Petroleum refining (3 of 3)

 Overview
 Chemistry
 Process Description
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Fuel Gas
Sulfur
C3 LPG
C4 LPG
Kerosene
Premium Gasoline
Regular Gasoline
Auto Diesel
Heating Oil
Haring Diesel
Heavy Fuel Oil
Bunker Oil

 In an alkylation process:
 Olefins + iso-paraffins  alkylate product (C7-C8 naphtha)
 The basic purpose of alkylation is to enhance the
octane number of the feed stock.
 Typical Components:
 propylene (C3)
 butylenes (C4) + isobutane
 amylene (C5 )
 Butylene is the most widely used olefin because of the
high quality of the alkylate produced.
 Isobutane is typically the iso-paraffin used
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 Typical Scheme:
 Gas Plant separates Olefins:
 For Polimerization
 For Alkylation
 In refining, the alkylation unit produces a high-quality gasoline blendstock by
combining two LPG-range molecules to form one gasoline-range molecule.
 This involves reacting isobutane with some type of light olefin, typically either
propylene or butylene coming from the FCC.

 Alkylate is a key component in reformulated
gasoline.
 In US and Europe about alkylate is about 11-12%
and 6%in the gasoline pool respectively.
 Alkylation processes are becoming important due
to growing demand for high octane gasoline
 Legislative regulations requirements
 Low RVP
 Low Sulphur
 Low toxics.
 Alkylate is an ideal blend stock to meet these
requirement.

 In 1930s, catalyst used was aluminium
chloride (AlCl3) catalyst
 New catalyst was replaced by HF + sulfuric
acid.
 The process of HF alkylation produces high
octane blend stock from:
 iso-parraffin (mainly iso butane)
 olefin (propylene, butylene and amylenes)
 NOTE:
 Replacing high risk toxic liquid acids (HF and
H2SO4)
 Vs. solid acid catalysts is challenging goal iso-
parraffin alkylation technology.
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 Note that alkylation of amylenes obtained from
C5 fraction of FCC can be another route to
increase the availability of alkylate.
 Alkylation of C5 cut from FCC can significantly
reduce RVP of finished gasoline pool.
 C5 alkylate, Amylene alkylation has two fold
advantage:
 It increase the volume of alkylate available
 Decreases Reid vapor pressure & Olefinic content of
gasoline blend stocks

 Some of the other side reaction is the formation of paraffin
 This typically will boil above and below the desired product.
 Impurities in the feed acid and normal operating practices
all can contribute to additional side reactions.
 Key factors:
 Maintaining proper composition of reaction mixture:
 isobutene olefins + the HF acid
 Maintaining the proper reaction environment:
 Correct contacting
 controlled temperature
 Freedom from surges.
 Making a proper separation of the reactor effluent into its
various components

 Alkylate is considered a premium gasoline
 This mainly because of:
 Low sulfur content - Alkylate has no sulfur
 Low aromatics content - Alkylate contains no aromatics
 Low vapor pressure - Alkylate has a low vapor pressure
 C4 alkylate has an RVP of 2.6 psi
 C3 alkyate has an RVP of 3.8 psi
 C5 alkylate has an RVP of 4.0 psi
 High octane - Alkylate has medium to high octane depending on the type
 C4 alkylate has a RON of 94-98 (MON of 92-94)

 To avoid olefin polymerization, high isobutane to olefin ratios are used.
 Typical isobutene to olefin ratios are 5:1 to 15:1
 Acid catalysts are used.
 Sulphuric acid (H2SO4)

 Reaction operating temperature
 10 - 20°C using H2SO4
 Reaction temperature
 4.4 bar for H2SO4
 When H2SO4 is used refrigeration is used.
 Reaction increases C7-C8 HC
 Propane and Butanes must be recycled
 Removal of gases are required

 Alkylation reactor:
 The reactor is arranged as a series of CSTRs
with acid fed in the first CSTR and feed
supplied to different CSTRs.
 This arrangement is for maximizing the
conversion.
 In the alkylation reactor it is important to
note that the olefin is the limiting reactant
and isoparaffin is the excess reactant.

 The reactor has two phases:
 olefin + isoparaffin mixture which will be
lighter
 the alkylate stream which will be heavier
 Alkylate will be appearing as a bottom
fraction if allowed to settle.
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 Phase separator:
 the acid enters the organic rich stream
and will be subjected to phase separation
by settling.
 Olefin/isoparaffin mixture will be also
separated by gravity settling.
 Thus the phase separator produces three
streams:
 (a) olefin + isoparaffin rich phase
 (b) acid rich stream
 (c) alkylate rich stream.

 Olefin + Paraffin processing:
 The olefin + paraffin stream is first
subjected to compression followed by
cooling.
 When this stream is subjected to throttling
and phase separation:
 olefin + paraffin rich stream will be
generated.
 The propane rich stream from this stream is
generated as another stream in the phase
separator.

 Propane defractionator:
 The propane rich stream after cooling is
fed to a fractionator:
 propane is separated from the olefin +
isoparaffin mixture.
 The O-IP mix is sent back to mix with the
olefin feed.

 Caustic wash for alkylate rich stream:
 The caustic wash operation ensures to
completely eliminate acid concentration
from the alkylate.
 The feed mixture (olefin + C4 compounds)
are first subjected to caustic wash.
 During caustic wash, sulphur compounds
are removed and spent caustic is recycled
back to the caustic wash.
 Fresh caustic solution is added to take
care of the loss.

 Alkylate fractionation separation
 The alkylate is fed to a distillation column
that is supplied with:
 isobutane feed
 alkylate feeds
 This produces:
 isobutane as a top product
 alkylate + butane mixture as a bottom
product.
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 Debutanizer:
 The debutanizer separates butane and
alkylate using the concept of distillation.

 To avoid olefin polymerization, high isobutane to olefin ratios are used.
 Typical isobutene to olefin ratios are 5:1 to 15:1
 Acid catalysts are used.
 HF are used.
 Reaction operating temperature
 25 – 40°C using HF
 Reaction temperature
 7.8 bar for HF
 When HF is used, refrigeration is not used.

 The process is similar to the sulphuric acid
plant.
 However, additional safety issues make the
process complex.
 The feed is first subjected to drying
followed by pre-cooling.
 After pre-cooling the reaction mixture is
fed to a reactor.
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 Unlike CSTRs in series here impeller
reactors are used.
 The reactor consists of:
 cooling tubes to absorb the heat generated.

 The reaction products enters a settler:
 oil and the HF are separated.
 Since there can be traces of HF in the oil
rich phase and vice-versa additional
processing is followed.

 The HF re-run column removes traces of
oils from the bulk of the HF.
 Thus HF purified will be recycled back to
the reactor.
 The bottom product thus generated in this
unit is acid oils.

 A HF stripper is used to remove:
 the HF (top)
 the alkylate product (bottom)
 Eventually, the HF stripper produces HF
that is sent back to the reactor and the
alkylate product.

 The alkylate product is sent to:
 a deisobutanizer
 The top product of the isobutanizer:
 C3-C4 mix

 The Depropanizer produces:
 Top  C3 pure
 Bottoms  C3 + iC4 traces
 These are recycled to the reactor
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 The final alkylate product is produced by
using a de-flourinator
 It will have a caustic wash
 Finally n-butane + alkylate is produced as
the bottom product
 N-Butane will typically be separated to
either:
 Isomerization unit
 LPG Plant

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 Overview
 Octane Number of HC
 Chemistry
 Process Conditions
 Process Description

 Reforming  Process in which low octane gasolines
will form high octane gasolines via chemical reactions
 Heavy naphtha which does not have high octane
number is subjected to reforming in the reformer unit
 Feedstock:
 Desulfurized Heavy Naphtha (from Naphtha Splitter)
 Produces:
 Light ends (to C3 Separator)
 Reformer Gas H2 (to HDS Naphtha)
 Reformate with high octane number (to gasoline blending
pool)
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 This unit produces high octane number
product that is essential to produce
premium grade gasoline as one of the
major refinery products.
 A reformer is regarded as a combination
of chemical and physical processes.
 Operating Conditions
 The initial liquid feed should be pumped
at a reaction pressure of 5 – 45 atm
 Preheated feed mixture should be heated
to a reaction temperature of 495 – 520°C.

 Catalytic reforming is the process of transforming:
 C⁷–C¹⁰ hydrocarbons with low octane numbers (heavy straight
run naphtha)
 to aromatics and iso-paraffins which have high octane numbers
 These gasolines are in fast growing demand

 The process will:
 Re-arrange or Re-structure the hydrocarbon
molecules in the naphtha feedstocks
 Breaks some of the molecules into smaller
molecules.
 It transforms low octane naphtha into:
 High-octane motor gasoline blending stock
 Aromatics rich in benzene, toluene, and
xylene
 Hydrogen
 LPG, liquefied petroleum gas as a byproduct.

 Typical RONC is (Research Octane Number) Order is given as follows:
 Paraffins < Branched paraffins < Naphthenes < Aromatics
 Note: Branched paraffins also have high octane.
 RONC (Octane number of naphtha) can be improved by reforming the hydrocarbon
molecule via Molecular rearrangement
 Such rearrangement takes place in reforming reactors in presence of catalyst by way of
numerous complex reactions.

 Typical Octane Number of Hydrocarbons:
 Recall that we aim to 90 +/- 5 Rating

 Following are the most prevalent main reactions in catalytic reforming
 Desirable
 Dehydrogenation of naphthenes to aromatics
 Conversion of paraffins to naphthenes
 Hydrogenation of oleffins to paraffins
 Dehydrocyclisation of paraffins to aromatics
 Non-Desirable
 De-alkylation of side chains of napthenes
 Hydrocracking of paraffins to lower molecular weight compounds
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 Proposed Mechanisms for Reforming Reactions

 Naphthene Dehydrogenation:
 Highly endothermic  cause decrease in temperatures
 Highest reaction rates
 Promoted by metallic function
 Aromatics formed have high B.P  end point of gasoline rises
 The reaction is promoted by the metallic function of catalyst
 Examples:
 Methyl cyclohexane  Toluene + H2
 DMCP  Toluene + H2

 Paraffin Dehydro-cyclisation:
 Multiple-step mechanism
 Lower rate of reaction than dehydrogenation
 Dehydrogenation followed by a molecular rearrangement to form a naphthene
 Naphthene is typically converted to the subsequent product
 Examples
 i-paraffins  aromatics of paraffins
 n-heptane  toluene + H2
 Favourable Conditions
 High temperature
 Low pressure
 Low space velocity
 Low H2:HC ratio

 Linear Paraffin Isomerisation
 Promoted by Acidic Function
 Branched isomers increase octane rating
 Small heat effect
 Fairly rapid reactions.
 Example:

 -->
 Favourable Conditions
 Low pressure
 Low space velocity
 H2/HC ratio no significant effect

 Naphthenes dehydro-Isomerisation
 A ring re-arrangement reaction
 Formed alkyl-cyclohexane dehydrogenate to aromatics
 Octane increase is significant
 Reaction is slightly exothermic

 Naphthenes Isomerization
 Required for subsequent dehydration to aromatic
 Difficulty of ring re-arrangement
 High risk of opening/breaking and paraffin formation
 Slightly endothermic

 Undesirable reactions:
 Hydrocracking*
 Hydrogenolysis
 Hydrodealkylation
 Further Alkylation
 Transalkylation
 Coking*
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 Coking:
 Coking is very complex group of chemical reactions.
 Linked to heavy unsaturated products such as poly-
nuclear aromatics.
 Traces of heavy olefines and di-olefines promote
coking.
 High feed FBP favors coking.
 Poor feed distribution in the reactor promotes coking
favored by high temperature

 Hydrocracking
 Exothermic reactions
 Slow reactions
 Consume hydrogen
 Produce light gases
 Lead to coking
 Causes are high paraffin concentration in feed
 Favourable conditions:
 High pressure

 Summary of Process conditions:
 Temperature
 Pressure
 H2:HC ratio
 Catalyst
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 Temperature is the most important operating parameter
 By simply raising or lowering reactor inlet temperature,
operators can raise or lower the octane number of the product.
 Pressure of Reactor
 Reforming reaction pressure ranges (5 – 35 atm)
 Decreasing pressure increases dehydrogenation of naphthenes
and dehydrocyclisation of paraffins
 This favours an increase in production of aromatics and
hydrogen
 It also increases catalyst coking  shorter cycle life
 Higher pressure causes higher rates of hydrocracking
 This reduces reformate yield BUT decreases coking of catalyst
resulting in longer cycle life.

 H2:HC Ratio
 Main purpose of hydrogen recycle is to increase hydrogen partial pressure in the reaction.
 H2 reacts with coke precursors removing them from the catalyst reforming polycyclic
aromatics.
 Higher the H2/HC ratio, higher the cyclic length.
 Two main reasons for reducing H2:HC ratio
 Reduction in energy costs for compressing and circulating H2.
 Favors naphthene dehydrogenations and dehydro-cyclisation reaction
 Thumb Rule:
 From 8 to 4 carbon increase in 1.75 times
 From 4 to 2 carbon increase 3.6 times

 Typical Catalyst
 Monometallic:
 Pt
 Bimetallic
 Pt, Rhenium
 Acid Activity:
 Halogens/silica incorporated in alumina base.
 Metallic Function
 It promote dehydrogenation and hydrogenation.
 It also contribute to dehydrocyclisation and isomerisation.
 Acid Function:
 It promotes isomerisation, the initial step in hydrocracking, participate in paraffin
dehydrocyclisation.

 Temporary Poisons
 Temporary poisons are those impurities which can be removed during various pretreatment
process like sulphur, nitrogen
 Permanent Poisons:
 Permanent Poisons are those impurities present in the feed which is irreversible damage to
the catalyst

 Basic steps in catalytic reforming involve
 Feed preparation
 Naphtha Hydrotreatment (see previous section on
Hydrotreatment)
 Pre-heating
 Temperature Control
 Reaction
 Catalytic Reforming
 Catalyst Circulation
 Regeneration
 Product separation
 Removal of gases and Reformate by fractional
Distillation
 Separation of aromatics in case of Aromatic production
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 Common Processes in Industry:
 Semi Regenerative catalytic reforming
 Cyclic catalytic reforming
 Continuous catalytic reforming(CCR)

 Common Processes in Industry:
 Semi Regenerative catalytic reforming
 Cyclic catalytic reforming
 Continuous catalytic reforming(CCR)
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 The feed is mixed with recycled hydrogen
 Subsequently, it is heated before sending to
reactor

 Since the reactions are highly endothermic,
several combinations of reactor + heaters are
used.

 The products from the final reactor are cooled.
 Often this is carried out with heat recovery
principle in which heat is recovered using the
fresh feed to the first reactor.
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 Product mixture enters a phase separator
which separates the hydrogen gas stream
from the liquid stream.

 The hydrogen produced in the phase
separator is compressed and sent back to the
first reactor.
 Excess hydrogen generated in the reactions is
taken out as a bleed stream
 Typically, it is used in:
 Hydrotreating & Crackers

 The liquid stream from the phase separator is
sent to a debutanizer distillation column
 It will separate butanes and lower alkanes
from the reformate product.

 Catalyst regeneration (not shown in the flow
sheet) needs to be carried out to regain
catalyst activity.
 This can be in different modes of operation:
 Semi-regenerative
 Cyclic
 Continuous.
Regeneration

 In this type of reformers the catalyst generally
has a life of one or more years between
regeneration.
 The time between two regeneration is called a
cycle.
 The catalyst retains its usefulness over
multiple regeneration.

 A semi-regenerative process uses low platinum and regeneration is required only
once a year.
 The dehydrogenation is highly endothermic and large temperature drop as the
reaction proceeds.

 Multiple reactors with intermediate reheat is required.
 Dehydrogenation of naphthene takes place in first reactor and requires less catalyst.
 Last reactor for isomerization of paraffins.
 Typical catalyst distribution in three reactors are 20, 30 and 50%.

 Catalyst Regeneration
 Performance of the catalyst decreases with
respect to time due to deactivation
 Coke formation
 Contamination on active sites
 Agglomeration
 Catalyst poisoning
 Activity could be restored if deactivation
occurred because of coke formation or
temporary poisons.
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 Objective of Regeneration
 Surface area should be high
 Metal Pt should be highly dispersed
 Acidity must be at a proper level
 Regeneration changes by the severity of the
operating conditions
 Coke formation can be offset for a time by
increasing reaction temperatures.
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 Cyclic reformers run under more severe operating conditions for improved octane
number and yields.
 Individual reactors are taken offline by a special valving and manifold system and
regenerated while the other reformer unit continues to operate.

 In these reformers the catalyst is in moving bed and
regenerated frequently.
 This allows operation at much lower pressure with a
resulting higher product octane, C5+, and hydrogen yield.
 These types of reformers are radial flow and are either
separated as in regenerative unit or stacked one above
the other.

 Check out this video:
 https://www.youtube.com/watch?v=FQ0ImB6eozY
 http://www.youtube.com/watch?v=6H0wpm70ckw
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 FCC Overview
 FCC Chemistry
 FCC Process Description

 The unit is one of the most important units of the
modern refinery.
 Feedstock:
 Hydrotreated Heavy Vacuum Gas Oils (from Heavy Gas
Oil Hydrodesulfurization Unit)
 Products
 Gaseous FCC Products (to Gas Treating Unit)
 Unsaturated light ends (to Alkylation Unit)
 Light cracked naphtha (to Gasoline Blending Pool)
 Heavy cracked naphtha (to Gasoline Blending Pool)
 Cycle oil (to Gas Oil Blending Pool)
 Slurry (to Fuel Oil Blending Pool)
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 Product Obtained
 Light gas
 H2, C1, and C2s
 LPG
 C3s and C4s – includes light olefins
 Gasoline
 C5+ high octane component for gasoline pool or light
fuel
 Light cycle oil (LCO) blend component
 diesel pool or light fuel
 Heavy cycle oil (HCO) Optional heavy cycle oil
product for fuel oil or cutter stock
 Clarified oil (CLO) or decant oil
 slurry for fuel oil
 Coke by-product consumed in the regenerator to
provide the reactor heat demand

 SR-Gasoline - straight run gasoline does
not have good octane number (40 – 60)
 It must be upgraded to obtain the desired
octane number (85 – 95)

 Typical processes in which gasoline increases in
octane rating:
 Cracking
 Reforming
 Isomerization
 These processes provide around 50% of all
transportation fuel and 35% of total gasoline pool.

 Typically cracking involves:
 thermal decomposition
 catalytic decomposition
 Since heat is required:
 Cracking reactions are carried out in furnaces
 These are supplied with either:
 Fuel oil
 Fuel gas
 Natural gas

 When cracking is carried out without a catalyst
 higher operating temperatures and pressures are required.
 This is called as thermal cracking.
 This was the principle of the old generation refineries.
 Now a days, cracking is usually carried out using a
catalyst.
 The catalyst enabled the reduction in operating pressure
and temperature drastically.
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 Proposed Mechanism:
 Cracking is an endothermic reaction (requires heat)
 Long chain paraffins converted to olefins
 Straight chain paraffins converted to branched
paraffins
 Alkylated aromatics converted to aromatics and
paraffins
 Ring compounds converted to alkylated aromatics
 Dehydrogenation of naphthenes to aromatics and
hydrogen
 Undesired reaction:
 Coke formation due to excess cracking

 About the catalyst:
 Acid treated silica-alumina was used as catalyst.
 20 – 80 mesh size catalysts used for FCCR and 3 – 4 mm pellets used for
MBRs
 During operation, poisoning occurs with Fe, Ni, Vd and Cu
 Zeolites:
 Have demonstrated vastly superior activity, gasoline selectivity, and
stability characteristics compared to original amorphous silica alumina
catalyst
 Today’s FCC catalysts Porous spray dried micro-spherical powder
 Particle size distribution of 20 -120 micron & particle density ~ 1400
kg/m3
 Comprising Y-Zeolite in many derivatives of varying properties
 Supplied under various grades of particle sizes & attrition resistance
 Continuing improvement metal tolerance, coke selectivity

 These very much depend upon the feed stock and type of cracking either thermal or
catalytic used.
 Cracking is a gas-phase reaction.
 Entire feedstock needs to be vaporized.
 Short reaction times (1 – 3 sec.) provides:
 good quality product
 less coke formation
 For vacuum gas oil:
 thermal cracking  600°C and 20 atms
 catalytic cracking  480°C and 0.7 – 1 atm

 Main reactions involved in catalytic cracking are
 Cracking
 Isomerisation
 Dehydrogenenation
 Hydrogen transfer
 Cyclization
 Condensation
 Alkylation and dealkylation

 Reactions:
 Paraffins  Smaller paraffins + olefins
 Alkyl naphthene  naphthene + olefin
 Alkyl aromatic  aromatic + olefin
 Multiring naphthene  alkylated naphthene
with fewer rings

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 The process technology consists of 3 flowsheets:
 Cracking coupled with main distillation column
 Regenerator
 Stabilization of naphtha.

 Three basic functions in the catalytic cracking
process are:
 Reaction
 Feedstock reacts with catalyst and cracks into different
hydrocarbons;
 Regeneration
 Catalyst is reactivated by burning off coke; and
recirculated to reactor
 Fractionation
 Cracked hydrocarbon stream is separated into various
products like:
 LPG and gasoline
 light cycle oil and heavy cycle oils

 Typical operating parameter of FCC
 Raw oil feed at heater inlet : 114 cubic meter /h
 Furnace outlet temperature : 291oC
 Reactor feed temperature : 371oC
 Reactor Vapour temperature : 549oC

 FCC Reactor
 Feed enters the cracking reactor.
 Fluidized catalytic cracking (FCC) reactors are
used.
 The cracked product from the reactor enters a
main distillation column
 It procures:
 Gas
 unstabilized naphtha
 light gas oil
 heavy gas oil
 Slurry

 It is therefore subjected to stabilized by
continued processing.
 The slurry enters a phase separation unit
which separates:
 decant oil
 heavier product.
 The heavier product is recycled back to the
cracking reactor.

 Naphtha Stabilization
 The naphtha obtained is unstabilized, as it
consists of various hydrocarbons.
 The unstabilized naphtha subsequently enters a
unsaturates gas plant
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 In the unsaturates (olefins) gas plant:
 the gas obtained from the main distillation
column is sent to a phase separator.
 The phase separator separates:
 lighter hydrocarbons
 heavier hydrocarbons.
 The phase separator is also fed with the
unstabilized naphtha.

 The unstabilized naphtha from the main
column is first fed to a primary absorber
 Here, absorbtion of heavier hydrocarbons in
the gas stream will occur.

 The gas leaving the primary absorber is sent to
a secondary absorber
 Here, the light gas oil from main distillation
column is used as a absorbent
 It will further extract any absorbable
hydrocarbons into the light gas oil.
 The rich light gas oil enters the main
distillation column

 The naphtha generated from the phase
separator is sent to stripping
 Here, it will continue further consolidation and
stabilization of naphtha.
 Now, stabilized naphtha is sent to distillation
in debutanizer and depropanizer units.

 The debutanizer unit removes:
 Butanes & lower hydrocarbons
 Stabilized Naphtha.
 The naphtha obtained as bottom product in the
debutanizer is termed as:
 “debutanized stable naphtha or gasoline”

 The butanes and other hydrocarbons are sent
to:
 depropanizer unit
 Butanes
 Propanes and lighter hydrocarbons.
 Butanes are obtained as lower product
 Propanes are obtained with other lighter
hydrocarbons as the top product in the
depropanizer unit.

 FCCR
 The basic principle of the FCCR is to enable the
fluidization of catalyst particles in the feed stream at
desired pressure and temperature.
 Another issue for the FCCR is also to regenerate the
catalyst by burning off the coke in air (see in regenerator)
 Therefore, the reactor unit should have basically two
units:
 Reactor (FCCR)
 Catalyst regenerator (CR).
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 Components:
 Riser
 Cyclone unit

 Riser
 In the riser (a long tube), the feed is allowed to get in contact
with the hot catalyst.
 The hot catalyst is enabled to rise through lift media in the
riser.
 The lift media is usually steam or light hydrocarbon gas.
 The riser contact time is about 250 milliseconds (0.25s)
 The riser is eventually connected to cyclone units.

 The cyclone units:
 They receive the catalyst and finished product.
 The catalyst that enters the cyclone unit is fully coked
 It needs to be sent to a regenerator to regain its lost activity.
 In here:
 Separation of the hydrocarbon vapors and catalyst as a solid
fluid operation will occur
 The catalyst falls down to the vessel that houses the riser and
cyclone units.
 The catalyst in the vessel is subjected to stream stripping in
which direct contact with steam is allowed to remove
hydrocarbons from the catalyst surface.

 Regenerator
 The spent catalyst which is relatively cold enters the
regenerator unit.

 Here air enters the vessel through a sparger set up.
 The catalyst is subsequently burnt in the air.
 This enables both:
 Heating the catalyst (which is required to carry out the
endothermic reaction)
 Removing the coke so as to regain the activity of the coke.
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 The catalyst + air after this operation will enter the cyclone
separator unit.
 Unlike the FCCR, the CR does not have a riser.
 Air enters a dense phase of catalyst
 It also enables the movement of the catalyst to a dilute phase
of catalyst + air

 The cyclone separators separate:
 flue gas
 Catalyst
 As a Solid-fluid operation.
 The activity regained catalyst is sent to the riser
through a pipe.
 Catalyst temperature is increased to 620 – 750°C
 The flue gas is obtained at 600 - 760°C and is sent for
heat recovery unit to generate steam.

 Read this Document:
 https://pubs.rsc.org/en/content/articlehtml/2015/cs/c5cs00376h

 Check this video:
 https://www.youtube.com/watch?v=u1tKTd3meUY
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 Check this video:
 https://www.youtube.com/watc
h?v=KeegB3q_668

 Overview
 Chemistry
 Process Technology

 Depending upon the intensity of the hydroprocessing
operation, the hydroprocessing is termed as:
 Hydrotreating
 Hydrocracking.
 During hydrotreating 
 Sulphur and nitrogen concentration in the final products is
reduced along with the saturation of olefins and aromatics.
 However, boiling range of the final products will be similar to
that of the feed stock.
 During hydrocracking 
 Heavier molecules react with hydrogen to generate lighter
hydrocarbons.

 Distillate hydrocracking is a refining process for conversion of:
 Heavy gas oils
 Heavy diesels
 Heavy distillates
 Into 
 light distillates (naphtha, kerosene, diesel, etc.)
 base stocks for lubricating oil manufacture.
 The process consists of:
 Hydrocarbon Feed reacts with hydrogen gas
 This is done in presence of a catalyst
 Typically, specific conditions are required:
 Temperature
 pressure
 space velocity

 Hydrocracking processes uses a wide variety of
feed stocks:
 Naphtha
 Atmospheric gas oil
 Vacuum gas oils
 Coke oils
 Catalytically cracked light and heavy cycle oil
 Cracked residue
 Deasphalted oils (DAO)
 Produces high quality product with excellent
product quality with low sulfur contents.

 The development of upgrading technology for
heavier stocks having:
 high sulfur
 nitrogen
 heavy metal (Ni, V)
 Conversion of low quality feed stocks 
 high quality products
 gasoline, naphtha, kerosene, diesel
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 The history of the hydrocracking process
goes back to late 1920 when
hydrocracking technology for coal
conversion was developed in Germany.
 During World War II, two stage
hydrocracking were applied in Germany,
USA and Britain.
 However, real breakthrough in
hydrocracking process was with the
development of improved catalyst due to
which processing at lower pressure.

 There has been continuous development in the
hydrocracking technology:
 Process conditions
 Catalyst in reaction
 Some of the important development in
hydrocracking has been:
 Mild hydrocracking (MHC)
 Resid hydrocracking (RHC)

 Mild hydrcracking (MHC) is characterized by relatively
low conversion (20-40%)
 Typically, Conventional HydroCracking gives 70-100%
 BUT!
 MHC route produces low sulphur diesel
 New mild hydrocracking route produces 10 ppm sulphur
diesel
 Under mild pressure.
 MHC allows increasing diesel production via:
 VGO hydroconversion.

 The yield of middle distillates obtained from
hydrocrackers is much more than that obtained from
other processes
 Hydrocracker does not yield coke or pitches as by
product.
 No post treatment is required for the hydrocracker
products.
 Most sulfur is transformed to H2S
 A typical hydrocracking reaction is as follows:
 C22H46 + H2  C16H34 + C6H14
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 Hydrocracking processes involved two types of catalyst:
 Hydrotreatment catalyst
 Hydrocracking catalyst
 Hydrotreating (Pretreat) Catalyst
 The main objective of pretreat catalyst is to remove organic
nitrogen from the hydrocracker feed allowing
 (i) Better performance of second stage hydrocracking catalyst, and
 (ii) The initiation of the sequence of hydrocracking reactions by
saturation of aromatic compounds
 Pretreat catalyst must have adequate activity to achieve above
objectives within the operating limits of the hydrogen partial
pressure, temperature and LHSV.

 Pretreat catalyst must have adequate activity to achieve
above objectives within the operating limits of the hydrogen
partial pressure, temperature and LHSV.
 (a) amorphous oxides (e.g. silica– alumina)
 (b) a crystalline zeolite (mostly modified Y zeolite) plus binder
(e.g. alumina)
 (c) a mixture of crystalline zeolite and amorphous oxides.

 Acid sites provide cracking activity.
 Examples are:
 Crystalline zeolite
 Amorphous silica alumina
 Mixture of crystalline zeolite
 Amorphous oxides
 Metals will provide hydrogenation dehydrogenation activity
 Examples:
 Noble metal (Pd, Pt)
 Non-noble metal sulphides (Mo, Wo or Co, Ni)]
 These metals catalyze the hydrogenation of feedstocks making them more reactive for
cracking and hetero-atom removal
 It is also a good reducing agent for the coke
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 Zeolite based hydrocracking catalysts have
following advantages:
 Greater acidity  greater cracking activity;
 Better thermal/hydrothermal stability
 Better naphtha selectivity
 Better resistance to nitrogen and sulphur
compounds
 Low coke forming tendency
 Easy regenerability.

 Hydrocracking process is a catalytic cracking process which takes place in the presence of an
elevated partial pressure of hydrogen
 It is facilitated by a catalyst having acidic sites and metallic sites.
 During hydrocracking process:
 hydrotreating reactions
 hydrocracking reactions
 Various hydrotreating reactions are:
 Hydrodesulphurization
 denitrogenation
 Hydrodeoxygenation
 Hydrometallization
 Olefin hydrogenation
 Partial aromatics saturation.

 Various hydrocracking reactions are splitting of C-C bond and or C-C rearrangement
reaction (hydrisomerisation process):
 Hydrogenation and dehydrogenation catalyst.

 CATALYST DEACTIVATION
 Catalyst activation may occur due:
 Coke Depositions
 Condensation of poly-nuclear and olefinic compounds into
high molecular weight which cover active sites.
 Metal Accumulation:
 Occurs at the pore entrances or near the outer surface of
the catalyst

 CATALYST REGENERATION
 Catalyst regeneration is done by burning off the carbon,
and Sulphur
 It is then re-circulated in the recycle compressor,
injecting a small quantity of air and maintaining catalyst
temperature above the coke ignition temperature.
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 The Hydrocracker unit consists of the following sections:
 Furnace
 First stage Reactor section
 Second stage Reactor section
 High pressure separator
 Fractionation Section
 Light Ends Recovery section

 In single stage process both:
 treating and cracking steps are combined
 In the 2-Stage Reactor we follow this:
 Feed along with recycle unconverted residue from
the fractionator is first hydrotreated in a reactor
 Then it will be combined with stream gases are fed
to second reactor
 Cracking takes place in the presence of hydrocracking
catalyst.
 In the single stage process the catalysts work under
high H2S and NH3 partial pressure.

 Feedstock:
 Cycle oils
 Coker distillates
 Product:
 High quality jet fuel
 Diesel production
 Cracking is promoted:
 silica-alumina sites of the catalyst.
 Hydrogenation promoted:
 palladium, molybdenum sulphide or tungsten sulphide
compounds.
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 Since catalyst gets poisoned with organic nitrogen
compounds, hydrotreater catalytic reactors are used
before hydrocracking reactors to safeguard the
hydrocracking catalysts.
 Excess hydrogen also aids in preventing catalyst
coking.
 Operating conditions of the hydrocracking reactor:
 340 – 425°C
 70 – 200 bar.

 Reactors use fixed or moving bed reactors
 *Fixed Bed Reactors are more common
 Fixed Bed (Packed)
 Cold shot reactors are used in which cold H2 is used to
cool the hot streams.
 Guard reactors are used before hydrocracking catalyst
within the reactor column itself
 Moving beds
 Feed allows movement of the catalyst for good mixing.

 After reaction, the product gets mixed
with water and enters a three phase
separator.
 The three phase separator generates
three streams:
 sour water stream
 organic stream
 gas stream.
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 The gas stream again enters a phase separator to
remove entrained organic stream after cooling.
 The gas stream is subjected to H2S removal using
amine scrubber.
 Sour Water is sent to treatment

 The organic stream eventually enters a steam
stripper further stabilize the organic stream.
 In this fractionator, a gas stream and a sour water
stream are generated.
 Eventually, the stabilized organic stream is sent to
a multi-product fractionators to generate:
 light naphtha
 heavy naphtha
 Kerosene
 Diesel
 Residue
 Steam is used to enhance the product quality.

 Installation of a Hydrocracker:
 https://www.youtube.com/watch?v=cl-C-0RglA8

 Diesel Hydrotreatment & Hydrocrackers
 https://www.youtube.com/watch?v=SvhbVt5RKCg
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 Catalyst of hydrocrackers
 Zeolites
 Unit Operation
 Chemistry
 https://www.youtube.com/watch?v=YqokaZ1e
5MY

 Visbreaking
 Coking:
 Delayed Coking
 Fluid Coking
 Flexicoking
 Asphalting:
 Solvent Deasphalting (SDA)
 Deasphalted Oil (DAO)
 Blown Asphalt

Fuel Gas
Sulfur
C3 LPG
C4 LPG
Kerosene
Premium Gasoline
Regular Gasoline
Auto Diesel
Heating Oil
Haring Diesel
Heavy Fuel Oil
Bunker Oil
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 NOTE:
 Do not confuse with FCC Units which is cracking

 Typical Content
 25%  Residue
 From that 25%:
 21% Coke

 The idea is to upgrade the residue product obtained from
the vacuum distillation unit.
 Significant amounts of vacuum residue is obtained from
various crude oils.
 For instance Arabian heavy oil produces 23.2 vol% vacuum
residue product.

 The residue consists of heavier hydrocarbons with
molecular weights ranging from 5000 – 10000.
 Thermal cracking is most preferable for the vacuum
residue.
 The vacuum residue also consists of other metals such as
vanadium and nickel.

 Typically, vacuum residue is subjected to six different operations:
 Vacuum residue desulphurization (VRDS)
 Residue fluid catalytic cracking (RFCC)
 Visbreaking*
 Coking*
 Deasphalting
 Gasification

 Cracking of heavy residue is most commonly used method for upgradation of
residues.
 This involves of decomposition of heavy residues by exposure to extreme
temperatures in the presence or absence of catalysts.
 THERMAL CRACKING:
 Cracking at elevated temperatures in the absence of catalyst:
 Visbreaking, delayed coking, Fluid coking etc.
 CATALYTIC CRACKING:
 Cracking in presence of catalyst
 FCC , Hydrocracking, DCC
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 Residue treatment still is improving the economics of the refinery through:
 low value fuel gas
 Mid/heavy distillates
 petroleum coke.
 Heavy residues are a mixture molecules consisting of an oil phase and an
asphaltene phase in physical equilibrium with each other in colloidal form.
 Asphaltenes are high molecular weight, relatively high atomicity molecules
containing high levels of metals.

 In TC the long molecules thus depleting the oil phase in the residue.
 Asphaltene cracking is the most difficult component to process
 At a certain condition asphaltenes is disturbed and asphaltenes precipitate.
 Under condition of thermal cracking, hydrocarbons, when heated, decompose into smaller
hydrocarbon molecules.

 Process Variables:
 Feed stock properties
 Cracking Temperature
 Residence time,
 Pressure
 Thermal Cracking:
 Medium
 High
 Ultra High
 Cracking with higher Temperature and with very short residence time
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 Visbreaking is a mild thermal cracking of vacuum
 It is a non-catalytic thermal process.
 It produces light products and 75–85% cracked material of lower
viscosity that can be used as fuel oil.
 Typical Yield:
 gas 1-2%
 naphtha 2-3%
 gas oil 5-7%
 furnace oil 90-92%.
 Vacuum residue is the heaviest distillation product and it
contains two fractions:
 heavy hydrocarbons
 very heavy molecular weight molecules
 (such as asphaltene and resins)

 In the resid, resins are holding asphaltene and keep them
attached to the oil.
 The cracking of resin will result in precipitation of
asphaltene forming deposits in the furnace and will also
produce unstable fuel oil.
 The cracking severity or conversion is limited by the storage
stability of the final residual fuel.

 Soaker drum utilizes a soaker drum in conjunction with a
fired heater to achieve conversion
 It reduces the viscosity and pour point of heavy petroleum
fractions so that product can be sold as fuel oil.
 It gives 80 - 85% yield of fuel oil and balance recovered as
light and middle distillates.

 The possible reactions in visbreaking are:
 Paraffinic side chain breaking
 which will also lower the pour point;
 Cracking of naphthens rings
 at temperature above 482 °C (900 °F);
 Coke formation by:
 polymerization, condensation, dehydrogenation and dealkylation;

 Visbreaking:
 Mild thermal cracking (low severity)
 Mild (470-500oC) heating at 50-200 psig
 Improve the viscosity of fuel oil
 Low conversion (10%) to 4300F
 Residence time 1-3 min
 Heated coil or drum

 Reaction:
 Splitting of C-C bond.
 Oligomerization and cyclisation to naphthenes of olefinic
compounds.
 Condensation of the cyclic molecules to polyaromatics.
 Side reactions
 Foramation of H2S, thiophenes, mercaptans, phenol
 Two classes of reactions occur during visbreaking
 Cracking of side chained aromatic compounds to produce
short chained aromatics and paraffins
 Cracking of large molecules to form light hydrocarbons
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 Visbreaking Conditions:
 Inlet Temperature: 305-3250C (15-40 bar)
 Exit: 480-5000C (2-10 bar)
 With soaking 440-4600C (5-15 bar)
 Feed: 900C pretreated with VB tar to 3350C
 Advantages:
 15% reduction in fuel oil
 Larger running time between two decoking operations
 Coke deposit rate 3-4 times slower than in conventional units.
 Better selectivity towards gas and gasoline productivity

 Variables:
 Feed rate & circulation
 Furnace transfer temperature
 Fractionation pressure & temperature
 Reflux flow
 Tar quench to transfer line
 tar quench to fractionator bottom
 tar quench to visbreaker tar stripper bottom
 Stabilizer temperature and pressure.

 There are two types of visbreakers:
 Coil visbreaking
 Thermal cracking occurs in the coil of the
furnace
 It yields a slightly more stable visbreaker
products, which are important for some
feedstocks and applications.
 It is generally more flexible and allows the
production of heavy cuts, boiling in the vacuum
gas oil range.
 Soak visbreaker
 Cracking occurs in a soak drum.
 Now all the new visbreaker units are of the
soaker type.

 Vacuum or atmospheric residue feedstock is heated and
then mildly cracked in the visbreaker furnace.
 Reaction T:
 850 to 900 °F (450 to 480 °C)
 Operating pressures:
 3 bar to as high as 10 bar.
 Coil furnace visbreaking is used and the visbroken
products are immediately quenched to stop the
cracking reaction.
 Coil cracking is described as:
 high temperature
 short residence time route

 The quenching step is essential to prevent:
 coking in the fractionation tower.
 The gas oil and the visbreaker residue are most
commonly used as quenching streams
 After quenching, the effluent is directed to the lower
section of the fractionator where it is flashed.
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 The fractionator separates the products into gas,
gasoline, gas oil and visbreaker tar (residue).
 The gas oil withdrawn from the fractionator is steam-
stripped to remove volatile components

 The un-stabilized naphtha and fuel gas recovered as
overhead products must be treated
 Then they will be added to further unitS:
 catalytic reforming
 Hydrotreatment
 Alkylation
 Blending pools

 The visbreaker bottoms are:
 withdrawn from the fractionator
 heat exchanged with the visbreaker feedstock
 mixed with stripped gas oil (optional)
 Routed to storage.

 Soaker cracking usually requires less capital
investment, consumes less fuel and has longer on-
stream times.
 Some visbreakers employ a soaker between the
visbreaker furnace and the quenching step
 This is similar to the conventional thermal cracking
processes.

 Conversion is mainly a function of two operating
parameters:
 temperature and residence time.
 soaker cracking is a low temperature
 long residence time route.

 Coking is very severe form of thermal cracking and
converts the heaviest low value residue to valuable
distillates and petroleum coke.
 Coking refers to extreme thermal cracking process.
 Feed is heated to 480 – 510°C and left for some time so
that coke and lighter products form.
 Since coking is a batch reaction, there can be different
ways to carry out coking.

 Recycle is used to further convert heavy distillate
fractions to lighter products
 Mechanism of coke formation:
 The colloidal suspension of the asphaltenes and resin
compounds is distorted
 This will form precipitation of highly cross linked structure of
amorphous coke
 The compounds are also subjected to cleavage of the aliphatic
groups.
 Polymerisation and condensation of the free aromatic radicals
 Grouping of the large number of these compounds to such a
degree that dense high grade coke is eventually formed
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 The process involves thermal conversion of vacuum
residue or other hydrocarbon residue resulting in:
 fuel gas
 LPG
 Naphtha
 gas oil
 Coke
 Various types of coking processes are:
 delayed coking
 fluid coking
 flexi coking

 Delayed Coking
 Delayed coking is a type of thermal cracking
in which the heat required to complete the
coking reactions is supplied by a furnace
 Coking itself takes place in drums operating
continuously on a 24 h filling and 24 h
emptying cycles.

 Delayed coking process is used to crack heavy oils
into:
 more valuable light liquid products
 solid coke as byproducts.

 Operates in semi batch mode
 Moderate (900-960oF) heating at 90 psig
 Soak drums (845-900oF) coke walls
 Coked until drum solid
 20-40% on feed, Yield 430°F, 30%
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 Feed:
 Vacuum residue
 FCC residual
 Cracked residue.
 Product:
 Gases
 Naphtha
 Fuel oil
 Gas oil
 Coke.

 Engineering Variables
 Batch
 Semi continuous or continuous
 Capacity and size factors
 Coke removal equipment
 Coke handling
 Storage
 Transportation
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 Feed stocks variables
 Chracterisation factor
 degree of reduction
 conardson carbon
 sulphur content
 Metallic constituents.

 Operating Variables
 Temperature
 Pressure
 recycle ratio
 transfer temperature
 coke chamber pressure

 Typical delayed coking consists of:
 a furnace to preheat the feed
 coking drums
 Fractionator column

 Fluid coking is non-catalytic fluid bed process
where residue is coked by spraying into a fluidized
bed of hot, fine coke particles.
 Higher temperature with shorter contact time than
delayed coking results in increased light and
medium hydrocarbons with less cake generation.
 Shorter residence time can yield higher quantities
of liquid less coke, but the product have lower
value

 The heated feed is fed to a fluidized bed where
coke particles with finer particle sizes would aid
fluidization.
 After coking, lighter products are withdrawn as
overhead vapour and coke thus formed is
removed continuously.
 The fluid coker also has an additional scrubber
which will remove heavier compounds from the
vapour (if any) and send them back with the feed
stream.
 Here, the feed stream absorbs heavier
hydrocarbons from the vapour generated.
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 This is required as it is difficult to keep
heavier hydrocarbons in the feed phase only
due to pertinent high temperatures.
 The coke after coking reaction is cold coke.
 Therefore, to generate hot coke, a burner
unit is used to heat the coke using
exothermic CO2reaction.
 The offgases from the burner are sent to
cyclones, scrubbing and then to the vent.
 The hot coke thus obtained is recycled back
to the fluidized bed or taken out as a net
coke product.

 The coker products are fed to a complex
distillation column i.e., main column
supplemented with side columns.
 From the complex distillation column,
naphtha, water, light gas oil and heavy gas oil
are obtained.
 Additional complexities in the distillation unit
are

 Feed entering the distillation column but not
the coker unit
 This is to facilitate the removal of light ends from
the feed (if any) and don’t subject them to
cracking.
 This is also due to the reason that light ends are
valuable commodities and we don’t want to loose
them to produce cheap coke product.
 In this case, the bottom product from the
distillation column is fed to the furnace for pre-
heating and subsequent coking operation.

 Live steam in distillation
 This is to facilitate easy removal of lighter
hydrocarbons in various sections.
 Circulatory reflux (Pump around units)
 This is to facilitate good amount of liquid
reflux in various sections of the main column.
For Further details, of the above two issues,
please refer to the crude distillation lecture
notes.
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 It is continuous process involves thermal cracking
in a bed fluidized coke and gasification of the
coke produced at 870 oC.
 This process contains an additional step of
gasification
 Temp: The gases leaving the gasifier is low
calorific value fuel gas at 800-1500 kcal/m3
(4200 to 5000 kJ/std m3 and is burned in the
furnace or power plants.

 It can be applied to wide variety of feed stocks.
 UOP UniflexTM Process [ Haizimann et al.,2010]:
 It is high conversion, commercially proven technology, that
processes low quality residue streams, like vacuum
residue, to make very high quality distillate products.
 The process utilizes thermal cracking to reduce molecular
weight of the residue in the presence of hydrogen and a
proven proprietary, nano sized catalyst to stabilize the
cracked products and inhibit the formation of coke
precursors.
 The main products from uniflex are naphtha and diesel
with a yield of greater than 80 vol%.

 De-propanation
 Stripper bottoms  lubes
 Asphalt
 Straight Run / Raw Bitumen (Paving)
 Oxidized Asphalt (Air Blown)

 Propane de-asphalting
 uses supercritical propane (meaning the
propane is at temperature and pressure
conditions above its critical point) as a
solvent
 Propane extracts and separate the lower-
boiling, lower-density hydrocarbon oil
molecules from the asphaltene molecules.
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 The extraction occurs in a vertical tower
 Operating Pressure 500 psi; 34 atmospheres
 Bottom temperature of 105 °F (40 °C)
 Top temperature of 140 °F (60 °C).
 The propane enters the extraction tower at
the bottom
 It travels upward counter-current to the
asphalt that enters the tower at the top and
flows downward.

 The propane solvent in the separated
deasphalted oil from the top of the extraction
tower is stripped out and recycled for re-use in
the extraction tower.
 The solvent-free, deasphalted oil (DAO) may
then be used as a feedstock component in
other petroleum refinery processes such as:
 Fluid catalytic cracker
 Hydrocracker
 This increases the production a gasoline
blending components

 Solvent in the asphalt stream from the bottom
of the extraction tower is also stripped from
the asphalt stream and recycled for re-use.
 The solvent-free asphalt may then be marketed
as end-product petroleum asphalt.
 Alternatively, all or some of the asphalt
product from the de-asphalter:
 May be processed in an air-blowing process
 This produces what is known as air-
blown or oxidized asphalt.

 Air-blowing process consists of:
 Air compressor to blow air through the liquid
asphalt
 Typical temperature ranges from 235 to 290 °C
 Combustion must be avoided (25 °C below the flash
point of the feedstock asphalt)
 Oxygen will oxidize tha sphalt
 Relevant Variables:
 the rate of air injection
 the system temperature
 the amount of time that the asphalt is kept in
contact with the air.

 The air-blown product asphalt has a higher
temperature softening point than asphalt
which has not been air-blown and that is a
desirable property for certain uses of
petroleum asphalt
 The end-product petroleum asphalt is
typically maintained at a temperature of
about 150 °C during storage at the petroleum
refinery as well as during transportation to
the asphalt end-users.
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 Air-blowing - The manufacturing process used to make oxidized roofing asphalts in
which air is blown through an asphalt flux.
 An exothermic oxidation reaction occurs, yielding asphalt that is:
 harder, more viscous, less volatile, and less temperature-susceptible than the asphalt flux used
as the feedstock to the process.
 Asphalt treated by blowing air through it at elevated temperatures to produce physical
properties required for the industrial use of the final product.

 Oxidized Asphalt is a solid or semi-black solid material and gradually liquid when heated
 It is predominant content in:
 bitumen, asphalten and aromatic resins.
 The oxidized asphalt or blown asphalt is producing in different softening point and penetration like
as:
 115/15
 90/40
 85/25
 85/25
 90/15
 95/25
 75/25
 150/5
 105/30
 90% is used in paving roads

 https://www.youtube.com/watch?v=fpyTdlziAZs
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 Check out this!
 https://www.youtube.com/watch?v=qZZigKd43NU

 https://www.youtube.com/watch?v=zP1mpDJ0KIk

 Hydrogen Gas Production
 Syngas (CO+H2)
 Amine Treating  DEA
 MEROX Unit
 Sulfur Recovery  Claus Process
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 Hydrogen gas is very versatile
 Fuel
 Hydrogenation
 Ammonia production

 Following methods:
 Treatment of gas mixture  manufacture of coke, olefins by steam cracking and catalytic
reforming
 Decomposition of hydrocarbons -into carbon and hydrogen, from partial oxidation or
steam treatment
 Water decomposition  electrolysis, thermochemical cycles
 Most commonly:
 Partial Oxidation (heavy, non-valuable hydrocarbons)

 Raw Materials:
 Oxygen / Air
 Fuel oil
 Product Gases

 ‘Synthesis gas’ is commonly used to describe two basic gas mixtures:
 synthesis gas containing CO, hydrogen (CO+H2)
 synthesis gas containing hydrogen and nitrogen for the production of ammonia (N2+H2)
 Methane and synthesis gas are important petrochemical feedstock for manufacture
of a large number of chemicals
 which are used directly or as intermediates
 Many of these products are number of which
are finding use in:
 Plastic
 synthetic fiber
 Rubber
 pharmaceutical and other industries.

 Petrochemical derivatives based on synthesis gas
and carbon monoxide have experienced steady
growth due to large scale utilization of methanol
 It is an excellent basis for the synthesis of some
valuable petrochemicals such as
 ammonia, methyl alcohol, formic acid, acetic acid,
formaldehyde, phosgene and others.
 Recent market studies show that there will be a
dramatic increase in demand of CO and syngas
derivatives.

 Methanol is the largest consumer of
synthesis gas.
 The reformed gas is to meet certain
requirements with regard to its
composition.
 It is characterized by the stoichiometric
conversion factor, which differs from case
to case

 Various raw materials for synthesis gas production
are:
 natural gas, refinery gases, naphtha, fuel
oil/residual heavy hydrocarbons and coal.
 Although coal was earlier used for production of
synthesis gas, it has now been replaced by
petroleum fractions and natural gas.
 Petrocoke is the emerging source for Synthesis
gas.
 Coal is again getting importance alone are with
combination of petroleum coke.

 SYNGAS production technologies:
 steam methane reforming
 naphtha reforming
 auto-thermal reforming
 oxygen secondary reforming
 partial oxidation of heavy
hydrocarbons, petroleum coke and
coal.

 Various steps involved in synthesis gas production through steam reforming are:
 Desulphurization of gas
 Steam reforming and compression
 Separation of CO2
 Various available synthesis gas generation schemes are:
 Conventional steam reforming
 Partial oxidation
 Combined reforming
 Parallel reforming
 Gas heated reforming

 Partial oxidation:
 O2, N2, H2, CO, CO2, H2S
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 The basis for the manufacture of synthesis gas at
the beginning of the nineteenth century was:
 The interaction of coal with water vapor at high
temperatures

2 2C H O CO H  

 Another method to prepare synthesis gas from
natural gas is called the partial oxidation of
methane as follows (Steam Reforming)

 This method is called steam reforming in which
 Nickel is used as a catalyst
 Temperatures ranging from 800 to 850°C
 Pressure ranging between 25 and 40 atmospheres
4 2 23CH H O CO H  

 Another method to prepare synthesis gas from natural gas
is called the partial oxidation of methane as follows:

 At temperatures of more than 1500°C
 Pressure range between 130 and 140 atmospheres.
1
4 2 22 2CH O CO H  

 A) Naphtha has to be gasified before reaction, which means
additional equipment, energy and financial costs.
 B) The sulfur content in naphtha is high and its removal is
more difficult than the removal of hydrogen sulfide in natural
gas
 for example, naphtha has to be treated carefully and costly to
remove sulfur before use.
 C) A larger quantity of carbon dioxide is formed when using
naphtha, and disposal costs are greater:


 2.1 2 2 22 3.05Naptha CH H O CO H  
4 2 23CH H O CO H  
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 Hydrogen sulfide is most commonly obtained by its separation from sour gas, which is
natural gas with high content of H2S.
 Fuel gases must be treated as well
 Hydrocarbons can serve as a source of hydrogen in this process

 The processes that use aqueous solutions of various alkylamines (commonly referred to
simply as amines) to remove hydrogen sulfide (H2S) and carbon dioxide (CO2) from gases
 It is a common unit process used in refineries, and is also used in petrochemical plants,
natural gas processing plants and other industries.

 In oil refineries, that stripped gas is mostly
H2S, much of which often comes from a
sulfur-removing process called
hydrodesulfurization.
 This H2S-rich stripped gas stream is then
usually routed into a Claus process to
convert it into elemental sulfur.

 In fact, the vast majority of the 64,000,000 metric tons of
sulfur produced worldwide in 2005 was by-product sulfur from
refineries and other hydrocarbon processing plants.
 Another sulfur-removing process is the WSA Process which
recovers sulfur in any form as concentrated sulfuric acid.
 In some plants, more than one amine absorber unit may share a
common regenerator unit.
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 Many different amines are used in gas treating:
 Diethanolamine (DEA) Monoethanolamine (MEA)
 Methyldiethanolamine
 (MDEA) Diisopropanolamine
 (DIPA) Aminoethoxyethanol
 Diglycolamine (DGA)
 The most commonly used amines in industrial plants are the alkanolamines DEA,
MEA, and MDEA

 Here is an example from:
 Hydrocrackers
 The yellow unit is actually an
Amine Gas Treating Unit
 Verify inlet/outlets

 The amine concentration in the absorbent aqueous solution is an
important parameter in the design and operation of an amine gas
treating process.
 Depending on which one of the following four amines the unit was
designed to use and what gases it was designed to remove, these
are some typical amine concentrations, expressed as weight
percent of pure amine in the aqueous solution:
 Monoethanolamine
 About 20 % for removing H2S and CO2
 Diethanolamine
 About 20 to 25 % for removing H2S and CO2
 Methyldiethanolamine
 About 30 to 55 % for removing H2S and CO2
 Diglycolamine
 About 50 % for removing H2S and CO2
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 Feed = Sour Gas
 Make-up Water
 The chemistry involved in the amine treating of
such gases varies somewhat with the particular
amine being used.
 For one of the more common amines,
monoethanolamine (MEA) denoted as R-NH2, the
chemistry may be expressed as:
 R-NH2 + H2S ⇌ R-NH3+ + SH−

 The process includes an:
 absorber unit
 regenerator unit
 In the absorber:
 the downflowing amine solution absorbs H2S and CO2
from the up-flowing sour gas
 Top Outlet:
 a sweetened gas stream as a product
 i.e., a gas free of hydrogen sulfide and carbon dioxide
 This is our “FINAL Product” or “Sweetened Product”
 Bottom Outlet:
 an amine solution rich in the absorbed acid gases.

 The resultant "rich" amine is then routed into the
regenerator (a stripper with a reboiler)
 This produces regenerated or "lean" amine that is
recycled for reuse in the absorber.
 Gaseous material is removed on “tops”
 The stripped overhead gas from the regenerator is
concentrated H2S and CO2.
 The “bottoms” is Lean amine, ready to be reused
 HEN is used to take advantage of heat flow

 Merox is an acronym for mercaptan (thiol) oxidation.
 It is a proprietary catalytic chemical process developed
by UOP used in oil refineries and natural gas processing
plants to remove mercaptans from:
 LPG, propane, butanes
 light naphthas
 kerosene
 jet fuel
 The Merox process requires an alkaline environment
which, in some process versions, is provided by:
 Aqueous sodium hydroxide (NaOH), a strong base
 Ammonia, which is a weak base.
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 For treatment of light feed stocks such as LPG,
no sweetening is required as mercaptans are
nearly removed by extraction.
 However, feed containing higher molecular
weight mercaptans and may require a
combination of Merox extraction and
sweetening using catalyst.
 Catalysts promote the oxidation of mercaptans
to disulphide using air as the source of oxygen.

 To improve:
 lead susceptibility of light gasoline
 the response of gasoline stocks to oxidation inhibitors
added to prevent gum formation during storage
 odor on all stocks
 To reduce:
 the mercaptans content to meet product specifications
 The sulphur content of LPG and light naphtha products
 The Sulphur content of coker FCC olefins to save acid
consumption in alkylation

 Process
 Pretreatment (Remove H2S and Naphthenic
Acids by dilute Alkali Solution)
 Extraction (Remove Caustic soluble
Mercaptans)
 Sweetening(Oxidation of mercaptans to
disulphides)
 Post Treatment (Remove Caustic Haze)
 Caustic Settler
 Wash Water
 Sand
 Clay Filters

 The LPG (or light naphtha) feedstock enters the
prewash vessel
 It then flows upward through a batch of caustic
which removes any H2S that may be present in
the feedstock.
 The coalescer at the top of the prewash vessel
prevents caustic from being entrained and
carried out of the vessel.

 The feedstock then enters the mercaptan
extractor and flows upward through the contact
trays where the LPG intimately contacts the
downflowing Merox caustic that extracts the
mercaptans from the LPG.

 The sweetened LPG exits the tower and flows
through:
 a caustic settler vessel to remove any entrained
caustic
 a water wash vessel to further remove any residual
entrained caustic
 a vessel containing a bed of rock salt to remove any
entrained water.
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 The caustic solution leaving the bottom of the
mercaptan extractor ("rich" Merox caustic) flows
through a control valve which maintains the
extractor pressure needed to keep the LPG
liquified.
 It is then injected with UOP's proprietary liquid
catalyst (on an as needed basis)
 It then flows through a steam-heated heat
exchanger

 It is injected with compressed air before
entering the oxidizer vessel where the
extracted mercaptans are converted to
disulfides.
 The oxidizer vessel has a packed bed to
keep the aqueous caustic and the water-
insoluble disulfide well contacted and well
mixed.

 The caustic-disulfide mixture then flows into
the separator vessel where it is allowed to
form a lower layer of "lean" Merox caustic and
an upper layer of disulfides.
 The vertical section of the separator is for the
disengagement and venting of excess air and
includes a Raschig ring section to prevent
entrainment of any disulfides in the vented air.
 The disulfides are withdrawn from the
separator and routed to fuel storage or to a
hydrotreater unit. The regenerated lean Merox
caustic is then pumped back to the top of the
extractor for reuse.

 A typical Jet Fuel Merox Unit
 Verify:
 Units
 Feedstock / Raw Materials
 Recycle ratio
 Typical Vessels

 Most Hydrogen Sulfide was converted from:
 Crude Oil distillation gases
 Hydrocrackers
 Hydrotreaters
 FCC
 Cokers
 Hydrogen Sulfide must be converted to Sulfur

 Sulphur recovery now has become one of the
most critical aspects of sulphur management
and affects emission sulphur dioxides
significantly in the refinery.
 There are two sulphur recovery processes are
 Claus process(used earlier)
 Super Claus process
 Conventional Claus process has only 99% sulphur
recovery.
 In order to meet the Sulphur emission standards
now Claus process has been improved
substantially to meet the standards.
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 A typical SUPER CLAUS sulphur recovery unit consist of following sections:
 Combustion Chamber
 Claus reactor
 Super Claus Reactor
 Incinerator
 Degassing Section

 The big difference between SUPER CLAUS catalyst
and Claus Catalyst is that the reaction is not
equilibrium based.
 Therefore, the conversion efficiency is much higher
than the equilibrium limited Claus reaction.
 SUPER CLAUS is a non-cyclic process that has
repeatedly shown:
 simplicity in operation
 high reliability
 sulphur guarantees up to 99.3 %

 Sour Process
 Stringent environmental regulations have necessitated higher
recovery of H2S from sour water stripper unit design.
 Super Sour process
 Ensures minimum H2S loss.
 The process employ additional hot feed flash drum upstream
of cold feed surge drum.

 In a furnace reactor, H2S is partially oxidized with air to produce water and SO2.
 The reaction is highly exothermic.
 Therefore, steam is generated using the products from the furnace reactor.
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 The remaining H2S is then sent to a converter at about 250°C to allow the reaction
between H2S and SO2 and produce Sulphur and water.
 The emanating product is at 290°C.
 The second reactor (H2S to SO2) is having severe equilibrium limitations.

 Therefore, it is sent to two to three reactors for maximizing conversion.
 After each converter, the product stream is cooled and sent to another reactor.
 Subsequently, Sulphur is removed as a product from the coolers.

 Finally tail gas is obtained from the last converter which consists of unreacted H2S,
N2 and O2.
 The tail gas requires treatment as well.
 This is because the gas consists of components such as H2S, CS2 etc.

 The tail gas is fed with air to another burner and converter that converts sulphur
compounds to H2S.
 The H2S thus generated is separated using amine scrubbers.
 The H2S thus recovered is sent as a recycle stream to the partial oxidation reactor.

 Check out this file:
 https://www.ecotech.com/wp-
content/uploads/2015/03/H2S-
1100-user-manual-Rev-1.5.pdf

 Claus Process:
 https://www.youtube.com/watch?v=b9Rhl9kVvZ0
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 Finally! You made it!

1. Introduction
2. Overview of the Petroleum Refining
3. Crude Oils
4. Products
5. Crude Oil Distillation (ATM/Vacuum)
6. Hydrotreatment
7. Gas Processing
8. Polymerization Unit
9. Isomerization Unit
10. Alkylation Unit
11. Catalytic Reforming
12. Fluid Catalytic Reformers (FCC)
13. Hydrocracking
14. Thermal Cracking & Coking
15. Secondary Processes
16. Conclusion
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Petroleum refining (3 of 3)

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