IOCL project report(chemical engineering)

Project Report
Indian Oil CorporationLimited
Gujarat Refinery
(A Govt. Of India Undertaking)
Duration: 06/05/2019 to 01/06/2019
Submitted to: Submitted by:
Mr. Vijendra Kumar Ashutosh Choubey (0701CM161008)
AM(MS,L&D)
IOCL
In partial fulfilment of requirements for the degree of
Bachelor’s in Chemical Engineering
UJJAIN ENGINEERING COLLEGE, UJJAIN(M.P.)

INDEX
 Preface
 Acknowledgement
 About IOCL
 Vision
 Refineries
 Pipelines
 Atmospheric Unit
 Fluid Catalytic Cracking Unit
 Diesel Hydrotreater Unit
 Sulphur Recovery Unit
 Efficiency of Furnace
 Heat balance of reactor-regenerator

PREFACE
Industrial training plays a vital role in the progress of future
engineers. Not only does it provide insights about the future concerns,
it also bridges the gap between theory and practical knowledge. We
are fortunate that we were provided with an opportunity of
undergoing industrial training at INDIAN OIL CORPORATION
LTD., Vadodara. The experience gained during this short period was
fascinating to say the least. It was a tremendous feeling to observe the
operation of different units and processes. It was overwhelming for us
to notice how such a big refinery is being monitored and operated
with proper coordination to achieve desired results. During our
training we realised that in order to be a successful chemical engineer
one needs to put his/her concepts into action. Thus, we hope that this
training serves as a stepping stone for us in future and help us to carve
a niche for ourselves in this field.

ACKNOWLEDGMENT
We are highly obliged to Learning and Training Department (Gujarat
Refinery) for providing us this opportunity to intern at IOCL,
Vadodara. Our special thanks to Mr P.K. Verma (SPNM, AU), Mr.
Pankaj Patil (CPNM, FCCU), Mr. A. Chaturvedi (SPNM, DHDT),
Mr. K.N. Kotecha (SPNM, SRU), Mr. Vijendra Kumar (AM (MS,
L&D)). Also, we want to thank Mr. Harsh Khanchandani (PNM,
FCCU) and Mr. K. Banerjee (PNM, FCCU) for helping us with our
project.
We would also like to express our sincere gratitude to Dr.A.K.
Dwivedi (HOD, Chemical Engg. Dept.) and Dr. Hemant Parmar,
(Coordinator, Internship) for allowing us to do our industrial training
at IOCL, Vadodara (Gujarat).

ABOUT IOCL
Indian Oil Corporation is India's largest commercial enterprise, with a
sales turnover of Rs. 438710 crore and profits of Rs. 19,106 crore the
year 2017-18. The improvement in operational and financial
performance for 2017-18 reflected in the market capitalization of the
company, which grew to-fold, from Rs. 95564 crores as on 31st
March 2017 to Rs. 187948 crores as on 31st March 2018. In view of
its rising share price and market capitation, India Oil was included in
the Nifty 50 index (NSE benchmark index of 50 best performing
corporates). Indian Oil is ranked 161st among the world's largest
corporates (and first among Indian enterprises) in the prestigious
“Fortune Global 500” listing for the year 2016.
As India's flagship national oil company, with a 33000-strong work
force currently, Indian Oil has been meeting India's energy demands
for over half a century. With a corporate vision to be ‘The Energy of
India’ and to become ‘A globally admired company’. Indian Oils
business interests straddle the entire hydrocarbon value-chain from
refining pipeline transportation and marketing of petroleum products
to exploration and production of crude oil and gas, marketing of
natural gas and petrochemicals, besides conversion into alternative
energy and globalisation of downstream operations.
Having setup subsidiaries in Sri Lanka, Mauritius and the UAE, the
Corporation is simultaneously scouting for new business opportunities
in the energy markets of Asia and Africa. It has also formed about 20
joint ventures with reputed business partners from India and abroad to
pursue diverse business interests.

VISION
Indian Oil’s ‘Vision with Values’ encompasses the Corporation’s new
aspirations – to broaden its horizons, to expand across new vistas, and
to infuse new-age dynamism among its employees.
Adopted in the company’s Golden Jubilee year (2009), as a ‘shared
vision’ of Indian Oil People and other stakeholders, it is a matrix of
six cornerstones that would together facilitate the Corporation’s
endeavours to be ‘The Energy of India’ and to become ‘A globally
admired company. More importantly, the Vision is infused with the
core values of Care, Innovation, Passion and Trust, which embody the
collective conscience of the company and its people, and have helped
it to grow and achieve new heights of success year after year.

REFINERIES
 Mathura Refinery
The Mathura Refinery, owned by Indian Oil Corporation, is located in
Mathura, Uttar Pradesh. The refinery processes low sulphur crude
from Bombay High, imported low sulphur crude from Nigeria, and
high sulphur crude from the Middle East.
The refinery, which cost Rs.253.92 crores to build, was commissioned
in January, 1982. Construction began on the refinery in October 1972.
The foundation stone was laid by Indira Gandhi, the former Prime
Minister of India. The FCCU and Sulphur Recovery Units were
commissioned in January, 1983.
 Digboi Refinery
The Digboi Refinery was set up at Digboi in 1901 by Assam Oil
Company Ltd. The Indian Oil Corporation Ltd (IOCL) took over the
refinery and marketing management of Assam Oil Company Ltd. with
effect from 1981 and created a separate division. This division has
both refinery and marketing operations. The refinery at Digboi had an
installed capacity of 0.50 MMTPA (million metric tonnes per annum).
The refining capacity of the refinery was increased to 0.65 MMTPA
by modernization of refinery in July, 1996. A new delayed Coking
Unit of 1,70,000 TPA capacity was commissioned in 1999.
 Gujarat Refinery
The Gujarat Refinery is an oil refinery located at Koyali (Near
Vadodara) in Gujarat, Western India. It is the Second largest refinery
owned by Indian Oil Corporation after Panipat Refinery. The refinery
is currently under projected expansion to 18 MMTPA.

 Haldia Refinery
The Haldia Refinery for processing 2.5 MMTPA of Middle East
crude was commissioned in January, 1975 with two sectors-one for
producing fuel products and the other for Lube base stocks.
 Barauni Refinery
Barauni Refinery in the Bihar state of India was built in collaboration
with the Soviet Union at a cost of Rs.49.4 crores and went on stream
in July, 1964.The initial capacity of 1MMTPA was expanded to 3
MMTPA by 1969. A Catalytic Reformer Unit (CRU) was also added
to the refinery in 1997 for production of unleaded motor spirit. The
present capacity of this refinery is 6.1 MMTPA.
 Bongaigaon Refinery
Bongaigaon Refinery is an oil refinery and petrochemical complex
located at Bongaigaon in Assam. It was announced in 1969 and
construction began in 1972.
 Guwahati Refinery
Guwahati Refinery is the country's first Public Sector Refinery as
well as Indian Oil's first Refinery serving the nation since 1962. Built
with Rumanian assistance, the initial crude processing capacity at the
time of commissioning of this Refinery was 0.75 MMTPA and the
Refinery was designed to process indigenous Assam Crude. The
refining capacity was subsequently enhanced to 1.0 MMTPA.

 Paradip Refinery
Paradip refinery is the 11th refinery being set up by Indian Oil
Corporation Ltd. in Paradip town in the state of Odisha. The installed
capacity of refinery was 15 MMTPA.
 Panipat Refinery
The Panipat refinery is the most technically advanced public sector
refinery in India. It supplies petroleum products to the state of
Haryana and the north-west region including Punjab, Chandigarh,
Himachal, Uttaranchal, Jammu & Kashmir, Rajasthan and Delhi.

CRUDE DISTILLATION UNIT
 Short Description of Crude Distillation
In petroleum refining, the Crude Distillation Unit (CDU) (often
referred to as the Atmospheric Distillation Unit) is usually the
first processing equipment through which crude oil is fed. Once in the
CDU, crude oil is distilled into various products, like naphtha,
kerosene, and diesel, that then serve as feedstocks for all other
processing units at the refinery.
As oil is being fed into the CDU, the first thing that happens is the
crude is heated to a temperature between 130 -140°C. This allows
salts, which can be harmful to some equipment, to be removed at the
desalter. The now desalted crude continues through the system into
the heater where it is further heated. Next, it is fed into the
atmospheric column where the vapours and liquids separate. Residues
are stripped out at the bottom of the column. The products (naphtha,
kerosene, diesel and gas oil) are taken from the side of the column
and moved through the refinery for further processing.
 Detail description of Atmospheric unit (AU)
The pre-flash column has 9 trays below and 19 trays above the flash
zone. The overhead vapours are condensed in the condensers and
collected in column. A part of gasoline is sent to the stabilizer while
the rest is returned to preflash column as reflux. The bottom product
is further heated and vaporised in furnace and is introduce into the
flash zone of the main fractionator.
6 stripping trays are provided below the flash zone where the reduced
crude is stripped with steam to remove the lighter constituents. The
stripping section vapours together with the vapourised portion of the
feed are fractionated on the trays above the flash zone to yield liquid
side products and vapour overhead stream. HSD is withdrawn as the
first side stream from the 14th tray along with HSD CR. The product

stream enters the stripper. The HSD and HSD CR are cooled by heat
exchange with crude before it is routed to storage to main fractionator
respectively.
1. Kerosene and Kero CR are withdrawn from the 26th tray. They
are cooled by heat exchange with crude after which the
circulating reflux is returned to fractionator and the product is
cooled and send to storage.
2. The ATF is withdrawn from 34th tray and after exchanging heat
with crude is further cooled and sent to storage.
3. The overhead naphtha vapours are condensed in the condensers
and collected in column.
4. The stripped RCO from bottom of the fractionator exchanges
the heat with crude and are then cooled in coolers before storing
to the storage.
Unstabilized gasoline is sent to stabiliser along with CRU drag
stream. It is preheated by successive heat exchanger with RCO and
bottom product of the stabiliser. the overhead is condensed and a part
is withdrawn as LPG and rest is sent to column as reflux. The reboil
heat is supplied by circulating part of stabilizer through furnace.
Stabilize naphtha from stabiliser bottom after preheating with feed is
routed to column for storage.

 Crude Desalting:
Crude Desalting is a water washing operation performed at the
refinery site to get additional crude oil clean up.
Crude Oil Desalting consists of
 Purifying process.
 Remove salts, inorganic particles and residual water from crude
oil.
 Reduces corrosion and fouling.
Desalting process is used for removal of the salts, like chlorides of
calcium, magnesium and sodium and other impurities as these are
corrosive in nature. The crude oil coming from field separator will
continue to have some water/brine and dirt entrained with it. Water
washing removes much of the water-soluble minerals and entrained
solids (impurities). Desalting process consists of three main stages:
heating, mixing and settling
The desalting process is completed in following steps:
 Dilution water injection and dispersion
 Emulsification of diluted water in oil
 Distribution of the emulsion in the electrostatic field
 Electrostatic coalescence
 Water droplet settling
Crude oil passes through the cold preheat train and is then pumped
to the Desalters by crude charge pumps. The recycled water from
the desalters is injected in the crude oil containing sediments and
produced salty water. This fluid enters in the static mixer which is
a crude/water disperser, maximizing the interfacial surface area for
optimal contact between both liquids.
The wash water shall be injected as near as possible emulsifying
device to avoid a first separation with crude oil. Wash water can
come from various sources including relatively high salt sea water,
stripping water, etc. The static mixers are installed upstream the

emulsifying devices to improve the contact between the salt in the
crude oil and the wash water injected in the line.
The oil/water mixture is homogenously emulsified in the
emulsifying device. The emulsifying device (as a valve) is used to
emulsify the dilution water injected upstream in the oil. The
emulsification is important for contact between the salty
production water contained in the oil and the wash water. Then the
emulsion enters the Desalters where it separates into two phases by
electrostatic coalescence.
The electrostatic coalescence is induced by the polarization effect
resulting from an external electric source. Polarization of water
droplets pulls them out from oil-water emulsion phase. Salt being
dissolved in these water droplets, is also separated along the way.
The produced water is discharged to the water treatment system
(effluent water). It can also be used as wash water for mud washing
process during operation.

Fluid Catalytic Cracking Unit
Fluid catalytic cracking (FCC) is one of the most important
conversion processes used in petroleum refineries. It is widely used to
convert the high-boiling, high-molecular
weight hydrocarbon fractions of petroleum crude oils into more
valuable gasoline, olefinic gases, and other products.
 Process Description:
The Fluidised Catalytic Cracking Unit consist of three section:
1. Catalyst Section
2. Fractionation Section
3. Gas Concentration Section
1. Catalyst section:
The catalytic section consists of regenerator and reactor which
together with standpipes and riser, form the catalytic circulation
system. The catalyst circulates up the Riser to the reactor
comma down through the stripper to the region narrator across
the reason letter standpipe, and back to the riser.
Fresh feed and recycle stream, known as combined feed are
vaporized and heated to the reactor temperature by the hot
catalyst. This mixture of oil paper and catalyst travels up to the
Racer into the reactor. Gas oil commenced to cried immediately
when it contains the hot catalyst and the rise in continuous until
the oil vapour are disengaged from the catalyst in the reactor.
The crack the product are in the vapour phase goes from the
reactor to the fractionating through reactor vapour line.

2. Fractionating Section:
In the fractionating section, the reactor vapour are fractionated
into wet gas, unstabilized gasoline, heaviness After, life cycle,
heavy cycle and clarified oil.
3. Gas Concentration Section:
Compressed wet gas and unstable gasoline go to the gas
concentration unit where they are separated into fuel gas, LPG
and stabilized gasoline. LPG and gasoline are washed with
caustic in the treating section of gas concentration unit. Amine
treatment of LPG is done prior to caustic wash.

Diesel Hydrotreater Unit (DHDT)
 Purpose of the process:
The objective of diesel hydrotreater unit (DHDT) is to produce a low
sulphur product along with cetane improvement.
The gas oil feed to the DHDT unit is a mixer of straight run distillates
and cracked feed stocks; the unit also stabilizes hydrotreater naphtha
from an existing DHDS unit. The cracked feed is comprised of TCO
from the FCC unit and CGO from the Coker unit. The design capacity
is 2.2MMTPA of the feed stock for all cases.
The unit is design for a stream factor of 8000 hours per year and a
turndown ratio of 50% of the unit design capacity.
There are two primary products from the DHDT unit:
• Stabilized diesel which is routed to battery limits for the diesel pool.
• Stabilized naphtha which is sent to battery limits.
In addition, there is a sweet LP of gas stream from the LP Amine
absorbers sent to the fuel gas system and a sweet MP of class stream
from the MP Amin absorber routed for hydrogen recovery. Also,
there is a rich amine stream that is sent to the amine regeneration unit.
 Brief description of the process:
The liquid feed to DHDT unit consists of SRGO, TCO and CGO
streams supplied from the battery limits. This gas oil blend flows to
the feed surge drum through feed filter. The oil is then pumped to the
reactor circuit using the reactor feed pumps.
Make-up hydrogen is supplied to the make-up compressor and is then
compressed by the makeup hydrogen compressor to the reaction
circuit pressure. The makeup and recycle gas streams are combined

with the liquid oil feed , and the total feed mixture is heat exchanged
against the reactor effluent in the reactor feed or effluent exchangers.
The preheated oil and hydrogen mixture is heated to the reactor inlet
temperature in the reactor feed heater. The total feed mixture is then
passed through hydrotreater reactors. The hydrotreater reactors
contain two reactors in series with each containing three catalyst bed.
Recycle hydrogen is used as the quench gas for maintaining the
desired WABT in each of the catalytic beds.
The reactor effluent is cooled by preheating the stripper feed in the
stripper feed preheater then by preheating the oil or hydrogen feed in
the feed or effluent exchangers. Additional cooling is obtained by heat
exchanging the effluent with the stripper field in the reactor effluent
or MP liquid exchanger in order to prevent any deposit of ammonium
salts.
The reactor effluent then is cooled in the reactor effluent air cooler
before being collected in the vertical HP cold separator drum. The
vapour phase from this drum consists of the recycle and quench gas
and is circulated by the recycles compressor. The hydrocarbon or
water separation obtained in this drum is not sharp, and hence both
the hydrocarbon rich and water rich streams are let down and once
again combined in the downstream MP cold separator for recovering
a hydrogen rich off-gas stream. Sour water is separated from the
hydrocarbon liquid phase and is routed to the battery limits. The
hydrogen liquid flows to the stripping section.
The recycle hydrogen stream leaving the HP cold separated drum is
routed to the HP amine absorber for H2S removal. The lean amine is a
35 wt% MDEA solution, supplied to the HP Amine absorber using
the HP lean amine pumps. To prevent the possibility of foaming in the
absorber, the lean amine stream is first heated using LP steam in the
HP amine heater. The rich amine from the HP absorber bottoms is
sent to the LP Amine absorber for flashing of the dissolved right ends.
The recycle gas leaving the HP Amine absorber is then routed to the
recycle compressor through the recycle compressor drum. The H2
compressor discharge is split into recycle gas and quench gas streams

and the recycle gas is combined with the make-up hydrogen and gas
oil feed to enter reactor circuit. The quench gas flows to each reactor
bed.
The diesel stripper fractionates the diesel product from the reactor
effluent liquid stream, which is supplied from the MP cold separator
drum. The wild naphtha product from the stripper overhead is
stabilized and H2S stripped in the stabilizer liquid feed from the MP
separator is first preheated by heat exchange with the reactor effluent
in the reactor effluent or MP liquid exchangers and stripper bottoms
in the stripper feed/bottoms exchanger followed by the final heating
in heat exchange with the reactor effluent in the stripper feed
preheater. The hot feed enters the stripper and is stripped of H2S and
light ends using superheated MP stream to produce wet diesel
product. The stripper overhead vapor is condense in an air-cooled
condenser and the trim condenser. The off-gas product from the
stripper reflux drum is sent to the LP Amine absorber. The water
condensate containing dissolved H2S and traces of NH3 is removed
from the stripper reflux drum water boot. This water is recycled to the
water drum for water injection upstream of the reactor effluent air
cooler. The stripper flux is returned to the tower and wild naphtha
overhead product is sent to the stabilizer.
The stripped wet diesel product is pumped to the battery limits. Heat
is recovered from the diesel product by preheating the stripper feed in
stripper feed or bottoms exchangers. The diesel product is further
cooled in the gas oil air cooler and gas oil trim cooler. The dissolved
water in the diesel product settles out as free water as the product
stream is cooled. The entrained water is removed in the the coaleser
and salt dryers before the diesel product is routed to the battery limits.
The wind naphtha product from the stripper overhead is combined
with the stripper naphtha from the DHDS unit, and the total naphtha
feed stabilized and H2S stripped in the stabilizer. This stabilizer feed
is preheated with the stabilizer bottom stream in the stabilizer
feed/bottom exchanger, and enter the stabilizer. The stabilizer is
reboiled using the hot diesel product from the stripper bottoms.

Stabilized naphtha product from the column is cooled by the cooling
water and send to storage.
The stabilizer overhead vapour is condensed with cooling water in
condenser. Off-gas product from the stabilizer reflux drum containing
H2S, butanes and lighter components, is sent for amine scrubbing to
the LP amine absorber. The entire liquid product from the stabilizer
reflux drum is returned to the stabilizer as a reflux stream.
Amine scrubbing in the unit is accomplished in three amine towers,
the HP amine absorber, the MP amine absorber, and the LP amine
absorbers. Lean amine (35 wt% MDEA) is supplied at the battery
limits at a pressure sufficient to be routed directly to the LP absorber.
The MP absorber scrubs the off-gas from the diesel stripper and
stabilizer as well as any dissolved light end from the HP and the MP
rich amine streams, which are laid down and filled to the LP absorber
sump.
To avoid foaming on the top tray(beds), the lean amine for the MP
absorber is heated along with the lean amine for the HP absorber in
the common exchanger using LP steam. Similarly, LP steam heats the
lean for the LP absorber in the LP amine heater.
The sour gas in the route to the LP and the MP absorber through KO
drums. The sweet of gas product from the LP absorber is routed to the
fuel gas mixture change system as well as provision for alternative
routing to the FCC unit. The off-gas product from the MP absorber is
routed to PSA II in HGU unit as well as a provision for alternative
routing to the fuel gas tank mixing system. Rich amine from the LP
absorber bottom is pumped into the battery limits.
 Thermodynamics and Kinetics:
The sulphur compound (mercaptans, sulphides, disulphides) are easily
hydrogenated to saturated or aromatic corresponding hydrocarbons
with product ions of H2S. The nitrogen compound is hydrogenated to
saturated or aromatic hydrocarbons with production of ammonia. This

reaction required higher H2 partial pressure and are exothermic. These
reactions are favoured by low temperature and high pressure.
Thermodynamics also show that in opposition to aromatics, the
hydrogenation of diolefins are almost complete at 3000C with the low
hydrogen partial pressure.
The operating conditions result from a compromise between the
above thermodynamic consideration and the minimum temperature
required by kinetics of reaction.
 Chemical reactions
The chemical reactions in hydrogenation process are of two types
desired and undesired.
The desirable reaction includes elimination of S, N, O, metals,
saturation of olefins and diolefins and saturation of aromatics.
The undesirable reaction is the reaction which results in a loss of
valuable components of the feed or decrease of catalyst activity.
1. Desirable reaction
a. Desulfurization reaction
Mercaptans, sulphides and disulphides react easily leading to the
corresponding saturated or aromatic compounds. Sulphur combine
into cycles of aromatic structures like thiophene is more difficult to
eliminate. This reaction is exothermic and produce hydrogen sulphide
and consume hydrogen.
Examples:
Mercaptans
R-SH + H2 R-H +H2S
Sulphide
R-S-R +2H2 2R-H + H2S

b. Denitrogenation reaction
This reaction lead to ammonia formation they are highly exothermic.
R-NH2 +H2  RH + NH3
c. Hydrogenation of oxygenated compound
Hydrogenation of bond C-O alcohol and phenols.
R-OH + H2  R-H + H2O
d. Hydrogenation of aldehydes
R-CHO + 2H2  R-CH3 + H2O
e. Hydrogenation of olefin compound
This reaction is highly exothermic. Olefins and diolefins are
converted to saturated compound.
f. Hydrogenation of aromatic compound
Generally, this reaction is exothermic and are favoured by low
temperature and high pressure. For a given pressure when the
temperature increases the hydrogenation rate increases first to reach a
maximum and then decreases as the temperature continue to increase
and it increases rapidly with the pressure.
g. Demetalization
The organometallic compounds (Pb, Cu, Ni, Va) are cracked and the
metals are trapped on the catalyst.

2. Undesirable reaction
The maximum productivity is achieved by limiting the undesirable
reaction.
a. Hydrocracking
It is an undesirable reaction because it consumes H2 , reduces the
product yield and the hydrogen purity of the recycle gas. It is limited
by the selection of catalyst with low hydrocracking capacity working
at low temperature. Hydrocracking increases with the temperature.
Example:
R-CH2-CH2-R’ +H2R-CH3 +R’-CH3
b. Coking
Under the design operating conditions heavy molecules which are
observed on the acidic sites of the catalyst may be condensed and
progressively polymerize on the catalyst and form coke. This coke
deposition is the main cause of catalyst activity reduction.
3. Amine H2S absorption
H2S or HSH is a weak acid and ionizes in H2O to form H2 ion and
sulphide ion.
H2S + H2O  H3O+
+ HS-
Since it is fairly weak acid only a fraction of the H2S will ionize
methyl diethanol amine is a weak base and ionizes in water to form
amine ion and hydroxyl ion.
When H2S dissolves in the solution containing the amine ion it reacts
to form a weakly bonded salt of the acid and base.
The sulphide ion is thus absorbed by the amine solution. This salt
formation reaction is not proceeded to completion, as the arrows
indicate an equilibrium level of H2S remains in the hydrocarbon
stream. The overall reaction can be summarised by the following
equation.

SULPHUR RECOVERY UNIT (SRU)
H2S removed in the ARU and SWS unit is sent to the sulphur
recovery unit (SRU) as acid gas. SRU recovers H2S as elemental
sulphur through the Claus reaction. Reactions occur in two stages: the
flame reaction stage and the catalytic reaction stage. The sulphur
recovery rate of the Claus process is about 95 to 97%.
 Amine Recovery Unit (ARU)
Rich MDEA (Methyl di-ethanol amine) solvent is collected from
amine absorbers in DCU, HGU, DHDT, VGO and ISOM units to rich
solvent flash drums. This rich solvent contains H2S with some CO2
and hydrocarbons. The rich solvent is flashed to remove light
hydrocarbons. The flash drum is designed to remove heavier
hydrocarbons through a skimming layer.
The rich solvent is pumped through a lean/rich solvent exchanger
where the rich solvent is heated by the lean solvent stream. The rich
solvent then flows to the free tray of the solvent regenerator stripping
vapours from the regenerator reboiler strips H2S from the solvent. The
H2S rich gas stream from the overhead of the regenerator is sent to the
SRU for conversion to elemental sulphur.
The lean solvent stripped of H2S, is pumped from the bottom of the
regenerator and is first cooled by the lean/rich exchanger, and further
cooled by an air cooler and a trim cooler. The lean solvent is filtered
and stored in a surge tank before being sent back to the anime
absorber in DCU, HCU, DHDT, VGO, FCC LPG MEROX and
ISOM unit.

 Chemistry of amine regeneration:
Primary reactions of MDEA with H2S are as follows:
2R2NH + H2S  (R2NH2)2S
(R2NH2)2S +H2S 2R2NH2HS
The reactions shown above proceed to the left at higher temperature.
At elevated temperature as exist in the regenerator column, the
reactions are reversed with sulphur salts being decomposed and acid
gas released.
 Sour water stripping:
The sour water as a feed received from DCU, HGU and SRU are
collected in refinery sour water surge drum after cooling up to 400C in
refinery feed sour water cooler. Hydrocarbons are flashed in surge
drum along with small quantity of ammonia and hydrogen sulphide.
Flashed vapours are routed to SRU. The entrained hydrocarbon liquid
is separated through partition baffle and drain to OWS. The
hydrocarbon free sour water is then routed to refinery sour water
storage tank.
The sour water is then pumped through refinery feed/ bottom
exchanger where sour water is heated by stripped water. After then
sour water flow to the feed tray of refinery sour water stripper,
stripping vapours from the refinery stripper reboiler strips ammonia
and hydrogen sulphide from sour water. The ammonia and hydrogen
sulphide rich overhead vapours are routed to SRU for the conversion
to elemental sulphur.
The stripped water from refineries sour water stripper bottom is
pumped through refinery stripped water pump and is first cooled by
refinery feed/bottom exchanger and further cool by refinery stripped
water cooler.

 Claus sulphur recovery
 Reaction furnace
The front end of the process uses the modified Claus process. In this
one third of the hydrogen sulphide in the acid gas streams is bond to
form a SO2. The SO2 then reacts with the balance of hydrogen
sulphide to form elemental sulphur and water in vapour phase.
Ammonia present in sour water SWS acid gas is destroyed during the
combustion process.
H2S +3/2O2  SO2 + H2O + Heat
2H2S + SO2 2H2O + 3/n Sn + Heat
2NH3 + 3/2O2 3H2O + N2
Other forms of sulphur are
3S2 S6 + Heat
4S2 S8 + Heat
Other combustibles in the gas will burn along with the H2S forming
other combustion products. Also, H2S will dissociate at high
temperature forming hydrogen and elemental sulphur.
CH4 + 2O2  CO2 + 2H2O
CO2 + 2H2S  COS + H2O
COS + H2S  CS2 + H2O
2H2S +Heat  S2 + H2
COS + H2O  H2S + CO2
The combustion reaction is carried out in the burner and combustion
chamber of the reaction furnace. The hot gases are cooled in the
waste heat exchanger and subsequently in sulphur condenser where
steam is generated by removing the heat of reaction. Sulphur vapour
is condensed and removed in the sulphur condenser.

 Ammonia description
At higher ammonia concentration, serious problems maybe
encountered caused by formation of solid nitrogen, sulphur salts such
as ammonium sulphate and ammonium hydrosulphate. They tend to
deposit in and plug catalytic reactor, condenser etc.
For such ammonia content in feed the preferred and proven solution is
the use of two zone reaction furnace. In two zone reaction furnace all
the combustion air is supplied to combustion chamber with a portion
of acid gas and all the sour water stripper acid gas. The quantity of
amine acid gas to the combustion chamber is restricted to control the
elevated temperature. This is done to ensure complete destruction of
the ammonia. The remainder of the amine acid gas is feed to the
reaction chamber of the reaction furnace. The temperature in the
combustion chamber should be maintained at 2370- 2680°F.
The moles of H2S is fed to combustion chamber must always be
greater than one third of the total H2S in all feed streams to prevent
SO3 formation and to ensure total O2 combustion. The minimum
recommended operating range for total H2S to combustion chamber is
40% of the total in the combined feeds.
The combustion gases from the burner and combustion chamber flow
into the reaction chamber of the reaction furnace where adequate
residence time is provided for the sulphur reaction and ammonia
destruction. The reaction furnace zone volumes are based on
ammonia destruction kinetics. The residence time and reactor
configuration are optimised to ensure adequate ammonia destruction
and all levels of expected operation. This include ammonia and
hydrocarbon destruction. The hot gases flow from the reaction
chamber to the waste heat exchanger.
 Catalytic Conversion Cycles:
The conversion reaction of the remaining Sulphur dioxide and
hydrogen sulphide occurs at lower temperature in catalytic converters
with an accompanying temperature is rise liquid sulphur is condensed

and removed from the vapour stream first downstream of the furnace
and then after each catalytic converter in order to reduce the
concentration of sulphur in the vapour favouring the continuous of
reaction.
The amount of sulphur recovered from catalytic converter decreases
because of the reduction in concentration of SO2 and H2S in the
vapour stream. The conversion reaction and the catalytic converter
improves as the inlet temperature of is lowered but converter
temperature should be above sulphur dew point temperature. The
recovery of liquid sulphur from each converter stage is favoured by a
low condenser outlet temperature.
 COS / CS2 Formation and Destruction:
In the combustion of the H2S in reaction furnace small amount of CS2
and carbonyl sulphide are formed from carbon dioxide and any
hydrocarbons present in feed gas. These sulphur compounds are
resistant to conversion to elemental sulphur at lower temperature.
However, these components can be recovered as sulphur with proper
design.
 Tail Gas Unit:
The tail gas unit remove the remaining sulphur from the claus unit tail
gas by a combination of chemical reaction and absorption. The
process involves:
1. Converting all the remaining sulphur in claus tail gas to H2S.
2. Selectively absorbing the H2S from the rest of the tail gas
constituents.
3. Stripping the absorbed H2S from the solvent and returning it to
the front of the claus unit for subsequent sulphur recovery.
 TGU reaction system
Claus tail gas from outlet of final sulphur condenser flows to the TGU
reaction system for reduction of sulphur compounds to H2S.

Reduction is achieved by first heating the tail gas, adding reducing
hydrogen and passing the hot gas over a sulphide proprietary catalyst.
An indirect superheated HP steam heater for tail gas heating is used.
Tail gas enters the preheater at 2700Fand exits 5500F. The hot gas mix
passes through a catalyst bed in hydrogenation reactor where sulphur
dioxide and other compounds are converted to H2S exothermically.
Carbon monoxide is equivalent to hydrogen as reducing agent with
somewhat higher reactivity.
The hot reactor effluents is cooled before going to absorber. This is
done by steam production in TGU waste heat exchanger and further
temperature reduction in quench column where the bulk of water of
combustion is removed from the process stream.
 TGU amine system:
Cooled water effluent enters the absorber/regenerator system where
MDEA is circulated to remove the hydrogen sulphide. The use of
selective amine is required in the TGU process to prevent the build-up
of inert within the claus loop. If the inert were not effectively
removed the entire process would be chocked off. The H2S is then
stripped from the TGU solvent in the solvent regenerator and the
regenerator overhead is recycled back to the claus plant where the
absorber overhead, or TGU tail gas, is incinerated. Sour water is
removed from the quench column water circuit under flow control to
maintain the water balance in the unit. The main overall reaction
occurring in the absorber are:
H2S + R2CH3N  R2CH3NH+ + HS- ----------------------(k1)
CO2 + R2CH3N + H2O  R2CH3NH+ + HCO3
- -------------------(k2)
Where k1>>k2 and R is C2H4OH
The H2S reach much faster with the solvent then does CO2 is
kinetically limited. The absorption of H2S and the selectivity of the
H2S over CO2 are enhanced at lower operating temperature. The
selective absorption of this H2S over CO2 will be optimised when the

temperature rise of the solvent across the observer is at a minimum.
High solvent flow rate will cause both H2S and CO2 to be absorbed.
 Tail gas Incineration (TGU)
Tail gas from TGU flows to the incinerator where residual sulphur is
converted to SO2 for discharge. It is accomplished at high temperature
by combustion of fuel gas with excess air. The exchanger in this unit
has the dual function of generating high pressure steam and
superheating high pressure steam.
 Sulphur handling:
Internal steam coils are used to keep the sulphur in the molten state.
The liquid sulphur is collected in the sulphur storage tank. Typically,
claus sulphur in the pit contains 250 - 300 ppm by wt. of H2S. More
H2S appears to be present at higher temperature than lower
temperature.
The key points of H2S behaviour and liquid sulphur can be
summarised as follows:
1. H2S exist in claus sulphur primarily as hydrogen polysulfide.
2. As the typical sulphur pit temperature is reached to hydrogen
polysulfide will decompose to dissolve H2S.
3. Dissolved H2S will then evolve as gaseous H2S.
4. It has been observed that the evolution of gaseous H2S will
increase with agitation of the liquid sulphur.

Fluidised Catalytic Cracking Unit
Fluid catalytic cracking (FCC) is one of the most important
conversion processes used in petroleum refineries. It is widely used to
convert the high-boiling, high-molecular
weight hydrocarbon fractions of petroleum crude oils into more
valuable gasoline, olefinic gases, and other products. Cracking of
petroleum hydrocarbons was originally done by thermal cracking,
which has been almost completely replaced by catalytic cracking
because it produces more gasoline with a higher octane rating. It also
produces by-product gases that have more carbon-carbon double
bonds (i.e. more olefins), and hence more economic value, than those
produced by thermal cracking.
The feedstock to FCC is usually that portion of the crude oil that has
an initial boiling point of 340 °C or higher at atmospheric
pressure and an average molecular weight ranging from about 200 to
600 or higher. This portion of crude oil is often referred to as heavy
gas oil or vacuum gas oil (HVGO). In the FCC process, the feedstock
is heated to a high temperature and moderate pressure, and brought
into contact with a hot, powdered catalyst. The catalyst breaks the
long-chain molecules of the high-boiling hydrocarbon liquids into
much shorter molecules, which are collected as a vapour.

 Detail process description of the following section:
Process flow of the entire unit can be divided into sections
1. Feed preheater section
2. Catalyst section
3. Fractionation section
4. Gas concentration section
5. Caustic treatment section
6. Steam generation section
1. Feed Preheater Section:
The purpose of the feed preheater is to recover the heat from the
furnace flue gas which increases the efficiency of the furnace by
reducing the useful heat lost in the flue gas. As a consequence, the
flue gases are also conveyed to the flue gas stack (or chimney) at a
lower temperature, allowing simplified design of the conveyance
system and the flue gas stack. It also allows control over the
temperature of gases leaving the stack.
The reactor-regenerator is the heart of the FCC process. In a modern
cat cracker, virtually all the reactions occur in 1.5 to 3.0 seconds
before the catalyst and the products are separated in the reactor. From
the preheater, the feed enters the riser near the base where it contacts
the regenerated catalyst .The heat absorbed by the catalyst in the
regenerator provides the energy to heat the feed to its desired reactor
temperature. The heat of the reaction occurring in the riser is
endothermic (i.e., it requires energy input). The circulating catalyst
provides this energy.

2. Catalytic Section:
Modern FCC catalysts are fine powders with a bulk density of 0.80 to
0.96 g/cm3 and having a particle size distribution ranging from 10 to
150 µm and an average particle size of 60 to 100 μm. The design and
operation of an FCC unit is largely dependent upon the chemical and
physical properties of the catalyst. The desirable properties of an FCC
catalyst are:
 Good stability to high temperature and to steam
 High activity
 Large pore sizes
 Good resistance to attrition
 Low coke production
A modern FCC catalyst has four major components:
crystalline zeolite, matrix, binder, and filler. Zeolite is the active
component and can comprise from about 15 to 50 weight percent of
the catalyst.
The preheated high-boiling petroleum feedstock (at about 315 to
430 °C) consisting of long-chain hydrocarbon molecules is combined
with recycle slurry oil from the bottom of the distillation column and

injected into the catalyst riser where it is vaporized and cracked into
smaller molecules of vapor by contact and mixing with the very hot
powdered catalyst from the regenerator. All of the cracking reactions
take place in the catalyst riser within a period of 2–4 seconds. The
hydrocarbon vapours "fluidize" the powdered catalyst and the mixture
of hydrocarbon vapours and catalyst flows upward to enter the
reactor.
The reactor is a vessel in which the cracked product vapours are:
a. separated from the spent catalyst by flowing through a set of
two-stage cyclones within the reactor
b. the spent catalyst flows downward through a steam stripping
section to remove any hydrocarbon vapour before the spent
catalyst returns to the catalyst regenerator. The flow of spent
catalyst to the regenerator is regulated by a slide valve in the
spent catalyst line.
Since the cracking reactions produce some carbonaceous material
(referred to as catalyst coke) that deposits on the catalyst and very
quickly reduces the catalyst reactivity, the catalyst is regenerated by
burning off the deposited coke with air blown into the regenerator.
The combustion of the coke is exothermic and it produces a large
amount of heat that is partially absorbed by the regenerated catalyst
and provides the heat required for the vaporization of the feedstock
and the endothermic cracking reactions that take place in the catalyst
riser. For that reason, FCC units are often referred to as being 'heat
balanced'.
The hot catalyst leaving the regenerator flows into a catalyst
withdrawal well where any entrained combustion flue gases are
allowed to escape and flow back into the upper part to the regenerator.
The flow of regenerated catalyst to the feedstock injection point
below the catalyst riser is regulated by a slide valve in the regenerated
catalyst line. The hot flue gas exits the regenerator after passing
through multiple sets of two-stage cyclones that remove entrained
catalyst from the flue gas.

3. Fractionation section:
The purpose of the fractionator is to desuperheater and recover liquid
products from the reactor vapour. The hot product vapours from the
reactor flow into the fractionator near the base. Fractionation is
accomplished by condensing and revalorising hydrocarbon
components as the vapor flows upward through trays in the tower.
The operation of the main column is similar to a crude tower, but with
two differences. First, the reactor effluent vapours must be cooled
before any fractionation begins. Second, large quantities of gases will
travel overhead with the unstabilized gasoline for further separation.
In Fractionation components like flue gas, LPG, gasoline, heavy
naphtha, Light cycle oil (LCO) and clarified oil is obtained.

Unstabilized gasoline and light gases pass up through the main
column and leave as vapor. The overhead vapor is cooled and
partially condensed in the fractionator overhead condensers.
Hydrocarbon vapor, hydrocarbon liquid, and water are separated in
the drum.
The heaviest bottoms product from the main column is commonly
called slurry or decant oil. Above the bottom product, the main
column is often designed for three possible side cuts:
 Heavy cycle oil (HCO)—used as a pump around stream,
sometimes as recycle to the riser, but rarely as a product
 Light cycle oil (LCO)—used as a pump around stream,
sometimes as absorption oil in the gas plant, and stripped as a
product for diesel blending
 Heavy naphtha—used as a pump around stream, sometimes as
absorption oil in the gas plant, and possible blending in the
gasoline pool.

4. Gas Concentration Section:
The FCC gas plant separates the unstabilized gasoline
and light gases into the following:
 Fuel gas
 C3 and C4 compounds
 Gasoline
C3's and C4's include propane, propylene, normal butane, isobutane,
and butylene. Propylene and butylene are used to make ethers and
alkylate, which are blended to produce high-octane gasoline. LPG and
gasoline are washed with caustic in the treating section of gas
concentration unit. Amine treatment of LPG is done in the caustic
treatment section. Amine and caustic solutions are used to remove
these impurities. The amine solvents known as alkanolamines remove
both H2S and CO2. Hydrogen sulphide is poisonous and toxic.
Amines remove the bulk of the H2S, primary amines also removethe
CO2. An amine absorber removes the bulk of H2S from the sour gas.
The sour gas leaving the sponge oil absorber usually flows into a
separator that removes and liquefies hydrocarbon from vapour. The
gas from the separator flows to the bottom of the H2S contactor where
it contacts a counter current flow of the cooled lean amine from the
regenerator. The treated fuel gas leaves the top of the H2S absorber,
goes to a settler drum for the removal of entrained solvent, and then
flows to the fuel system. Rich amine from the bottom of the H2S
contactor goes to a flash separator to remove dissolved hydrocarbons
from the amine solution. The rich amine is pumped from the separator
to the amine regenerator.
5. Caustic Treatment section:
Mercaptans are organic sulphur compounds having the general
formula of R-S-H. As stated earlier, amine treating is not effective for
the removal of mercaptan. There are two options for treating
mercaptans. In each option, the mercaptans are first oxidized to

disulphides. One option, extraction, dissolves the disulphides in
caustic and removes them. If the LPG and the gasoline contain high
levels of H2S, a caustic prewash is needed to protect the catalyst.
Chemical Reactions:
RSH + NaOH +catalystRSNa + H2O
2RSNa + ½ O2 + H2O + catalyst RSSR +2NaOH
6. Steam Generation Section:
A boiler or steam generator is used wherever a source of steam is
required. HCO goes to the steam generator. Most of the fluid catalytic
cracking (FCC) units produce gases rich in carbon monoxide (CO).
This gas is burnt in a CO boiler to recover heat and produce steam. Re
The supply of CO gas is normally not sufficient to ensure destruction
of CO. Supplementary fuel raises the temperature of the CO gases to
the ignition point and assures the complete burning of the
combustibles in the CO gas stream. There are two types of CO boilers
prevalent in the industry. One design has a separate combustion
chamber for CO burning and is followed by a convection section for
heat recovery. The other type is a water wall type boiler which is
modified to burn CO gas.
The exhaust emissions from the CO boiler must be controlled reliably
and predictably to meet operating company requirements. Principal
constituents for control in the exhaust emissions include carbon
monoxide (CO), oxides of nitrogen (NOx), sulphur oxides (SOx) and
particulate matter. The design of UOP Callidus CO combustors
ensures complete conversion of CO to CO2 while minimising NOx
emissions with post-combustion treatment of sulphur and particulates.

PROJECT OBJECTIVE
 To calculate the efficiency of the furnace used in the Fluidised
Catalytic Cracking Unit.
Theory: the efficiency of a furnace is the ratio of useful output to heat
input. The furnace efficiency is calculated by:
1. Direct Method
2. Indirect Method
1. Direct Method: The efficiency of a furnace can be computed by
measuring the amount of fuel consumed per unit weight of
material produced from the furnace.
Efficiency of the furnace =
𝐇𝐞𝐚𝐭 𝐢𝐧 𝐭𝐡𝐞 𝐬𝐭𝐨𝐜𝐤
𝐇𝐞𝐚𝐭 𝐢𝐧 𝐭𝐡𝐞 𝐟𝐮𝐞𝐥 𝐜𝐨𝐧𝐬𝐮𝐦𝐞𝐝
The quantity of the heat to be imparted (Q) to the stock can be
found from the formula
Q = m×Cp× (T2-T1)
Where
Q = Quantity of heat
m = weight of the material
Cp = Specific Heat
T2 = Final Temperature
T1 = Initial temperature

DATA GIVEN:
Vacuum Gas oil (VGO) Flow rate = 160 m3
Outlet Temperature = 3420C
Inlet Temperature = 3180C
Gas flow rate = 90 m3
Specific Density of fuel gas = 0.5
Specific heat of flue gas =0.65 kcal/kg/0C
Gross Calorific Value of fuel oil = 11326 kcal/kg
Heat generated = 90 × 11326 × 0.5
= 509670 kcal/kg
Heat liberated = 160×894.1×0.65×(342-318)
= 2231673.6 kcal/kg
Efficiency =
(Heat liberated−Heat generated)
Heat generated
×100
= 77.18%

2. Indirect Method:
Furnace efficiency using indirect method, various parameters
that are required like sensible heat, loss due to evaporation of
moisture in fuel, radiation heat loss from the surface of furnace
and many more.
Efficiency is determined by subtracting all the heat losses from
100.
DATA GIVEN:
Vacuum Gas oil (VGO) Flow rate = 160 m3
Outlet Temperature = 3420C
Inlet Temperature = 3180C
Excess Oxygen = 10%
Gas flow rate = 90 m3
Specific Heat of Flue gas = 0.47 kcal/kg/0C
Flue gas Temperature after air preheater = 1800C
Ambient temperature = 400C
Mass of the fuel = 1 kg
Theoretical air required to burn 1 kg of fuel = 14 kg
Gross Calorific Value of fuel oil = 11326 kcal/kg
1. Sensible Heat Loss in the flue gas:
Excess air =
O%
21−O%
× 100
=
10
11
× 100 = 90.9% excess air
Total air supplied = Theoretical air ×(1 +
Excess air
100
)
= 14 × (1.909)

= 26.726 kg/kg of fuel
Sensible heat loss = m×Cp×∆T
= Weight of the flue gas
= Actual mass of air supplied /kg of fuel +
mass of fuel
= 26.726+1 = 27.726 kg/kg of fuel
Heat loss = 27.726 × 0.47 × (180-40)
= 1824.37
% heat loss in the flue gas =
1824.37 ×100
11326
= 16.11%
2. Loss due to evaporation of water formed due to hydrogen in fuel
% loss =
9×H×[584+0.45(Tflue−Tamb)]
GCV of fuel
×100
Where H = kg of the H2 in 1 kg of fuel oil (6-7% of the fuel)
Tflue = Flue Gas Temperature
Tamb = Ambient Temperature
GCV = Gross Calorific Value of the fuel
=
9×0.06×[584+0.45(180−40)]
11326
×100
= 3.08%
3. Radiation heat loss from surface of furnace
Consideration is made and in radiation loss we consider 4-5%
loss
Therefore, radiation loss from surface of furnace = 4%

Furnace Efficiency
1. Sensible heat loss in the flue gas = 16.11%
2. Loss due to evaporation of H2O formed from H2 in Fuel =3.08%
3. Radiation heat loss from surface of furnace = 4%
Total heat losses = 23.19%
Furnace efficiency = 100-23.19
= 76.81%

 To the heat balance of reactor regenerator in the fluidised
catalytic cracking.
Regenerator heat balance calculation:
1. Heat generated in the regenerator:
C to CO2 = 11016.85 kg/hr x 7832.4 kcal/kg
= 86.28 x106 kcal/hr
H to H2O =1222.43kg/hr x 28673.476 kcal/kg
= 35.05 x 106 kcal/hr
S to SO2 = 102.06 kg/hr x 2214.5 kcal/kg
= 0.226 x 106 kcal/hr
Total heat released in the regenerator
= (86.28 + 35.05 +0.226) x106 = 121.556 x 106 kcal/hr

2. Required heat to increase air temperature from blower
discharge to the regenerator dense phase temperature:
enthalpies of air at 190°C and at 709°C are 50.04 kcal/kg and
197.38 kcal/kg.
Therefore, the required heat is
= 0.184 x 106 kg/hr x (197.38-50.04)kcal/kg
= 27.11 x 106 kcal/hr.
3. Energy to desorb coke from the spent catalyst.
Desorption of coke = 12341.34 kg/hr x 806.2 kcal/kg
= 9.949 x 106 kcal/hr
4. Energy to heat the stripping steam:
Enthalpy of 50 psig-saturated steam=655.524 kcal/kg.
Enthalpy of 50 psig at 522°C = 844.56 kcal/kg.
Change of enthalpy = 5896.70 kg/hr x (844.56-655.524) kcal/kg
= 1.11 x 106 kcal/hr.
5. Energy to heat the coke on the spent catalyst:
=12362.8 kg/hr x 0.4 kcal/kg-°C x (709.4-522) °C
=0.925 x 106 kcal/hr.
6. Energy to heat the flue gas from regenerator dense phase to
regen crater flue gas temperature:
Enthalpy of flue gas at 709°C = 202.94 kcal/kg
and at 721°C = 209.05 kcal/kg.
The required heat is therefore
= 196784.03 kg/hr x (209.05- 202.94) kcal/kg
= 1.64 x 106 kcal/hr
7. Heat loss to surroundings: Assume heat loss from the stripper-
regenerator (due to radiation and convection) is 4% of total heat
of combustion, i.e. 0.04 x 121.556 x 106 kcal/hr
=4.86 x 106 kcal/hr.
8. Energy required to heat the spent catalyst from its reactor to the
regenerator temperature
= (121.556-27.11-9.49-1.11-0.925-1.64-4.86)x 106 kcal/hr
= 76.421 x106 kcal/hr.

Reactor heat balance calculation:
Heat into the reactor:
1. Heat with regenerator catalyst
= 1.437 x 106 kg/hr x 0.285 kcal/kg-°C x 709°C
= 290.55 x 106 kcal/hr
2. Heat with the fresh feed:
At a feed temperature of 312°C,the feed liquid enthalpy is
225.18 kcal/kg, therefore, heat content of the feed is
= 0.299 x 106 kg/hr x 225.18 kcal/kg
= 67.36 x 106 kcal/hr
3. Heat with atomizing steam:
From steam tables,
enthalpy of 68.1 kg saturated steam 653.8 kcal/kg,
therefore, heat with steam = 4540 kg/hr x 653.8 kcal/kg
=2.96 x 106 kcal/hr.
4. Heat of adsorption: The adsorption of coke on the catalyst is
an exothermic process; the heat associated with this
adsorption is assumed to be the same as desorption of coke in
the regenerator (i.e., 8.89 x 106 kcal/hr).
Total heat in = (290.55 + 67.36 + 2.96 + 8.89) x 106
= 369.76 x 106
kcal/hr.
Heat out of the reactor:
1. Heat with spent catalyst
=1.438x 106 kg/hr x 0.285 kcal/kg-°C x 522 °C
=214.12 x 106 kcal/hr.
2. Heat required to vaporize feed:
Enthalpy of reactor vapours 432.568 kcal/kg, therefore, heat
content of the vaporized products
= 299101.556 kg/hr x 432.568 kcal/kg
= 129.38 x 106 kcal/hr.

3. Heat content of steam:
Enthalpy of steam at 522°C = 844.564 kcal/kg, therefore,
heat content of steam 4540kg/hr x 844.564 kcal/kg
= 3.834 x 106 kcal/hr
4. Heat loss to surroundings
Assume heat loss due to radiant and convection to be 2% of
heat with the regenerated catalyst (i.e. 0.02 x 76.421 x 106
kcal/hr) =1.528 x 106 kcal/hr
Total heat out = (214.12+129.38+3.834+1.528)x 106
=368.89 x 106
kcal/hr
Calculation of heat of reaction
Total heat out = total heat in
Total heat out=368.89 x 106 kcal/hr
Total heat in = 369.76 x 106 kcal/hr.
Conclusion:
The total heat into the reactor is equal total heat out of the
reactor.

IOCL project report(chemical engineering)

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