PENEX PROCESSES
In worldwide production of automotive gasoline permanent tendency to the toughening
of not only its operating but also its ecological characteristics is observed. So, international and
domestic regulations to automotive gasoline considerably limit the content of benzene, aromatic
hydrocarbons, olefin hydrocarbons and sulfur. In 1970s the variants of hydrogenation of the
benzene, contained in the reformate, proceeding without the decrease of product octane number
have been offered. However for decrease of the total aromatics content the dilution of reformate
with high-octane nonaromatic components is required. This situation is complicated by refusal
from tetraethyl lead (TEL) and deficit of butane-butylene fraction (because of the lack of FCC
duty), which is used for the production of high-octane alkylate in the world practice.
Thereby the development of isomerization process is one of the effective methods for
solution of this problem. It allows the producing of commercial gasoline which corresponds to
the current and perspective requirements to the fuels and provides necessary flexibility of
processing.
1. TYPES OF ISOMERIZATION PROCESSES



Three types of industrial isomerization processes are worked out currently:
high-temperature isomerization process (360-440 °С) on fluorinated-alumina catalysts;
medium-temperature isomerization process (250-300 °С) on zeolite catalysts;
low-temperature isomerization process on chlorinated-alumina catalysts (120-180 °С)
and sulfated metal oxides (180-210 °С ).
2. THERMODYNAMIC AND KINETIC LAWS OF ISOMERIZATION PROCESS
The schemes of proposing processes are analogous generally. The differences are defines
by performances of usable catalysts due to their type. Main parameter which is the octane
number of produced isomerizate depends on process temperature. That‘s why we will dwell on
the issue of thermodynamic of isomerization reaction. First of all hydrocarbons isomerization
reaction is balanced reaction, and equilibrium yield of isoparaffins increases with temperature
reducing, but it can be reached only after an ―infinite residence time‖ of the feed in reaction zone
or an equivalent very small value for LHSV. On the other hand an increase in temperature
always corresponds to an increase in reaction velocity. So that at low temperature the actual yield
will be far below the equilibrium yield, because of low reaction velocity. On the contrary, at
higher temperature, the equilibrium yield will be more easily reached, due to a high reaction rate.
Consequently, at higher temperature the yield of isoparaffins is limited by the thermodynamic
equilibrium, and at lower temperature it is limited by low reaction rate (kinetic limitation)
(Figure3.1). The comparative estimation of isopentanes content in sum of pentanes for different
types of isomerization catalysts is represented below (Figure 3.2).
Figure 3.1. Dependence of n-paraffins conversion on reaction temperature
Figure 3.2. Comparative estimation of isomerization catalysts
The conversion level of n-paraffins on zeolite catalysts is low, as it is limited by thermodynamic
equilibrium. In the case of chlorinated-alumina catalysts and sulfated metal oxides conversion of
n-paraffins is higher because of high equilibrium content of isocomponents in product.
3. TECHNOLOGIES OF ISOMERIZATION PROCESS ON DIFFERENT CATALYSTS

ZEOLITE CATALYSTS
Zeolite catalysts are less active and used at higher operating temperature compared to
another types of catalysts, and consequently the octane number of isomerizate is low. However
they possess high resistance to impurities in the feed and capability for total regeneration in the
reactor of the unit. The technological scheme of this process is provided with fire-heaters for
heating hydrogen and feed mixture up to reaction temperature.
It is necessary high ratio of hydrogen to hydrocarbon feed (along with isomerization,
hydrogen is spent for hydrotreating and dearomatization of the feed); that‘s why compressor for
supplying of recycle hydrogen-rich gas and separator for separation of hydrogen-rich gas are
necessary .
Hysopar catalyst should be marked out among zeolite catalysts; it is the most progressive
in the world catalyst market, because it considerably exceeds all another catalysts by resistance
to impurities in the feed (available sulfur content is 100 ppm permanently and 200 ppm during
short periods of time)

CHLORINATED-ALUMINA
Chlorinated-alumina based catalysts are the most active and supply the highest
isomerizate yield and isomerizate octane. It should be noted that during isomerization catalysts
loose chlorine, consequently the activity is reduced. That‘s why chlorine compound injection to
the feed (usually ССl4) is provided for keeping of high activity. As a result, caustic soda washing
from organic chloride in special scrubbers is necessary. Considerable drawback is that this type
of catalyst is very sensible to poisonous impurities (to the oxygen compounds including water, to
nitrogen) and requires pretreatment and drying of the feed. In addition the problems occur at
regeneration
The first generation catalyst of UOP is I-8, which was improved later in more active I-80
type catalyst. The latest developments of UOP Company are high-performance I-8 Plus, I-82, I84 catalysts for Penex process and I-122, I-124 catalysts, which are used in Butamer process (nbutane isomerization process with purpose to produce isobutane, which is the feed for alkylation
unit). In development of new catalysts UOP has the target to decrease its platinum content
without losing the activity, thereby to reduce significantly its operating costs. It is not of small
importance for present-day refinery.

SULFATED METAL OXIDES BASED CATALYSTS
Sulfated metal oxides based catalysts get heightened interest last years as they combine
main advantages of medium-temperature and low-temperature catalysts. They are active,
resistant to poisonous impurities and able for regeneration. The only drawback, as for zeolite
catalysts, is necessity in compressor for recycling of hydrogen-rich gas.
The CИ-2 catalyst has an activity, which is higher than activity of PI-242 [5] and
characterized with unique sulfur resistance. If necessary, the process can be carried out without
pretreatment of the feed. In this case the octane number of isomerizate is reduced by 2 points, but
total lifetime (8-10 years) doesn‘t changes and service cycle is no less than 12 months. The feed
may contain considerable quantity of benzene which is hydrogenated efficiently on the catalyst.
The Pt/WO3-ZrO2 catalyst shows higher activity and selectivity in isomerization reaction
of n-alkanes compared to sulfated-zirconia catalysts. The advantage of this type of catalyst is
explained by rapid surface diffusion of hydrogen atoms, which are converted into protons and
hydrides on the Lewes acid sites, thereby increasing catalyst activity and selectivity.
TECHNOLOGIES SCHEMES OF ISOMERIZATION PROCESS
Penex Process
The Penex process has served as the primary isomerization technology for upgrading
C5/C6 light straight-run naphtha feeds since UOP introduced it in 1958. This process has a wide
range of operating configurations for optimum design flexibility and feedstock processing
capabilities. The Penex process is a fixed-bed procedure that uses high activity chloridepromoted catalysts to isomerize C5/C6 paraffins to higher octane branched components. The
reaction is conducted in the presence of a minor amount of hydrogen. Even though the chloride
is converted to hydrogen chloride, carbon steel construction is used successfully because of the
dry environment. For typical C5/C6 feeds, equilibrium will limit the product to 83 to 86 RON
(Research Octane Number) on a single hydrocarbon pass basis.
To achieve higher octane, UOP offers several schemes in which lower octane
components are separated and recycled back to the reactors. These recycle modes of operation
can lead to product octane as high as 93 RON.
Hydrocarbon Once-Through Penex Process
Hydrogen Once-Through Penex process flow scheme results in a substantial saving in
capital equipment and utility costs by eliminating product separator and recycle gas compressor.
The stabilizer separates the light gas from the reactor effluent (Fig3.3).
Typically, two reactors in series are used to achieve high on-stream efficiency. The
catalyst can be replaced in one reactor while operation continues in the other. One characteristic
of the process is that catalyst deactivation begins at the inlet of the first reactor and proceeds
slowly as a rather sharp front downward through the bed. The adverse effect that such
deactivation can have on unit on-stream efficiency is avoided by installing two reactors in series.
Each reactor contains 50% of the total required catalyst. Piping and valving are arranged to
permit isolation of the reactor containing the spent catalyst while the second reactor remains in
operation. After the spent catalyst has been replaced, the relative processing positions of the two
reactors are reversed. During the short time when one reactor is off-line for catalyst replacement,
the second reactor is fully capable of maintaining continuous operation at design throughput,
yield, and conversion.
Several factors are considered when choosing a process flow scheme. One of the most
important aspects is desired product octane.
The hydrocarbon once-through flow scheme is the most widely used isomerization
process for producing moderate octane upgrades of light naphtha. Economically efficient ―onethrough‖ scheme without any recycle can be used with minimum investment in realization of
isomerization process Figure (3.3).
Figure (3.3) Block diagram of “one- through” process
TABLE 3.2 Typical Estimated Yields for Once-through Processing
Penex Process With Recycle And Fractionation
Separation and recycle of unconverted normal C5 and C6 paraffins and low octane C6
isoparaffins back to the reactor, produce a higher octane product. The most common flow
scheme uses a deisohexanizer (DIH) column to recycle methylpentanes, n-hexane, and some C6
cyclics. It is the lowest capital cost option of the recycle flow schemes and provides a higher
octane isomerate product, especially on C6 rich feeds. In the Penex/DIH process the stabilized
isomerate is charged to a DIH column producing an overhead product containing all the C5 and
dimethylbutanes. Normal hexane and some of the ethylpentanes are taken as a side-cut and
recycled back to the reactors. The small amount of bottoms (C7+ and some C6 cyclics) can be
sent to gasoline blending or to a reformer.
The addition of a deisopentanizer (DIP) or a super DIH will achieve the highest octane
from a fractionation hydrocarbon recycle flow scheme. In this scheme, both low octane C5 and
normal and isoparaffin C6 are recycled to the Penex reactors .
The scheme with deisopentanizer (DIP) before the reactor section allows the producing of
isomerizate with high octane number, increasing of conversion level of n-pentanes and reducing
the reactor duty simultaneously. The technology is reasonable in the case of isopentanes content
in the feed more than 13-15 % Figure (3.4)
Figure (3.4) Block diagram of process with DIP
The scheme with deisohexanizer (DIH) after the isomerization reactor is the simplest way
to produce the isomerizate with higher octane number. In this case non-converted low-octane
components (methylcyclopentane and n-hexane) are recycled into reactor. However the given
scheme allows only increasing of hexanes conversion, but doesn‘t raise the content of
isopentanes in the product (Figure 3.5). The scheme of the process may include both
deisopentanizer and deisohexanizer (with DIP and DIH)
Figure (3.5) Block diagram of process with DIH
TABLE 3.3 Typical Estimated Yields for Deisohexanizer Processing
Scheme with recycle of n-pentane (with DIP and DP) requires providing with
depentanizer of isomerizate after the reaction section and deisopentanizer before the reactor.
Schemes with recycle of n-pentane and n-hexane. For total conversion of all linear paraffins (not
only n-С6 but also n-С5) into isomers, their total recycle is necessary which can be realized by set
of distillation columns (with DIP, DIH and DP) or by adsorption on molecular sieves.
The method of adsorption on molecular sieves (in liquid or vapor phase) is based on capability of
pores with definite size to adsorb selectively the molecules of n-paraffins. The next stage is
desorption of n-paraffins from pores and its recycle to the feed stock. Stages of adsorption and
desorption are repeated in cycles or pseudo-continuously.
Penex / Molex Process
This flow scheme uses Molex technology for the economic separation and recycle of nparaffin from the reactor effluent. The Molex process is an adsorptive separation method that
utilizes molecular sieves for the separation of n-paraffins from branched and cyclic
hydrocarbons. The separation is effected in the liquid phase under isothermal conditions
according to the principles of the UOP Sorbex separations technology. Because the separation
takes place in the liquid phase, heating, cooling and power requirements are remarkably low.
Sorbex is the name applied to a particular technique developed by UOP for separating a
component or group of components from a mixture in the liquid phase by selective adsorption on
a solid adsorbent. In broad outline, Sorbex is a simulated moving bed adsorption process
operating with all process streams in the liquid phase and at constant temperature within the
adsorbent bed. Feed is introduced and components are adsorbed and separated from each other
within the bed. A separate liquid of different boiling point referred to as ‗desorbent‘ is used to
displace the feed components from the pores of the adsorbent. Twoliquid streams emerge from
the bed – an extract and a raffinate stream, both diluted with desorbent. The desorbent is
removed from both product streams by fractionation and is recycled to the system.
A simplified schematic flow diagram of a gasoline Molex unit is shown in Fig. 3.6. The
adsorbent is fixed while the liquid streams flow down through the bed. A shift in the positions of
liquid feed and withdrawal, in the direction of fluid flow through the bed, simulates the
movement of solid in the opposite direction. It is, of course, impossible to move the liquid feed
and withdrawal points continuously. However, approximately the same effect can be produced
by providing multiple liquid access lines to the bed, and periodically switching each net stream
to the next adjacent line.
A liquid circulating pump is provided to pump liquid from the bottom outlet to the top
inlet of the adsorbent chamber. A fluid-directing device, known as a ‗rotary valve‘, is also
provided. The rotary valve functions on the same principle as a multiport stopcock.
Figure (3.6) Block diagram of Penex/Molex process
TABLE 3.4 Typical Estimated Yields for Molex Processing
UOP offers the processes with adsorption systems on the molecular sieves in vapour
phase (Penex/Iso Siv) and liquid phase (Penex/Molex (Figure 3.9)), and process, which
combines adsorptive separation of unconverted n-paraffins from isomers and deisohexanizing
Penex/DIH/PSA.
Penex-Plus technology, which is for processing of the feed with high benzene content
(from 7 up to 30 % vol. in the case of light straight-run gasoline fraction and light reformate
blend), includes feed treatment section which is hydrogenation of benzene.
TABLE 3.5 TYPICAL PENEX ESTIMATED INVESTMENT COST
TABLE 3.6 TYPICAL PENEX ESTIMATED UTILITY REQUIRMENT
TABLE 3.7 TYPICAL PENEX ESTIMATED OPERATING REQUIRMENTS
PENEX FLOW DIAGRAM DISCRIPTION
FEEDSTOCK REQUIREMENTS
To maintain the high activity of the Penex catalyst, the feedstock must be hydrotreated.
However, costly pre-fractionation to sharply limit the levels of C6 cyclic and C7 compounds is not
required. In fact, the Penex process affords the refiner with remarkably good flexibility in the
choice of feedstocks, both at the time of design and even after the unit has been constructed. The
latter is important because changes in the overall refinery processing scheme may occur in
response to changing market situations. These changes could require that the composition of the
isomerization feed be modified to achieve optimal results for the entire refinery.
The Penex system can be applied to the processing of feeds containing up to 15 percent
C7 with minimal or no effect on design requirements or operating performance. Generally, the
best choice is to operate with lower levels of C7+ material because these compounds are better
suited for upgrading in a reforming process. Charge containing about 5.0 percent or even higher
amounts of benzene is completely acceptable in the Penex chargestock and will not produce
carbon on the catalyst. When the feed has extremely high levels of benzene, a Penex-Plus unit is
recommended. (The ―Plus‖ section can be retrofitted to an existing Penex unit should the refiner
want to process high-benzene feedstock in an existing Penex unit.) The low-octane C6 cut
recovered from raffinate derived from aromatic extraction operations typically contains a few
percent of olefins and is completely acceptable as Penex feed without pre-hydrogenation.
Sulfur is an undesirable constituent of the Penex feed. However, it is easily removed by
conventional hydrotreating. Sulfur reduces the rate of isomerization and, therefore, the product
octane number. Its effect is only temporary, however, and once it has been removed from the
plant, the catalyst regains its normal activity.
Water, other oxygen-containing compounds, and nitrogen compounds are the only
impurities normally found in the feedstock that will irreversibly poison the Penex catalyst and
shorten its life. Fresh feed and makeup hydrogen are dried by a simple, commercially proven
desiccant system.
PROCESS FLOW DIAGRAM
The UOP Penex Unit can be divided into ten sections.
A. Sulfur Guard Bed
B. Liquid Feed Driers
C. Makeup Hydrogen Driers
D. Feed Surge Drum
E. Exchanger Circuit
F. lsomerization Reactors
G. Stabilizer
H. Stabilizer Gas Scrubber
I. Separator and Compressor Section (Recycle Gas Units Only)
FIGURE 4.1 PENEX FLOW DIAGRAM
A. SULFUR GUARD BED
The purpose of the sulfur guard bed is to protect the Penex catalyst from sulfur in the
liquid feed. The hydrotreater will remove most of the sulfur in the Penex feed. The guard bed
reduces the sulfur to a safe level for H.O.T. Penex operation and serves as insurance against
upsets in the NHT which could result in higher than normal levels of sulfur in the feed. The
guard bed is loaded with UCP ADS-11 adsorbent, a nickel- containing extrudate designed to
chemisorb sulfur from the liquid feed. The feedstock is heated to the required temperature for
sulfur removal, usually 250°- (120°C) and passed down flow over the adsorbent. Once sulfur
breakthrough occurs, normally after one year or so of operation, the guard bed is taken off line
and reloaded with fresh adsorbent. The Penex Unit need not be shut down during the short period
of time required to reload the guard bed so long as the NHT is performing properly.
B. LIQUID FEED DRIERS
The liquid feed driers are used to dry the Penex liquid feed to less than 0.1 ppm H20. The
piping is designed so that either drier can be in the lead or the lag position in series flow
operation. Either drier can be operated individually while the other is being regenerated. The
driers are designed for a 48 hour cycle which includes 24 hours in the lead position, 7 hours
regenerating, and 3 hours cooling and 14 hours in the lag position. Proper drier operations are
essential in the Penex process since the catalyst is water intolerant. Typically, type 4A molecular
sieves are employed within the driers.
Charge to the liquid feed driers is hydrotreated LSR naphtha from the Naphtha
hydrotreater stripper bottoms. Additional recycle feed from a Molex unit or deisohexanizer may
be added in the future.
Liquid feed is pumped to the driers on flow control at 100°F (38°C). The low temperature
must be maintained for proper drying. Feed enters the lead drier, passes up flow through it,
crosses over to the bottom of the lag drier, passes up flow through it and is routed to the cold
combined feed exchanger in the reactor section.
The lead drier removes water from the feed to less than 0.1 wppm. The lag drier acts as a
guard bed. It reduces the water content even further but is essentially in standby until 1.0 wppm
H20 breaks through the lead drier (or until the scheduled regeneration period arrives). The water
content is continuously monitored with a Parametric moisture analyzer. This analyzer will
always be used to monitor the lead drier effluent. If the lead drier effluent reaches 1.0 wppm H20
content, it must be taken off line and immediately regenerated. The previous lag drier then
becomes the lead drier. On completion of the regeneration, series flow is re-established with the
regenerated drier now in the lag position. See the liquid feed drier regeneration procedure for
details of the regeneration.
C. MAKEUP HYDROGEN DRIERS
The makeup hydrogen driers are used to dry the Penex unit makeup hydrogen to less than
0.1 ppm H20. The function and design of these driers is very similar to the liquid feed driers.
Again, two vessels with interconnecting piping are used. the piping is designed so that in series
operation either drier can be in the lead or the lag position and so that either drier can be operated
individually. while the other is being regenerated. The driers are designed for a 48 hour cycle
with includes 24 hours in the lead position, 6 hours regenerating, 4 hours cooling and 14 hours in
the lag position. These driers, too, are essential for good Penex operations since the I-8 catalyst is
water intolerant. Type 4A molecular sieves are normally employed within the driers.
Makeup hydrogen to the driers comes from the Platforming unit at 100°F. (38°F). The
actual makeup rate will vary with chemical consumption due to benzene saturation, ring opening
and hydrocracking. The makeup gas flow rate is controlled by the H2O/HCBN ratio controller on
the fresh feedstock. The ratio is adjusted to maintain excess hydrogen at the reactor outlet. The
incoming gas passes up flow through the lead driers, crosses over to the bottom of the lag drier
and passes up flow through it to the cooler in the separator and compressor section. Like the
liquid feed driers, the lead makeup gas drier removes water from the makeup hydrogen to less
than 0.1 wppm The lag drier acts as a guard bed until 1.0 ppm H20 breaks through the lead drier.
Water content is continuously monitored with a moisture analyzer. This analyzer will always be
used to monitor the lead drier effluent. If the lead drier effluent reaches 1.0 ppm H20 contend it
must be taken off line and regenerated. The previous lag drier becomes the lead drier. On
completion of the regeneration, series flow is re-established with the regenerated drier now in the
lag position. See the makeup hydrogen direr regeneration procedure. for the details of the
regeneration.
D. FEED SURGE DRUM
The purpose of this drum is to provide liquid feed surge capacity for the Penex Unit.
Dried feed from the liquid feed driers is routed to this drum. A The feed surge drum is blanketed
with dry hydrogen gas originating from the outlet of the make-up gas driers with the feed surge
drum pressure being controlled by a PRC.
E. EXCHANGER CIRCUIT
The dried liquid feed from the feed surge drum is pumped by either of the two reactor
charge pumps through the reactor exchanger circuit on flow control. The reactor exchanger
circuit consists of the cold combined feed exchanger, the hot combined feed exchanger, and the
reactor charge heater. Prior to the entry of the liquid hydrocarbon into the cold combined feed
exchanger, it combines with the makeup hydrogen stream. After combining, the mixed
hydrocarbon-hydrogen stream passes through the exchanger circuit in the order previously
mentioned.. After the makeup gas combines with the feed a small quantity of catalyst promoter
(CCI4) is added. This promoter is pumped into the process either of the- two injection pumps.
The catalyst promoter is stored in a nitrogen blanketed storage drum. The cold combined feed
exchanger is equipped with a bypass which can be used to regulate the amount of combined feed
preheat. The bypass is regulated with a board mounted control valve. The combined feed is
finally brought up to the desired temperature in the reactor charge heater by a temperature
controller which resets the exchangers heating medium flow. The charge heater is equipped with
an automatic shutdown which is activated by low feed or low makeup gas flow. After exiting the
reactor charge heater, the heated combined stream then flows to the first reactor.
F. ISOMERIZATI ON REACTORS
The reactors are the heart of the process. The operation of them is such that a reactor will
be placed in series with the other reactor. At various times throughout the unit‘s history it will be
possible to have either reactor in the lead or tail position. Thermocouples are inserted into the
catalyst bed of each reactor to monitor the activity of the catalyst. After exiting the reactor
charge heater, the heated combined stream then flows to the first reactor. Upon exiting the first
reactor, the stream then passes to the hot combined feed exchanger where the first reactor‘s heat
of reaction is partially removed. The degree of temperature removal can be achieved by adjusting
the amount of exchanger bypassing with a temperature controller.The partially cooled stream is
then routed to the second reactor where the final process reactions are completed. The reactors
are equipped with hydrogen purge lines which are located at the inlet of each reactor. The
hydrogen purge is used to remove hydrocarbon from a reactor which is to be unloaded or to cool
a reactor during an emergency. Each purge is controlled by a board mounted flow controller.
In case of a high reactor temperature emergency the reactors are equipped with
depressuring lines to the flare system. The reactors are depressured from the outlet of the lag
reactor. The depressuring line is equipped with two motorized valves which can be operated
from the control room. After exiting the second reactor, the stream is then routed to the tube side
of the cold combined feed exchanger. The cold combined feed exchanger tube side effluent is
then routed to the stabilizer on pressure control.
G. STABILIZER
The purpose of this column is to separate any dissolved hydrogen, HCl and cracked gases
(C1, C2, and C3‘s) from the isomerate.The feed to this column is routed hot directly from the cold
before entering the stabilizer.The column is reboiled by either steam or hot oil. The reboiler heat
input is controlled by a FRC on the heating medium. The stabilizer column overhead vapor,
consisting of the light hydrocarbon components of the column‘s feed, is routed to an air or water
cooled condenser and then to the stabilizer receiver. To maintain pressure control on the column,
gas is vented on pressure control to the stabilizer gas scrubber. Liquid is pumped from the
receiver on level control with the stabilizer reflux pump. All liquid from the stabilizer overhead
receiver is refluxed to the column on tray No. 1. Bottoms product is routed to storage on level
control after first being cooled in the stabilizer bottoms cooler. If the stabilizer bottoms is sent to
a Deisohexanizer it is not cooled, but is charged hot to the column. Part of the stabilizer bottoms
is used for regenerating the driers.
H. STABILIZER GAS SCRILIBBER
The stabilizer off gas flows up flow through the stabilizer gas scrubber to remove
hydrogen chloride. The scrubbed gas leaves the top of the vessel and goes to fuel gas on
backpressure control. The hydrogen purity is monitored on the scrubbed off gas to determine the
moles of H2 leaving the system for the H2/C:H determination. Make-up caustic is pumped from
the refinery to the reservoir section of the gas scrubber when caustic addition is required. The
caustic in the reservoir section is pumped by the caustic recirculating pumps to the top of the
scrubbing section of the scrubber where a counter current contact with the rising acidic gas is
made. Caustic is also continuously circulated to the distributor under the packed section. The
flow rate of the circulating caustic can be monitored by a local flow indicator. Periodically a
portion of the caustic is withdrawn to the refinery spent caustic facilities as spent caustic. The
caustic level in the scrubber is maintained about 1-2 feet below the distributor under the packed
section.
I. SEPARATOR AND COMPRESSOR SECTION (Recycle Gds Units Only)
The separator and compressor section separates the reactor effluent into unstabilized
liquid product and recycle gas. The separator pressure is controlled by regulating the makeup
hydrogen flow rate. The equipment in this section is: the reactor product condenser, the product
separator, the recycle gas compressor, and if required, the make-up gas compressor suction
drum, and the make-up gas compressor.
In a Hydrogen Once Through Unit the product condenser, product separator and recycle
compressor are not used. In this unit the pressure in the reactor circuit is controlled using a back
pressure value on the stabilizer feed line.
Reactor effluent exits the reactor section and is partially condensed in the reactor product
condenser. It cools the effluent to about 100°F. The cooled liquid and gas then separate in the
product separator. Unstabilized liquid product is pressured out of the product separator on level
control to the stabilizer section. Recycle gas exits from the separator and goes through the
recycle gas compressor to the cold combined feed exchanger in the reactor section. Recycle gas
flow is controlled by a flow indicating controller which spills back to the product separator
(through the reactor product cooler), thereby controlling the amount sent forward. Recycle gas
flow and purity are controlled to maintain hydrogen to hydrocarbon mole ratio of about 2:1. Dry
make-up hydrogen, from the make-up hydrogen drier section, combines with spillback from the
make-up gas compressor. These gases pass through the make-up gas cooler and into the makeup
gas compressor suction drum. Any entrained hydrocarbons are knocked out and are manually
drained to an appropriate location. The make-up gas is compressed and combined with the
recycle gas to the reactor section. Make-up gas flow is controlled by the product separator
pressure recorder controller.
PROCESS CHEMISTRY
REACTION MECHANISMS
Paraffin isomerization catalysts fall mainly into either of two principal categories: (1)
those based on Friedel-Crafts catalysts as classically typified by aluminum chloride/hydrogen
chloride, or (2) dual-function hydro-isomerization catalysts. No attempt is made to present a
discussion of mechanisms of a degree of sophistication acceptable to a chemist specializing in
the area. The intention is simply to provide those practicing engineers who have not previously
had reason to consider isomerization with a basic introduction to the subject. Isomerization by
either Friedel-Crafts or duals function catalysts is generally thought to entail intramolecular
rearrangements of carboniumions as illustrated - for pentane:
(1) CH3-CH-CH2-CH2-CH3
CH3-C2H3-CH2-CH3
There appear to be two schools of thought regarding the Friedel-Crafts mechanism.
Perhaps each mechanism is operative and the disagreement is merely over their relative
importance under specific circumstances.
Friedel-Crafts isomerization is believed by some to require the presence of traces of olefins
or alkyl halides as carbonium ion initiators, with the reaction thereafter proceeding through
chain propagation. The initiator ion, which needs to be present in small amounts only, may be
formed by the addition of HCl or HAlCI4 to an olefin, which is present in the paraffin as an
impurity or which is formed by cracking of the paraffin:
(2) RCH=CH2 + HAlCl4
RCHCH3 + AlCl4
The initiator then forms a carbonium ion from the paraffin to be isomerized:
(3)
RCHCH3 + CH3-CH2-CH2-CH2- CH3
RCH2CH3 + CH3-CH-CH2CH2-CH3
Skeletal rearrangement then occurs:
(4) CH3-CH-CH2CH2-CH3
CH3-C-CH3-CH2-CH3
Isopentane is then formed and the chain propagated by the generation of a New normal
carbonium ion:
(5) CH3-CCH3-CH2CH3 + CH3-CH2-CH2-CH2-CH3
CH3-CH-CH2-CH3 + CH3-CH-CH2CH2-CH3
Naturally, the same sequence could have been illustrated starting with Isopentane and
ending with n-pentane and an iso-carbonium ion to propagate The chain, i.e. reactions (3), (4),
and (5) are reversible, as are all of the Reactions to be shown later. The composition of the final
mixture is, of course, that set by thermo-dynamic equilibrium, assuming that sufficient reaction
time has been provided. Another Friedel-Crafts route which has been suggested is direct hydride
ion abstraction:
(6) CH3-CH2-CH2-CH2-CH3 + AlCl3
CH3-CH-CH2-CH2-CH3 + HAlCl3
The carbonium ion, as before, rearranges
(7) CH3-CH-CH2-CH2-CH3
CH3-C -CH2-CH3
Finally, iso-pentane is formed:
(8) CH3-C-CH2-CH3 + HAlCl3
CH3-CH-CH2-CH3 + AlCl3
Abstraction of the hydride ion is energetically favored by the fact that the aluminum atom
can thereby complete its electron octet. Since there is always some hydrogen chloride present,
either by design or from hydrolysis of aluminum chloride by traces of water, a Bronsted
(protonic) acid could have been shown for Reactions (6) and (8) instead of a Lewis acid:
CH3-CH-CH2-CH2-CH3 (AlCl4)- +H2
(9) CH3-CH2-CH2·CH2CH3 + H + AlCl4
Some chemists feel uncomfortable with the above because of the required postulation of
hydrogen formation. The dual-function hydro-isomerization catalysts are thought by some to
operate through an olefin intermediate whose formation is catalyzed by the metallic component,
assumed for illustration purposes to be platinum: (10) CH3-CH2-CH2CH2CH3
CH3-CH2CH2CH=CH2 + H2
This reaction is, of course, reversible and, since these catalysts are employed under
substantial hydrogen pressure, the equilibrium is far to the left. However, the acid function of the
catalyst consumes the olefin by formation of a carbonium ion and thus permits more olefin to
form despite the unfavorable equilibrium. This step is entirely analogous to Reaction (2) shown
for Friedel- Crafts, except that it is better to denote the acid function by a more general.
(11) CH3-CH2-CH2—CH=CH2 + H + A-
CH3-CH2-CH2-CH-CH3 + A-
The usual rearrangement ensues:
(12) CH3-CH2-CH2-CH-CH2
CH3-CH2CCH3-CH3
The is olefin is then formed by the reverse analogue of (11):
(13) CH3-CH2C CH3-CH3 + A-
CH3-CH2-C=CH2 + H + A-
The iso-paraffin is finally created by hydrogenation:
(14) CH3-CH2-CCH3=CH2 + H2
CH3-CH2-CHCH3-CH3
Those dual—functi0nal hydro-isomerization catalysts which operate at very low
temperatures have stronger acid sites than those which require higher temperatures. In this case it
is possible that the necessary carbonium ion is formed by direct hydride ion abstraction from the
paraffin by the acid function of the catalyst:
(15) CH3-CH2-CH2CH2CH3 + H+A(16) CH3-CH-CH2-CH2-CH3
(17) CH3-CCH3-CH2·CH2 + A-+H2
CH3-CHCH2-CH2-CH3 + A- + H2
CH3-CCH3-CH2-CH3
CH3-CH-CH2-CH3 + H+A
The last reaction is in lieu of the displacement type chain propagation step (Reaction 5)
discussed earlier. Since the reaction with hydrogen is relatively fast, acid sites are readily
liberated for further reaction. This may account, at least in part, for the higher activity of such
dual functional catalysts.
Equilibrium limits the maximum conversion possible at any given set of conditions. This
maximum is a strong function of the temperature at which the conversion takes place. A more
favorable equilibrium exists at lower temperatures. Figure 5.1 shows the equilibrium plot for the
pentane system. The maximum isopentane content increases from 64 mol % at 260°C to 82 mol
% at 120°C (248°F). Neopentane and cyclopentane have been ignored because they seem to
occur only in small quantities and are not formed under isomerization conditions.
The hexane equilibrium curve shown in Fig. 5.2 is somewhat more complex than that
shown in Fig. 5.1. The methylpentanes have been combined because they have nearly the same
octane rating. The methylpentane content in the C6-paraffin fraction remains nearly constant over
the entire temperature range. Similarly, the fraction of 2,3-dimethylbutane is almost constant at
about 9 mol % of the C6 paraffins. Theoretically, as the temperature is reduced, 2,2dimethylbutane can be formed at the expense of normal hexane. This reaction is highly desirable
because nC6 has a RON of 30. The RON of 2,2-dimethylbutane is 93. Of course, the petroleum
refiner is more interested in octane ratings than isomer distributions.
Figure 5.3 shows the unleaded research octane ratings of equilibrium mixtures plotted
against the temperature characteristic of that equilibrium for a typical Chargestock. Both the C5
and the C6 paraffins show an increase in octane ratings as the temperature isreduced.
FIGURE 5.1. C5 Paraffin Equilibrium Plot.
FIGURE 5.2. C6 Paraffin Equilibrium Plot.
FIGURE 5.3 Unleaded RON Ratings Of Equilibrium Fractions.
Because equilibrium imposes a definite upper limit on the amount of desirable branched
isomers that can exist in the reactor product, operating temperatures are thought to provide a
simple basis for catalyst comparison or classification. However, temperature is only an
approximate comparison that at best can discard a catalyst whose activity is so low that it might
be operated at an unfavorably high temperature. Further, two catalysts that operate in the same
general low-temperature range may differ in the closeness with which they can approach
equilibrium in the presence of reasonable amounts of catalyst.
B. ALUMINIUM CHLORIDE
The isomerization catalysts employed during World War Il were all of the Friedel-Crafts
type. Those which contained aluminum chloride only were either a hydrocarbon/aluminum
chloride complex (the so-called sludge process) or they were manufactured in deposition onto a
support such as alumina or bauxite. They were intended to operate at very low temperatures
(120-2650F) and to approach the very favorable equilibrium composition characteristic of these
temperatures.
The catalyst tended to consume itself by reaction with the feedstock and/or product.
When temperature was raised a little in an effort to compensate for loss of catalyst and to speed
the reaction to effect more isomerization, light fragments were formed by cracking and these,
when vented, caused an excessive loss of the HCl promoter.
Corrosion of downstream equipment was also commonplace, due to the solubility of
aluminum chloride in hydrocarbon, to its relatively high volatility and to the difficulty of
removing it from the product by caustic washing. Aluminum chloride deposition in and plugging
of reboiler tubes was not uncommon. The process faced problems in sludge disposal which were
considered onerous even before the present acute awareness of environmental factors developed,
The fixed bed process sometimes experienced unpredictable amounts of isomerization.
C. HYDRO-ISOMERIZATI ON CATALYSTS (ABOVE 390 0F)
The operational problems which had characterized the wartime Friedel- Crafts type
isomerization plants, the advent of catalytic reforming which not only made hydrogen generally
available in refineries but also demonstrated the practicality of using noble metal containing
catalysts on a large scale, and the octane number race which postwar high compression engines
initiated all combined in the 1950‘s to spawn a spate of hydro-isomerization processes. These
catalysts generally contained a noble metal and some halide, operated at temperatures between
about 560°F and temperatures approaching those characteristic of catalytic reforming, employed
recycle hydrogen to prevent catalyst carbonization and utilized either no promoter or traces at
most. In general, they did not require an especially dry feedstock but did benefit from a low
sulfur content feedstock. Most achieved a close approach to the equilibrium characteristic of
their particular operating temperature. Because of their high operating temperatures and their
necessarily low conversions to iso-paraffins, these high temperature catalysts were quickly
replaced with the advent of the "third generation" low temperature catalysts.
D. HYDRO·ISOMERIZATI ON CATALYSTS (BELOW 390 °F)
―Low temperature‖ is considered rather arbitrarily for catalyst classification purposes as
anything below 390°F operating temperature. Typically these are fixed bed catalysts containing a
supported noble metal and a component to provide acidity in the catalytic sense. They operate in
a hydrogen atmosphere and may employ a catalyst promoter whose concentration in the reactor
may range from parts per million to substantially higher levels. They generally all require a dry,
low sulfur feedstock; however, they may differ importantly in their tolerance of certain types and
molecular weights of hydrocarbons. Hydrocracking to light gases is generally slight, so liquid
product yields are high. The type of catalyst used in the Penex unit is of this type. Apart from the
paraffin isomerization reactions which were discussed in detail in the proceeding pages, there are
several other important reactions including:
1. Naphthenes Ring Opening:
Penex feeds can contain up to 30% naphthene rings. The three Naphthenes which are
typically present in Penex feed are cyclopentane (CP), methyl cyclopentane (MCP) and
cyclohexane (CH). The naphthene rings will hydrogenate to form paraffins. This ring opening
reaction increases with increasing reactor temperature. At typical Penex reactor conditions, the
conversion of naphthene rings to paraffins will be on the order of 20-40 percent.
2. Naphthenes Isomerization:
The Naphthenes MCP and CH exist in equilibrium. Naphthene isomerization will shift
towards MCP production as temperature is increased.
3. Benzene Saturation:
Penex feeds can contain up to 4% benzene. The catalyst will saturate benzene to
cyclohexane. The catalyst-will saturate benzene to cyclohexane. This reaction proceeds very
quickly and is achieved at very low temperatures. Saturation of benzene is not equilibrium
limited at Penex conditions and conversion will be 100%. The saturation of benzene produces
heat. This heat generation limits the amount of benzene which can be tolerated in the Penex feed.
The platinum function on the Penex catalyst is responsible for benzene saturation. .
Hydrocracking; Hydrocracking occurs in the Penex reactor to a degree which depends on the
feed quality and severity of operation. Large molecules such as C7‘s tend to hydrocrack more
easily than smaller molecules. C5 and C5 paraffins will also hydrocrack to a certain extent. As
C5/C7 paraffin isomerization approaches equilibrium, the extent of hydrocracking increases. If
isomerization is pushed too hard, hydrocracking will reduce the liquid yield and increase heat
production. Methane, ethane, propane and butane are produced as a result of hydrocracking.
The various Penex Unit reactions are illustrated as under:
PARAFFIN ISOMERIZATION
NORMAL HEXAHE
CH3-CH2-CH2-CH2-CH2-CH3
24.8 ron
2 METHYL PENTANE
CH3
CH3-CH-CH2-CH2-CH3
0 73.4 ron
3 METHYL PENTANE
CH3
CH3-CH2-CH2-CH2-CH2-CH3
24.8 ron
CH3-CH2-CH-CH2-CH3
74.5 ron
2-2 DIMETHYL BUTANE
CH3
CH3-CH2-CH2-CH2-CH2-CH3
24.8 ron
CH3-C-CH2-CH3
CH3
91.8 ron
2-3 DIMETHYLBUTANE
CH3
CH3-CH2-CH2-CH2-CH2-CH3
24.8 ron
CH3-CH-CH-CH3
CH3
104.3 ron
PARAFFIN ISOMERIZATION
NORMAL PENTANE
CH3-CH2-CH2-CH2-CH3
61.8 ron
ISO PENTANE
CH3-CH-CH2-CH3
CH3
93.0 ron
HYDROCRACKING
NORMAL HEPTANE
PROPANE + BUTANE
CH3
CH3-CH2-Ch2-CH2-CH2-CH2-CH3 + H2
CH3-CH2-CH3 + CH3-CHCH3
ACIDIZING
Hydrogen
Iron
Iron
Chloride + Oxide
Chloride + Water
6HCl + Fe203
2FeCl3 + 3H2O
CHLORIDE PROMOTER
Carbon
Tet. + Hydrogen
Hydrogen
Chloride + Methane
CATALYST
CCI4 + 4H2
HEAT
4HCl + CH4
Perchloroethylene + Hydrogen
Ethane
Hydrogen Chloride +
CATALYST
C2Cl4 + 5H2
HEAT
4HCl + C2H6
CAUSTIC SCRUBBING
HCI + NaOH
NaCl + H2O
HZS + 2NaOH
Na2S + 2H2O
HZS + Na2S
2NaHS
HCI + Na2S
NaCl + NaHS
HCl + NaHS
NaCl + H2S
PROCESS VARIABLES
In the normal operation of a Penex unit having once set the operating pressure, fresh feed
rate and the recycle hydrogen and makeup hydrogen flows, it is usually only necessary to adjust
the reactor inlet temperatures. Nevertheless, it is to the operator‘s advantage that he has a
thorough understanding of the influence process variables will have on performance of the unit
and the life of the catalyst. Once the catalyst has been loaded into the unit, the manner in which
the catalyst is placed in service and the treatment it receives when in service will to a large extent
influence its effectiveness for making quality product as well as the length of service it will give.
In making any changes to the operation, the welfare of the catalyst must be given prime
consideration for it can be regarded as the heart of the operation on which the quality of the
results obtained will depend.
FEED FRACTIONATION
The economics of operating an isomerization unit are impacted by the feedstock
composition and the cut point between the Penex isomerization feedstock and the heavy naphtha
which is usually processed in a catalytic reformer. The prime consideration in splitting the
feedstock is to maximize the C5/C6 paraffins to the Penex unit keeping the C7+ material in the
naphtha splitter bottoms which will be processed in the catalytic reformer. A complication arises,
however, in the fact that the C6 cyclics (methylcyclopentane, cyclohexane and benzene) are also
present in the splitter feed and adecision on which unit to send them to must be made. Ideally,
the C6 cyclics which are high octane benzene precursors would be best processed in the reformer.
Some insight can be obtained by reviewing the accompanying figures regarding the component
boiling points and volatilities. It is observed that the benzene and normal hexane have essentially
the same volatility and, hence, cannot be separated by fractionation. Therefore, the decision must
be made whether to take both the normal hexane and benzene overhead in the naphtha splitter or
to send both components out the splitter bottoms to the reformer. By observing the conversion of
normal hexane in both processes, it is seen that a catalytic reformer leaves a significant level of
normal hexane in the C6 product. The question then becomes which is better to leave, low octane
n-C6 in the reformate or to sacrifice it to cyclohexane. In general, the best overall octane
(Penexate plus reformate) can be obtained by including the n-C6 and benzene in the Penex
feedstock., Most of the cyclohexane should be fractionated out of the Penex feed and sent to the
reformer. Hence, the Penex feed should include all of the C6 paraffins plus benzene and some
MCP and the reformer feed should contain most of the cyclohexane and C7+. This feed
preparation philosophy may shift as the percentage levels of MCP plus benzene increase in the
C6 fraction and as the reforming selectivity in converting n-C6 to benzene increases.
For those operations that process reformates in an aromatics recovery unit such as Sulfolane or
Udex, the C6 cyclics and n-C6 should be included in the reformer feedstock. The unconverted C6
paraffins may then be fractionated from the light raffinate and upgraded in the Penex unit.
In the discussion which follows, the "product isomer ratio‖ refers to either the percentage ratio
of isopentane to total C5 aliphatic paraffins, the percentage ratio of 2,3 dimethylbutane in the
total C6 aliphatic paraffins, or the percentage .ratio of normal hexane in the C6 aliphatic paraffins
in the stabilizer bottoms stream, The terms ―liquid feed," ―reactor charge" and "combined feed,‖
refer to the C5/C6 charge to the liquid feed dryers of the unit, the effluent from the liquid feed
dryers and the reactor charge plus makeup hydrogen and recycle hydrogen gas, respectively.
REACTOR TEMPERATURE
In general, reactor temperature is the main process control. A definite upper limit exists
for the amount of iso-paraffins which can exist in the reactor product at any given outlet
temperature.This is the equilibrium imposed by thermodynamics, and it can be reached only after
an infinite length of time, i.e. with an infinitely large reactor. in practice, therefore, the product
will contain slightly less iso-paraffins than this equilibrium concentration. As the reactor
temperature is raised to increase the rate of isomerization, the equilibrium composition will be
approached more closely. At excessively high temperatures, the concentration of iso-paraffins in
the product will actually decrease because of the downward shift in the equilibrium curve, even
though the high temperature gives a higher reaction rate.
In the case of the Hydrogen Once Through Penex designs, where the hydrogen content
of the reactor charge is so much lower than for recycle gas Penex, the reactor feed will contain a
considerable amount of liquid. The equilibrium concentration for isoparaffins in the liquid phase
is lower than in the vapor phase. This is a thermodynamic phenomenon related to the lower
Gibbs free energies for the components in the liquid phase. The attached figures illustrate the
equilibrium for both the vapor and liquid phases for isopentane and 2,2 DMB. The use of
temperatures higher than necessary to achieve a reasonable close approach to equilibrium
accomplishes neither other than to increase the amount of hydrocracking. Extremely high,
temperatures may lead to an increased rate of carbon laydown on the catalyst; however, the
carbon forming propensity of the catalyst is inherently so low that excessive hydrocracking
would normally be encountered before carbon formation problems would develop. It is
recommended, however, that UOP be consulted before temperatures above about 380°F are
employed.
A typical C5/C6 Penex unit is provided with two reactors in series with provision for
independent temperature control. in The first reactor system affects the bulk of the isomerization,
so long as most of the catalyst therein is still active. All of the benzene in the feed is
hydrogenated in the first reactor, even when the catalyst therein has lost its activity with respect
to paraffin isomerization. Some conversion of cyclohexane and methyl cyclopentane to hexanes
also occurs, as does some hydrocracking of C7 to C3 and C4. These three reactions (benzene
hydrogenation, naphthene conversion to hexane, and C7 hydrocracking) are exothermic and, for a
typical feedstock, contribute more to the temperature rise in the first reactor than does paraffin
isomerization, which is also exothermic.
Normally, the first reactor will be operated at such a temperature as to maximize the
concentration of isopentane and 2,2 dimethyl butane in its effluent. The concentrations attainable
and-the required outlet temperature will be influenced by the amount of active catalyst present
and by the amount of C6 cyclic and C7 components present in the feed, higher temperatures being
required with high concentrations of these components in the feed. By this procedure, the
required operating temperature for the second reactor is reduced and it is possible to operate
under conditions where the equilibrium is more favorable.
The optimum reactor temperatures are determined in the field by establishing a "base‖ set
of conditions and then varying the reactor temperatures, one at a time, from this base condition.
The performance of the reactors is determined by calculating the ―iso·ratios‖ of each reactor
effluent. Reactor effluent or product octane is not used for this determination since it is also
dependent on the relative amounts of C5 and C6 in the feedstock.
Another interaction may occur with feeds which are rich in C5 cyclics. Since these
materials tend to reduce the rate of paraffin isomerization, it may be beneficial with very rich
feed stocks to choose the first reactor system temperature to control the amount of` cyclics which
enter the second reactor system. By raising the first reactor system temperature, more of the
cyclics be converted to hexanes and the rate of isomerization in the second reactor system
thereby increased.
LIQUID HOURLY SPACE VELOCITY
This term, commonly shortened to LHSV, is defined as the volumetric hourly flow of
reactor charge divided by the volume of catalyst contained in the reactors in consistent units. The
design LHSV for C5/C5 Penex operation is- normally 1 to 2 and increasing the LHSV beyond
this could lead to lower product isomer ratios.
HYDROGEN TO HYDROCARBON MOL RATIO
For Hydrogen Once Through Penex units, this ratio is defined as the number of mols
hydrogen at the reactor outlet per mol of reactor charge passing over the catalyst and is specified
at 0.05 mols H2/mol H:C. The primary purpose of maintaining the ratio at or above the design is
to avoid carbon deposition on the catalyst and maintain enough H2 for the reactions to proceed. lf
necessary, the reactor charge rate is to be reduced to maintain the design hydrogen to
hydrocarbon ratio. The H2/H:C ratio is determined by measuring the total moles of hydrogen in
the stabilizer overhead gas and dividing by the total moles of fresh feedstock.
For Recycle Gas Penex units, the hydrogen to hydrocarbon ratio is specified at 1 to 2.
This ratio is defined as the moles of hydrogen per mole of reactor charge. Lower ratios will
generally increase the amount of liquid phase material in the reactor and may lead to flow
distribution.
PRESSURE
C5/C5 Penex units are normally designed to operate at 450 psig at the reactor outlet.
Methylcyclopentane and cyclohexane appear to adsorb on the catalyst and reduce the rate of
isomerization reactions. Higher pressure helps to offset this effect of the C6 cyclic compounds.
Lowering the unit pressure or operating at a slightly lower level would not affect the catalyst life
but the extent of isomerization would be influenced.
CATALYST PROMOTER
To sustain catalyst activity, the addition of chloride is necessary. At no time should the
plant be operated for longer than six hours without the injection of chloride. Whenever there is-a
catalyst chloride decency, the product isomer ratios will decrease (although not necessarily
instantaneously), other things being equal. Restarting the injection of chloride will tend to return
the activity of the catalyst to its previous level, but it is possible that full activity will not be
restored if a decline in activity, as a result no chloride injection has been observed. Carbon
tetrachloride and specific grades of perchloroethylene are the only approved sources of chloride.
CATALYST:
CATALYSTS OF HYDROISOMERIZATION PROCESS:
The major advantage of this catalyst was its low temperature activity (T< 200°C) due to
its high acidity. However the catalysts were sensitive towards water and oxygenates and in
addition had corrosive properties. Furthermore, chlorine addition during the reaction is necessary
to guarantee catalyst stability.
In the Hysomer process zeolite based catalysts were used which had the major advantage
of resistance to feed impurities. Industrially applied zeolites used today are Pt-containing,
modified synthetic (large-port) mordernite e.g. HS10 of UOP, or HYSOPAR from Sud- Chemie.
As higher hydrogen to hydrocarbon ratios are needed recycle compressors and separators are
required for this technology.
The isomerization of hydrocarbons < C6 is currently carried out very successfully using
bifunctional supported platinum catalysts. However, difficulties are encountered with
hydrocarbons larger than hexane since the cracking reactions become more significant over
platinum catalysts as the chain length increases. Catalysts used in state of the art isomerizationcracking reactors are bifunctional. They have a metal function providing de-hydrogenation and
hydrogen activation properties that are usually supplied by group VIII noble metals like Pt, Pd,
Ni or Co. The acid function is the support itself and some examples include acid zeolites,
chlorided alumina and amorphous silica alumina. Noble metals have a positive effect on the
activity and stability of the catalyst. However they have a low resistance to poisoning by sulfur
and nitrogen compounds present in the processed cuts.
In order to prepare a suitable catalyst for hydroconversion of alkanes, good balance
between the metal and acid functions must be obtained. Rapid molecular transfer between the
metal and acid sites is necessary for selective conversion of alkanes into desirable products.
Two of the attractive features of zeolite are that the catalysts are tolerant of contaminants
and that they are regenerable. The chlorinated alumina catalysts are very sensitive to
contaminants such as water, carbon oxides, oxygenate, and sulfur. Thus, feeds and hydrogen
must be hydrotreated and dried to remove water and sulfur. Furthermore, the chlorinated alumina
catalysts require the addition of organic chloride to the feed in order to maintain their activities.
This causes contamination in the waste gas of hydrogen chloride, a scrubber is needed to remove
such contamination.
The UOP BenSat process uses a commercially proven noble metal catalyst, which has
been used for many years for the production of petrochemical-grade cyclohexane. The catalyst is
selective and has no measurable side reactions. Because no cracking occurs, no appreciable coke
forms on the catalyst to reduce activity. Sulfur contamination in the feed reduces catalyst
activity, but the effect is not permanent. Catalyst activity recovers when the sulfur is removed
from the system.
ALUMINA
Alumina or aluminum oxide (Al2O3) is a chemical compound with melting point of about
2000°C and sp. gr. of about 4.0. It is insoluble in water and organic liquids and very slightly
soluble in strong acids and alkalies. Alumina occurs in two crystalline forms. Alpha alumina is
composed of colorless hexagonal crystals with the properties given above; gamma alumina is
composed of minute colorless cubic crystals with sp. gr. of about 3.6 that are transformed to the
alpha form at high temperatures. Figure 8.1 shows the shape of (Al2O3).
FIGURE 8.1. The shape of aluminium oxide
Identifiers
Aluminium oxide
The most common form of crystalline alumina, α-aluminium oxide, is known as
corundum. If a trace of the element is present it appears red, it is known as ruby, but all other
colorations fall under the designation sapphire. The primitive cell contains two formula units of
aluminium oxide. The oxygen ions nearly form a hexagonal close-packed structure with
aluminium ions filling two-thirds of the octahedral interstices.
TYPICAL ALUMINA CHARACTERISTICS INCLUDE:







Good strength and stiffness
Good hardness and wear resistance
Good corrosion resistance
Good thermal stability
Excellent dielectric properties (from DC to GHz frequencies)
Low dielectric constant
Low loss tangent
ZEOLITE
Zeolites are microporous crystalline solids with well-defined structures. Generally they
contain silicon, aluminium and oxygen in their framework and cations, water and/or other
molecules within their pores. Zeolites occur naturally as minerals or synthetic, Figure (2.6)
shows the shape of different types of zeolites.
Because of their unique porous properties, zeolites are used in a variety of applications
with a global market of several million tonnes per annum. In the western world, major uses are in
petrochemical cracking, ion-exchange (water softening and purification), and in the separation
and removal of gases and solvents. Other applications are in agriculture, animal husbandry and
construction. They are often also referred to as molecular sieves.
Zeolites have the ability to act as catalysts for chemical reactions which take place within
the internal cavities. An important class of reactions is that catalysed by hydrogen-exchanged
zeolites, whose framework-bound protons give rise to very high acidity. This is exploited in
many organic reactions, including crude oil cracking, isomerisation and fuel synthesis.
FIGURE 8.2. Structures and dimensions of different types of zeolite
Underpinning all these types of reaction is the unique microporous nature of zeolites,
where the shape and size of a particular pore system exert a steric influence on the reaction,
controlling the access of reactants and products. Thus zeolites are often said to act as shapeselective catalysts. Increasingly, attention has focused on fine-tuning the properties of zeolite
catalysts in order to carry out very specific syntheses of high-value chemicals e.g.
pharmaceuticals and cosmetics.
The following properties make zeolites attractive as catalysts, sorbents, and ionexchangers:
1.
2.
3.
4.
5.
6.
Well-defined crystalline structure.
High internal surface areas (>600 m2/g).
Uniform pores with one or more discrete sizes.
Good thermal stability.
Highly acidic sites when ion is exchanged with protons.
Ability to sorb and concentrate hydrocarbons.
The tetrahedral arrangements of [SiO4]-4 and [AlO4]-5 coordination polyhedra create
numerous lattices where the oxygen atoms are shared with another unit cell. The net negative
charge is then balanced by cations (e.g. K+ or NH4+). Small recurring units can be defined for
zeolites named, ‗secondary building units.
The primary building blocks of all zeolites are silicon Si+4 and aluminum Al+3 cations
that are surrounded by four oxygen anions O-2. This occurs in a way that periodic three
dimensional framework structures are formed, with net neutral SiO2 and negatively charged
AlO2.
The negative framework charge is compensated by cation (often Na+) or by proton (H+)
that forms bond with negatively charged oxygen anion of zeolite.
The secondary building blocks differ between different types of zeolites. In the top line of
Figure (8.2) the structure of a faujasite type zeolite is shown. The secondary building block of
this zeolite is a sodalite cage, which consists of 24 tetrahedral in the geometrical form of a cubooctahedron. The sodalite cages are linked to each other via a hexagonal prism.
KINETIC ANALYSIS
The main aim of the present study is to analyze the kinetics of hydroisomerization
process by assessing the effect of reaction time and reaction temperature on the performance of
the catalysts. The process feed involves light naphtha which contains many reactions.
Therefore, the hydroisomerization reaction has three stages as follows:
1- Adsorption of n- paraffin molecule on dehydrogenation- hydrogenation site followed by
dehydrogenation to n- olefins.
2- Desorption of n- olefin from the dehydrogenation sites and diffusion to a skeletal rearranged
site, which converts n- olefin into iso- olefin.
3- Hydrogenation of iso- olefin into iso- paraffin molecule.
In general, the hydroisomerization of n- paraffin can occur through the bifunctional
scheme shown below:
The hydroisomerization process of light naphtha is regarded as one of the complex
chemical reactions network, where such types of reactions take on a metal and acid sites of
catalysts.
Therefore, the mathematical modeling of the hydroisomerization process is a very
important tool in petroleum refining industries. It translates experimental data into
parameters used as the basis of commercial reactor process optimization.
In the hydroisomerization of alkanes it is supposed that the alkane is dehydrogenated to
an alkene on the metal site. The alkene is then protonated on the acid site to a carbenium ion,
which is subsequently isomerized to a branched carbenium ion. The branched carbenium ion
gives the proton back to the acid site, the resulting branched alkene is hydrogenated on the
metallic site. The branched alkane is formed, and can be desorbed from the catalyst surface. The
reaction mechanism scheme is shown in Figure (4.1)
The general reactions mechanism for isomerization of n-alkane
ASSUMPTION:
The catalytic hydroisomerization kinetic the following assumptions are taken into
account:
1. The system is isothermal and in steady state operation with first order reactions.
2. The reaction is carried out in the gas phase with constant physical properties and without
pressure drop.
3. The temperature and concentration gradients in the radial direction can be neglected.
The objective of kinetic study is to construct from the experimental results of the process,
a mathematical formulation that can be used to predict the kinetic parameters of the
hydroisomerization process. Therefore, the main aim of the present work is to estimate the
reaction parameters (reaction rate constant, activation energy and pre-exponential factor)
depending on the experimental work results under real isomerization conditions.
In present work, it is suggested kinetic model for the reactions of hydroisomerization for
light naphtha (n-paraffin) can be considered by the following scheme depending on the present
model assumptions which can be formulated to the following equations:
The Suggested Reactions Of Light Naphtha Isomerization
Let,
CA denotes the mole fraction of n-paraffin present at any time t,
CN the mole fraction of n-olefin,
Ciso the mole fraction of i-paraffin.
Then,
The mole balance can be formulated mathematically as follows:
Equation (1)
Equation (2)
By integration of equation (1)
CA = CA°
CA = CA°exp(- k1t)
at t= 0
we get,
Equation (3)
Substituting the equation (3) in equation (2) yield:
= k1CA°exp (- k1t) - k2CN
Rearrangement of equation (4) gives:
+ k2CN = k1CA°exp (- k1t)
Equation (4)
This is a linear first order differential equation as follows:
𝑑𝑦
+Py = Q
where P = k2 ,
𝑑𝑥
where 𝜌 integration factor which can be calculated from:
𝜌=𝑒
𝑃𝑑𝑥
𝜌=𝑒
𝑘2𝑑𝑡
where integration factor is : exp ( k2t)
now multiple equation (4) with integration factor
𝑑𝐶 N
𝑑𝑡
𝑑
exp ( k2t) + k2 exp ( k2t) CN = k1CA°exp (- k1t) exp ( k2t)
( k t)
[C
]
=
k1CA°exp (k - k t)
N exp
𝑑𝑡
2
2
1
Then by integrate of differential equation will give:
Q = k1CA°exp (- k1t)
𝑑[CN exp ( k2t)] = k1CA°
𝑒𝑥𝑝(k2- k1t) dt
[CN exp ( k2t)] = exp (k2-k1) t + A
Equation (5)
k2 - k1
where A is the integration constant, and it can be determined using the following
conditions:
t=0 ,
CN = 0
Thus :
[(0) exp ( k2( 0))] - exp (k2-k1) (0) = A
k2 - k1
Equation (6)
CN exp (- k2t) =
[exp (k2-k1) t - 1]
Then,
CN =
[exp (- k1t) - exp (- k2t) ]
Equation (7)
But, all products come from initial n-paraffin in the light naphtha feed, then,
CAO = CA+CN+Ciso
Equation (8)
Then substituting the equations (3) and (7) in equation (8), will give:
CAO = CAO exp(- k1t) +
[exp (- k1t) - exp (- k2t) ] + Ciso
Equation (9)
Rearrangement of equation (9) gives:
Ciso = CAO- CAO exp(- k1t) -
Ciso = CAO [1- exp(- k1t) -
[exp (- k1t) - exp (- k2t) ]
[exp (- k1t) - exp (- k2t) ]
Equation (10)
REACTOR MODEL
To develop a reaction model for an integral reactor, a material balance is made over the
cross section of a very short segment of the tubular catalyst bed, as shown in Figure (.2):
SEGMENT OF PACKED BED REACTOR.
The tubular-flow reactor is one in which there is no mixing in the direction of flow and
complete mixing perpendicular to the direction of flow. Above figure represents such a reactor
.Concentrations will vary along the length coordinate , z , but not radial coordinate ,r .We
conclude that the rate of reaction will vary with reaction length.
A steady- state mole balance on reactant gives:
𝑓𝑙𝑜𝑤 𝑟𝑎𝑡𝑒 – 𝑓𝑙𝑜𝑤 𝑟𝑎𝑡𝑒 + 𝑟𝑎𝑡𝑒 𝑜𝑓
𝑖𝑛
𝑜𝑢𝑡
𝑔𝑒𝑛𝑒𝑟𝑎𝑡𝑖𝑜𝑛
=
𝑟𝑎𝑡𝑒 𝑜𝑓
𝑎𝑐𝑐𝑢𝑚𝑢𝑙𝑎𝑡𝑖𝑜𝑛
Therefore, the volume element in the mole balance must be differential in length, but
extend across the entire diameter of the reactor. Tubular-flow reactors are normally operated at
steady state so that properties at any position are constant with respect to time. For such steadystate operation applied to the volume element ∆V, becomes
FA│V - FA│V+∆V + rA∆V = 0
FA│V - FA│V+∆V = - rA∆V
Divide by ∆V both the sides
FA│V - FA│V+∆V = - rA
∆V
applying limit ∆V
0
lim
∆V
0
FA│V - FA│V+∆V = - rA
∆V
dFA= - rA
dV
Equation (11)
For a flow system, FA has previously been given in terms of the entering molar
flow rate FA and the conversion X:
Equation (12)
FA = FAO - FAOX
Now for amount of mole convert in differential form ,
FAO dX= - rA
dV
Equation (13)
Integration with the limit V=0 when X=0 gives:
FAO
𝑥
0
𝑑𝑋 = -rA
FAO
𝑥
0
𝑑𝑋 = -rAV
𝑉
𝑑𝑉
0
But, the rate of reaction for first order is:
- rA = k1 CA
So,
V = FA
o
𝒙
𝒅𝑿
𝟎
- k1 CA
Now the the concentration conversion is,
Equation (14)
CA(1+εX) T = CAo (1 - X) To
ε = Voidage is the proportion of unoccupied volume (that is, gaps or empty spaces) in a
volume of some material. The term voidage is normally used to refer to the tiny spaces between
particles in a powder or granulated material like sand
CA = CAo
(1 − X) To
Equation (15)
(1+εX) T
put CA into equation (14) from equation (15)
V = FAo 𝒙 (1+εX) T 𝑑𝑋
𝟎 (1 − X)To
- k1CAo
V = FAo
[
k1CAo
V = FAo [
k1CAo
V = FAo
𝑥 1
0 1−𝑥
𝑑𝑥 +ε
𝑥 1
𝑑𝑥
0 1−𝑥
+ε
𝑥 X
0 1−𝑥
𝑇
𝑑𝑥]𝑇𝑜
𝑥 x−1+1
𝑇
𝑑𝑥]𝑇𝑜
0 1−𝑥
𝑥 1
0 1−𝑥
𝑑𝑥 +ε
𝑥 x−1
𝑑𝑥
0 1−𝑥
V = FAo [
k1CAo
𝑥 1
0 1−𝑥
𝑑𝑥 - ε
𝑥 1−x
0 1−𝑥
V = FAo [
k1CAo
𝑥 1
0 1−𝑥
𝑑𝑥 - ε
𝑥
0
[
k1CAo
+ε
𝑥 1
0 1−𝑥
𝑑𝑥 ]𝑇𝑜
𝑑𝑥 + ε
𝑥 1
0 1−𝑥
𝑑𝑥 ]𝑇𝑜
𝑥 1
0 1−𝑥
1 𝑑𝑥 + ε
𝑇
1
1
𝑇
1
𝑇
V = FAo [ ln (1−𝑥) (1 + ε) − ε x]𝑇𝑜
k1CAo
𝑇
𝑇
𝑑𝑥 ]𝑇𝑜
V = FAo [-ln(1 − 𝑥) – ε x − ε ln(1 − x)]𝑇𝑜
k1CAo
V = FAo [ ln (1−𝑥) +ε ln (1−x) – ε x]𝑇𝑜
k1CAo
𝑇
K1 = FAo [ ln 1 (1 + ε) − ε x] 𝑇
(1−𝑥)
𝑇𝑜
V
Equation (16)
CAo
From equation (16), the values of k1 are calculated for any component.
From Arrhenius equation plot Ln k1 vs 1/T, the slope represents –E/RT to
calculate the activity energy (E) and the intercept represents Ln k◦.
The relationship between Lnk1 vs 1/T using Arrhenius equation.
Lnk1 = Lnko -
𝐸
𝑅𝑇
Substitute values of k1 in equation (10) to calculate values of k2
Ciso = CAO [1- exp(- k1t) -
[exp (- k1t) - exp (- k2t) ]
Figure (5.36) Arrhenius plot
♦WHSV=1.5hr-1, ■WHSV=3hr-1, ▲WHSV=4.5hr-1
Figure (5.37) Arrhenius plot
♦WHSV=1.5hr-1, ■WHSV=3hr-1, ▲WHSV=4.5hr-1.
Figure (5.38) Arrhenius plot
♦WHSV=1.5hr-1, ■WHSV=3hr-1, ▲WHSV=4.5hr-1
Figure (5.39) Arrhenius plot
♦WHSV=1.5hr-1, ■WHSV=3hr-1, ▲WHSV=4.5hr-1
Figure (5.40) Arrhenius plot
♦WHSV=1.5hr-1, ■WHSV=3hr-1, ▲WHSV=4.5hr-1
Figure (5.41) Arrhenius plot
♦WHSV=1.5hr-1, ■WHSV=3hr-1, ▲WHSV=4.5hr-1.
Figure (5.42) Arrhenius plot
♦WHSV=1.5hr-1, ■WHSV=3hr-1, ▲WHSV=4.5hr-1.
Figure (5.43) Arrhenius plot
♦WHSV=1.5hr-1, ■WHSV=3hr-1, ▲WHSV=4.5hr-1.
Figure (5.44) Arrhenius plot
♦WHSV=1.5hr-1, ■WHSV=3hr-1, ▲WHSV=4.5hr-1.