
reactor effluent is used to preheat not only the liquid feed but also recycle gas
and stripper feed.
This unit has the HP separator (702-V-04) and the Flash drum (702-V-05).
The HP separator operates at 84kg/cm2g and 54 0C where three phases are
separated. The hyrdrocarbon liquid phase from HP separator is routed under
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level control to the Flash drum through Power recovery turbine (702-PRT-02)
to recover the energy. The sour water containing salts is also routed to the
bottom of the flash drum. The gases are further compressed in recycle
compressor (702-K-01) and recycled.
The Flash drum operates at 17.6kgcm2g and 540C where three phases are
again separated. The sour water is removed under level control and sent to
sour water stripping unit at battery limit. This water contains a large
concentration of highly toxic H2S and NH3, and should not be drained to the
sewer for any length of time. The hydrocarbon liquid phase is routed stripper
heat exchanger train (702-E-08A/B and 702-E-04A/B) where it is heated by
exchangers with reactor effluent to the required stripper inlet temperature of
260℃.
Fractionation Section
The function of the fractionation section is to separate sour gas and naphtha
from the diesel product. This can be accomplished with a column
fractionation scheme. The hydrocarbon liquid collected in the Flash drum
(702-V-05) is sent to a stripper column on level control. The feed is preheated
by reactor effluent. Stripping steam (MP steam) is used under flow control to
reboil the stripper. Steam added to the bottom of the tower helps strip light
ends from the bottoms.
Light ends and H2S gather at the top of the stripper and are partially
condensed in
Stripper overhead condenser (702-AC-04) and Stripper overhead trim cooler.
Corrosion inhibitor is injected into the stripper overhead line ahead of the
overhead condenser. Three phases are separated in the stripper overhead
drum (702-V-08). Sour water is combined with the sour water from the Flash
drum and sent to waste water stripping unit at battery limit.
The liquid hydrocarbon are pumped through the stripper reflux pumps (702-
P-07A/B) and spilt into three streams. First one is reflux which is returned to
top of the stripper under level control and second is other reflux routed to the
7th tray of the stripper under the flow control after being heated by the diesel
product – hot naphtha recycle exchanger (702-E-12). And third stream is
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Vapor stream is sent to the absorber knockout drum (702-V-09). This stream
is mixed with gas from the Flash drum before it enters Absorber knockout
drum. The gas from absorber knockout drum is routed to the bottom of the LP
absorber (702-C-04). Lean amine is fed directly from battery limit into the
absorber under flow control. After washing the H2S in the hydrocarbon gas
the amine gets collected in the bottom of the absorber. Rich amine is sent
under level control to the amine treating unit at battery limit. The overhead
gases from the absorber are routed to the Stripper gas amine knockout drum
(702-V-10) to remove the traces of amine in the carryover. The sweet gas is
routed to fuel gas header under the pressure control of the stripper receiver
(702-V-08).
The stripper bottom product exchanges against the returned naphtha reflux to
7th tray of the stripper and undergoes further cooling in the feed preheat
exchanger (702-E-01), the diesel product cooler (702-AC-05) and the diesel
product trim exchanger (702-E-14). Water is removed from the diesel product
in the diesel product coalescer (702-V-11) and the diesel product is sent under
flow control to the storage tank at battery limit.
Process Variables
Standard Operating Parameters:
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DEG C
REACTOR INLET / 102.9/99.8 116.8/114.2 102.1/98.9 102.4/
OUTLET PRESSURE 99.3
KG/CM2
H2 CONSUMPTION 3926 5015 4722 5377
KG/HR
V04 HIGH SEPARATOR 84.9 95 84.9 84.9
KG/CM2
STRIPPER FEED 260 260 260 260
TEMPERATURE DEG C
STEAM 6804 6804 6804 6804
CONSUMPTION
KG/HR
STRIPPER TOP 164 164 164 164
TEMPERATURE DEG C
STRIPPER TOP 7.5 8.5 7.5 7.5
PRESSURE
KG/CM2
REACTOR R1B1 2.5 2.5 2.5 2.5
DELTA P
KG/CM2
REACTOR R1B2 3.5 3.5 3.5 3.5
DELTA P
KG/CM2
REACTOR R2B1 2.5 2.5 2.5 2.5
DELTA P
KG/CM2
REACTOR R2B2 9 9 9 9
DELTA P
KG/CM2
GAS OIL RATIO 1100 1100 1100 1100
NM3/M3
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PURITY
MIN %
REACTOR BED 416 416 416 416
TEMPERATURE
( MAX)
DEG C
The proper operation of the unit will depend on the careful selection and
control of the processing conditions. Most of the variables have already been
mentioned in Section III (technology features) of this manual concerning
catalyst deactivation. However, as they are of utmost importance to the
performance of the unit the following discussion includes some of the steps
that can be taken to maintain them within acceptable limits.
Reactor Temperatures
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The amount of catalyst loaded into the reactors as well as other design
parameters are based on the quantity and quality of feedstock the unit is
designed to process. While minor changes in feed type and charge rate can be
tolerated, wide variations should be avoided since they will tend to reduce the
useful life of the catalyst.
The type of feed being processed is best indicated by its distillation range and
API gravity. An increase in the end point of the feed will make sulfur and
nitrogen removal more difficult, thus requiring higher reactor temperatures
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which, in turn, accelerate coke formation (See Figure IV-4). Coke deposition
is also accelerated by the fact that heavier feed contains more of the
precursors that favor coke formation.
Thus, it is seem that processing higher than design end point feeds will, at
best reduce the length of the operating cycle, and under extreme conditions,
may lead to an irrecoverable loss of catalyst activity. Therefore, every effort
must be made to maintain the end point of the feed within the design limits
by operating the crude and other units such that an acceptable feedstock is
obtained. Storage tanks used for the accumulation of feed to the unit must be
gas blanketed (with nitrogen) or floating roof tanks with torroidal seals in
order to minimize the formation of coke producing and polymer forming
materials that will foul the catalyst bed and Increase pressure drop.
A reduction in the API gravity of the feed for the same boiling range is an
indication of higher unsaturates content. This type of feed will result in
increased hydrogen consumption and higher temperature rise across the
catalyst bed. It also contains more of the materials that easily condense to
form coke in the reactor and associated equipment.
Pressure
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critical design parameter for achieving the desired degree of feed saturation
(See Figures IV-6 and IV-7).
The operating pressure (i.e. the high pressure separator pressure) must be
held constant at the design value for the following reasons:
2) A reduction of the operating pressure below the design level will have
a negative elect on the activity of the catalyst and will accelerate catalyst
deactivation due to coke formation (See Figure IV-8).
The importance of this variable to the satisfactory performance of the unit has
already been described in this manual. It is once again emphasized that the
unit should not be operated at less than design gas-to-oil ratio since rapid
catalyst deactivation will result (See Figure IV-9). The design ratio is set to
provide sufficient hydrogen for the reactions (assuming recycle gas purity is
maintained) and to provide enough heat sink to continuously remove the heat
of reaction.
The quantity of gas circulated from the high-pressure separator to the reactor
relatively large. The gas blow is required for the following purposes:
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1) Provides the excess H2 needed to assure that the reactions are carried
to completion.
Space Velocity
The design quantity of catalyst per unit of feed will depend upon feedstock
Properties, operating Conditions, and product quality required. The liquid
hourly space velocity (LHSV) is defined as follows:
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OPERATION OF CENTRIFUGAL
PUMPS :
Preliminary work :
1) Get a pump technician check if the pump is in good order, if starting
for the first time after maintenance work.
2) Have electricians check the pump motor, the switch and circuit breaker
in the unit
sub-station if starting up for the first time after maintenance work.
3) Check that the motor of the pump is properly earthed to prevent
electrical shock.
4) Ensure that protective covers are placed over exposed parts of the
pump (such as
the pump coupling).
5) Check that the bearing housing has adequate oil level of the required
quality.
6) Rotate with hand the pump shaft and make sure that it rotates freely.
7) Check that cooling water and seal oil systems for the pump are
operable, and there
is adequate flow in each line.
8) Install a proper pressure gauge on the discharge line of the pump.
9) Make sure that all blinds are removed from the necessary suction and
discharge
lines of the pump.
10) Check that all bleeder valves are closed on the suction and discharge
lines of the
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Warming-up a pump
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Start-up of a pump:
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Stopping of pump
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2. Gradually open the inlet valve on the cold side till it is wide open.
4. Start gradually opening the outlet valves on the cold side and
simultaneously gradually close the bypass- on the cold side. When the
outlet valve is wide open close the bypass valve completely.
Having started the flow of the colder fluid proceeds to commission the hot
side. Before commissioning the hot side fluid, it is advisable to recheck that
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no drains or plug points are open. Having confirmed that, commission the
hot side as follows:
2. Gradually open the inlet valve. Simultaneously gradually close the bypass
valve and open gradually the outlet valve. It is very essential to operate the
valves gradually to prevent any thermal shocks.
3. When the inlet and outlet valves are opened completely, shut off the bypass
valve completely.
While commissioning vent any air that may be present so that the heat
exchanger performance is not impaired.
Precaution must be taken while venting to ensure that there is no
flashing of liquid vented.
Flush out tube and/or shell side with gas oils if they ate handling heavy
oils or crude. Pumps H-5A or H-11A can be used for flushing purpose.
2. After the exchanger content cools down, drain the shell and tube
content into E-10 or in the industrial sewer near T-28. Sector 7 should
be informed in case of draining to industrial sewer for effective
draining, the vent or top plug should be opened.
3. Get blinds installed on top and bottom shell flanges (Both tube and
shell sides).
4. In case the tube bundle is to be pulled out, the tube side blinds should
be on the isolation valve flanges i.e. inlet valve D/S flange and outlet
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valve U/S flange. In addition there should be blinds on pump out line
valve (U/S flange) and flushing line valve (D/S flange).
5. Open top and bottom plugs on either side and connect steam hoses.
The hose connected on the bottom plug is to be put in the industrial
drain as to avoid spillage around the exchanger.
Steam the shell/tube side from top till the oil is expelled out.
(This step can be avoided in certain case depending on the fluid handed
and the nature of repairs to be done).
6. Using the same method, flush the tube and shell side with water till
clear water starts coming from the drain.
On prolonged contact with gasoline, the rubber hose may get clogged
and swollen. Checks should be done in this regard.
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MAINTENANCE
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The mechanical deptt. Deals with the machine parts. Here a very close
component various operations are required like turning, boring, slotting,
drilling etc.
For the maintenance purpose various machines are used as listed below:
Lathe machines
Shaper machines
Boring machines
Planners
Drilling machines
Grinding machines
INDUSTRIAL SAFETY
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Preventive measures:
Use of safety guards for reciprocating machine component.
Fencing of dangerous and rotating parts like revolving shafts.
Incorporating safety devices.
Rigid construction of heavy items like crains etc.
Maintenance clearance of shop floor.
Avoiding fire hazard.
Removal of metal chips with proper protection.
Electrical safety :
All metallic parts , externally accessible must be earthed.
Defective and worn out fitting must be replaced promptly.
Inflammable material should not be kept near electric appliances, live
wires or electric control panels.
Safety devices like fuse, circuit breaks, and over- tripping switches must
be used when required.
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REFERENCES
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