8/18/2019 PENEX Unit Data1-Libre http://slidepdf.com/reader/full/penex-unit-data1-libre 1/49 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 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 r 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).
In worldwide production of automotive gasoline permanent tendency to the tougheningof not only its operating but also its ecological characteristics is observed. So, international anddomestic regulations to automotive gasoline considerably limit the content of benzene, aromatic
hydrocarbons, olefin hydrocarbons and sulfur. In 1970s the variants of hydrogenation of thebenzene, contained in the reformate, proceeding without the decrease of product octane numberhave been offered. However for decrease of the total aromatics content the dilution of reformatewith high-octane nonaromatic components is required. This situation is complicated by refusalfrom tetraethyl lead (TEL) and deficit of butane-butylene fraction (because of the lack of FCCduty), 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 forsolution of this problem. It allows the producing of commercial gasoline which corresponds tothe current and perspective requirements to the fuels and provides necessary flexibility ofprocessing.
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 definesby performances of usable catalysts due to their type. Main parameter which is the octanenumber of produced isomerizate depends on process temperature. That‘s why we willthe issue of thermodynamic of isomerization reaction. First of all hydrocarbons isomerizationreaction is balanced reaction, and equilibrium yield of isoparaffins increases with temperaturereducing, but it can be reached only after an ―infinite residence time‖ of the feed in ror an equivalent very small value for LHSV. On the other hand an increase in temperaturealways corresponds to an increase in reaction velocity. So that at low temperature the actual yieldwill be far below the equilibrium yield, because of low reaction velocity. On the contrary, athigher 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 differenttypes of isomerization catalysts is represented below (Figure 3.2).
The conversion level of n-paraffins on zeolite catalysts is low, as it is limited by thermodynamicequilibrium. In the case of chlorinated-alumina catalysts and sulfated metal oxides conversion ofn-paraffins is higher because of high equilibrium content of isocomponents in product.
Figure 3.2. Comparative estimation of isomerization catalysts
Figure 3.1. Dependence of n-paraffins conversion on reaction temperature
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. Howeverthey possess high resistance to impurities in the feed and capability for total regeneration in thereactor of the unit. The technological scheme of this process is provided with fire-heaters forheating 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 compsupplying of recycle hydrogen-rich gas and separator for separation of hydrogen-rich gas arenecessary .
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 resistanceto impurities in the feed (available sulfur content is 100 ppm permanently and 200 ppm duringshort periods of time)
CHLORINATED-ALUMINA
Chlorinated-alumina based catalysts are the most active and supply the highestisomerizate yield and isomerizate octane. It should be noted that during isomerization catalystsloose chlorine, con sequently the activity is reduced. That‘s why chlorine compound injecthe feed (usually ССl4) is provided for keeping of high activity. As a result, caustic soda washingfrom 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, tonitrogen) and requires pretreatment and drying of the feed. In addition the problems occur atregeneration
The first generation catalyst of UOP is I-8, which was improved later in more active I-80type catalyst. The latest developments of UOP Company are high-performance I-8 Plus, I-82, I-84 catalysts for Penex process and I-122, I-124 catalysts, which are used in Butamer process (n-butane isomerization process with purpose to produce isobutane, which is the feed for alkylationunit). In development of new catalysts UOP has the target to decrease its platinum contentwithout losing the activity, thereby to reduce significantly its operating costs. It is not of smallimportance for present-day refinery.
SULFATED METAL OXIDES BASED CATALYSTS
Sulfated metal oxides based catalysts get heightened interest last years as they combinemain advantages of medium-temperature and low-temperature catalysts. They are active,resistant to poisonous impurities and able for regeneration. The only drawback, as for zeolitecatalysts, 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] andcharacterized with unique sulfur resistance. If necessary, the process can be carried out withoutpretreatment of the feed. In this case the octane number of isomerizate is reduced by 2 points, buttotal lifetime (8- 10 years) doesn‘t changes and service cycle is no less than 12 months. The feedmay contain considerable quantity of benzene which is hydrogenated efficiently on the catalyst.
The Pt/WO 3-ZrO 2 catalyst shows higher activity and selectivity in isomerization reactionof n-alkanes compared to sulfated-zirconia catalysts. The advantage of this type of catalyst isexplained by rapid surface diffusion of hydrogen atoms, which are converted into protons andhydrides on the Lewes acid sites, thereby increasing catalyst activity and selectivity.
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 widerange of operating configurations for optimum design flexibility and feedstock processingcapabilities. The Penex process is a fixed-bed procedure that uses high activity chloride-promoted catalysts to isomerize C5/C6 paraffins to higher octane branched components. Thereaction is conducted in the presence of a minor amount of hydrogen. Even though the chlorideis converted to hydrogen chloride, carbon steel construction is used successfully because of thedry 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 octanecomponents 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 incapital 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. Thecatalyst can be replaced in one reactor while operation continues in the other. One characteristicof the process is that catalyst deactivation begins at the inlet of the first reactor and proceedsslowly as a rather sharp front downward through the bed. The adverse effect that suchdeactivation 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 topermit isolation of the reactor containing the spent catalyst while the second reactor remains inoperation. After the spent catalyst has been replaced, the relative processing positions of the tworeactors 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 mostimportant aspects is desired product octane.
The hydrocarbon once-through flow scheme is the most widely used isomerizationprocess for producing moderate octane upgrades of light naphtha. Economically efficien-through‖ scheme withoutany recycle can be used with minimum investment in realization ofisomerization process Figure (3.3).
Separation and recycle of unconverted normal C 5 and C 6 paraffins and low octane C 6
isoparaffins back to the reactor, produce a higher octane product. The most common flowscheme uses a deisohexanizer (DIH) column to recycle methylpentanes, n-hexane, and some C 6
cyclics. It is the lowest capital cost option of the recycle flow schemes and provides a higheroctane isomerate product, especially on C 6 rich feeds. In the Penex/DIH process the stabilizedisomerate is charged to a DIH column producing an overhead product containing all the C 5 anddimethylbutanes. Normal hexane and some of the ethylpentanes are taken as a side-cut andrecycled back to the reactors. The small amount of bottoms (C 7
+ and some C 6 cyclics) can besent to gasoline blending or to a reformer.
The addition of a deisopentanizer (DIP) or a super DIH will achieve the highest octanefrom a fractionation hydrocarbon recycle flow scheme. In this scheme, both low octane C 5 andnormal and isoparaffin C 6 are recycled to the Penex reactors .
The scheme with deisopentanizer (DIP) before the reactor section allows the producing ofisomerizate with high octane number, increasing of conversion level of n-pentanes and reducingthe reactor duty simultaneously. The technology is reasonable in the case of isopentanes contentin the feed more than 13-15 % Figure (3.4)
The scheme with deisohexanizer (DIH) after the isomerization reactor is the simplest wayto produce the isomerizate with higher octane number. In this case non-converted low-octanecomponents (methylcyclopentane and n-hexane) are recycled into reactor. However the givenscheme allows only increasing of hexanes conversion, but doesn‘t raise the content ofisopentanes in the product (Figure 3.5). The scheme of the process may include bothdeisopentanizer and deisohexanizer (with DIP and DIH)
TABLE 3.3 Typical Estimated Yields for Deisohexanizer Processing
Scheme with recycle of n-pentane (with DIP and DP) requires providing withdepentanizer 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 setof 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 ofpores with definite size to adsorb selectively the molecules of n-paraffins. The next stage isdesorption 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 n-paraffin from the reactor effluent. The Molex process is an adsorptive separation method thatutilizes molecular sieves for the separation of n-paraffins from branched and cyclichydrocarbons. The separation is effected in the liquid phase under isothermal conditionsaccording to the principles of the UOP Sorbex separations technology. Because the separationtakes 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 ona solid adsorbent. In broad outline, Sorbex is a simulated moving bed adsorption processoperating with all process streams in the liquid phase and at constant temperature within theadsorbent bed. Feed is introduced and components are adsorbed and separated from each otherwithin the bed. A separate liquid of different boiling point referred to as ‗desorbent‘ isdisplace the feed components from the pores of the adsorbent. Twoliquid streams emerge fromthe bed – an extract and a raffinate stream, both diluted with desorbent. The desorbent isremoved 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 . Theadsorbent is fixed while the liquid streams flow down through the bed. A shift in the positions ofliquid feed and withdrawal, in the direction of fluid flow through the bed, simulates themovement of solid in the opposite direction. It is, of course, impossible to move the liquid feedand withdrawal points continuously. However, approximately the same effect can be producedby providing multiple liquid access lines to the bed, and periodically switching each net streamto the next adjacent line.
A liquid circulating pump is provided to pump liquid from the bottom outlet to the topinlet of the adsorbent chamber. A fluid- directing device, known as a ‗rotary valve‘, is alsoprovided. The rotary valve functions on the same principle as a multiport stopcock.
TABLE 3.4 Typical Estimated Yields for Molex Processing
UOP offers the processes with adsorption systems on the molecular sieves in vapourphase (Penex/Iso Siv) and liquid phase (Penex/Molex (Figure 3.9)), and process, whichcombines adsorptive separation of unconverted n-paraffins from isomers and deisohexanizingPenex/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 reformateblend), includes feed treatment section which is hydrogenation of benzene.
To maintain the high activity of the Penex catalyst, the feedstock must be hydrotreated.
However, costly pre-fractionation to sharply limit the levels of C 6 cyclic and C 7 compounds is notrequired. In fact, the Penex process affords the refiner with remarkably good flexibility in thechoice of feedstocks, both at the time of design and even after the unit has been constructed. Thelatter is important because changes in the overall refinery processing scheme may occur inresponse to changing market situations. These changes could require that the composition of theisomerization 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 percentC7 with minimal or no effect on design requirements or operating performance. Generally, thebest choice is to operate with lower levels of C 7
+ material because these compounds are bettersuited 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 producecarbon on the catalyst. When the feed has extremely high levels of benzene, a Penex-Plus unit isrecommended. (The ―Plus‖ section can be retrofitted to an existing Penex unit shoulwant to process high-benzene feedstock in an existing Penex unit.) The low-octane C 6 cutrecovered from raffinate derived from aromatic extraction operations typically contains a fewpercent 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 byconventional hydrotreating. Sulfur reduces the rate of isomerization and, therefore, the productoctane number. Its effect is only temporary, however, and once it has been removed from theplant, the catalyst regains its normal activity.
Water, other oxygen-containing compounds, and nitrogen compounds are the onlyimpurities normally found in the feedstock that will irreversibly poison the Penex catalyst andshorten its life. Fresh feed and makeup hydrogen are dried by a simple, commercially provendesiccant system.
PROCESS FLOW DIAGRAM
The UOP Penex Unit can be divided into ten sections.
A. Sulfur Guard BedB. Liquid Feed DriersC. Makeup Hydrogen DriersD. Feed Surge DrumE. Exchanger CircuitF. lsomerization ReactorsG. StabilizerH. Stabilizer Gas ScrubberI. Separator and Compressor Section (Recycle Gas Units Only)
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 bedreduces the sulfur to a safe level for H.O.T. Penex operation and serves as insurance againstupsets in the NHT which could result in higher than normal levels of sulfur in the feed. Theguard bed is loaded with UCP ADS-11 adsorbent, a nickel- containing extrudate designed tochemisorb sulfur from the liquid feed. The feedstock is heated to the required temperature forsulfur removal, usually 250°- (120°C) and passed down flow over the adsorbent. Once sulfurbreakthrough occurs, normally after one year or so of operation, the guard bed is taken off lineand reloaded with fresh adsorbent. The Penex Unit need not be shut down during the short periodof 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 H 20. Thepiping is designed so that either drier can be in the lead or the lag position in series flowoperation. Either drier can be operated individually while the other is being regenerated. Thedriers are designed for a 48 hour cycle which includes 24 hours in the lead position, 7 hoursregenerating, and 3 hours cooling and 14 hours in the lag position. Proper drier operations areessential in the Penex process since the catalyst is water intolerant. Typically, type 4A molecularsieves are employed within the driers.
Charge to the liquid feed driers is hydrotreated LSR naphtha from the Naphthahydrotreater stripper bottoms. Additional recycle feed from a Molex unit or deisohexanizer maybe added in the future.
Liquid feed is pumped to the driers on flow control at 100°F (38°C). The low temperaturemust 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 coldcombined 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 aguard bed. It reduces the water content even further but is essentially in standby until 1.0 wppmH20 breaks through the lead drier (or until the scheduled regeneration period arrives). The watercontent is continuously monitored with a Parametric moisture analyzer. This analyzer willalways be used to monitor the lead drier effluent. If the lead drier effluent reaches 1.0 wppm H 20content, 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 theregenerated drier now in the lag position. See the liquid feed drier regeneration procedure fordetails of the regeneration.
C. MAKEUP HYDROGEN DRIERS
The makeup hydrogen driers are used to dry the Penex unit makeup hydrogen to less than0.1 ppm H 20. 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 seriesoperation either drier can be in the lead or the lag position and so that either drier can be operatedindividually. while the other is being regenerated. The driers are designed for a 48 hour cyclewith includes 24 hours in the lead position, 6 hours regenerating, 4 hours cooling and 14 hours inthe lag position. These driers, too, are essential for good Penex operations since the I-8 catalyst iswater 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). Theactual makeup rate will vary with chemical consumption due to benzene saturation, ring openingand hydrocracking. The makeup gas flow rate is controlled by the H 2O/HCBN ratio controller onthe fresh feedstock. The ratio is adjusted to maintain excess hydrogen at the reactor outlet. Theincoming gas passes up flow through the lead driers, crosses over to the bottom of the lag drierand passes up flow through it to the cooler in the separator and compressor section. Like theliquid feed driers, the lead makeup gas drier removes water from the makeup hydrogen to lessthan 0.1 wppm The lag drier acts as a guard bed until 1.0 ppm H 20 breaks through the lead drier.Water content is continuously monitored with a moisture analyzer. This analyzer will always beused to monitor the lead drier effluent. If the lead drier effluent reaches 1.0 ppm H 20 contend itmust be taken off line and regenerated. The previous lag drier becomes the lead drier. Oncompletion of the regeneration, series flow is re-established with the regenerated drier now in thelag position. See the makeup hydrogen direr regeneration procedure. for the details of theregeneration.
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 blanketedwith 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 reactorcharge pumps through the reactor exchanger circuit on flow control. The reactor exchangercircuit consists of the cold combined feed exchanger, the hot combined feed exchanger, and thereactor charge heater. Prior to the entry of the liquid hydrocarbon into the cold combined feedexchanger, it combines with the makeup hydrogen stream. After combining, the mixedhydrocarbon-hydrogen stream passes through the exchanger circuit in the order previouslymentioned.. After the makeup gas combines with the feed a small quantity of catalyst promoter
(CCI 4) 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 feedexchanger is equipped with a bypass which can be used to regulate the amount of combined feedpreheat. The bypass is regulated with a board mounted control valve. The combined feed isfinally brought up to the desired temperature in the reactor charge heater by a temperaturecontroller which resets the exchangers heating medium flow. The charge heater is equipped withan automatic shutdown which is activated by low feed or low makeup gas flow. After exiting thereactor 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 willbe placed in series with the other reactor. At various times throughout the unit‘s history it will bepossible to have either reactor in the lead or tail position. Thermocouples are inserted into thecatalyst bed of each reactor to monitor the activity of the catalyst. After exiting the reactorcharge heater, the heated combined stream then flows to the first reactor. Upon exiting the firstreactor, the stream then passes to the hot combined feed exchanger where the first reaof reaction is partially removed. The degree of temperature removal can be achieved by adjustingthe amount of exchanger bypassing with a temperature controller.The partially cooled stream isthen routed to the second reactor where the final process reactions are completed. The reactorsare equipped with hydrogen purge lines which are located at the inlet of each reactor. Thehydrogen purge is used to remove hydrocarbon from a reactor which is to be unloaded or to coola 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 withdepressuring lines to the flare system. The reactors are depressured from the outlet of the lagreactor. The depressuring line is equipped with two motorized valves which can be operatedfrom the control room. After exiting the second reactor, the stream is then routed to the tube sideof the cold combined feed exchanger. The cold combined feed exchanger tube side effluent isthen routed to the stabilizer on pressure control.
The purpose of this column is to separate any dissolved hydrogen, HCl and cracked gases(C1, C2, and C 3‘s) from the isomerate.The feed to this column is routed hot directly frombefore entering the stabilizer.The column is reboiled by either steam or hot oil. The reboiler heatinput 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 acooled 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 thereceiver on level control with the stabilizer reflux pump. All liquid from the stabilizer overheadreceiver is refluxed to the column on tray No. 1. Bottoms product is routed to storage on levelcontrol after first being cooled in the stabilizer bottoms cooler. If the stabilizer bottoms is sent toa Deisohexanizer it is not cooled, but is charged hot to the column. Part of the stabilizer bottomsis used for regenerating the driers.
H. STABILIZER GAS SCRILIBBERThe 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 onbackpressure control. The hydrogen purity is monitored on the scrubbed off gas to determine themoles of H 2 leaving the system for the H 2 /C:H determination. Make-up caustic is pumped fromthe refinery to the reservoir section of the gas scrubber when caustic addition is required. Thecaustic in the reservoir section is pumped by the caustic recirculating pumps to the top of thescrubbing section of the scrubber where a counter current contact with the rising acidic gas ismade. Caustic is also continuously circulated to the distributor under the packed section. Theflow rate of the circulating caustic can be monitored by a local flow indicator. Periodically aportion of the caustic is withdrawn to the refinery spent caustic facilities as spent caustic. Thecaustic level in the scrubber is maintained about 1-2 feet below the distributor under the packedsection.
I. SEPARATOR AND COMPRESSOR SECTION (Recycle Gds Units Only)
The separator and compressor section separates the reactor effluent into unstabilizedliquid product and recycle gas. The separator pressure is controlled by regulating the makeuphydrogen flow rate. The equipment in this section is: the reactor product condenser, the productseparator, the recycle gas compressor, and if required, the make-up gas compressor suctiondrum, and the make-up gas compressor.
In a Hydrogen Once Through Unit the product condenser, product separator and recyclecompressor are not used. In this unit the pressure in the reactor circuit is controlled using a backpressure value on the stabilizer feed line.
Reactor effluent exits the reactor section and is partially condensed in the reactor productcondenser. It cools the effluent to about 100°F. The cooled liquid and gas then separate in theproduct 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 therecycle gas compressor to the cold combined feed exchanger in the reactor section. Recycle gasflow 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 gasflow 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 themake-up gas compressor. These gases pass through the make-up gas cooler and into the makeupgas compressor suction drum. Any entrained hydrocarbons are knocked out and are manuallydrained to an appropriate location. The make-up gas is compressed and combined with therecycle gas to the reactor section. Make-up gas flow is controlled by the product separatorpressure recorder controller.
Paraffin isomerization catalysts fall mainly into either of two principal categories: (1)those based on Friedel-Crafts catalysts as classically typified by aluminum chloride/hydrogenchloride, or (2) dual-function hydro-isomerization catalysts. No attempt is made to present adiscussion of mechanisms of a degree of sophistication acceptable to a chemist specializing inthe area. The intention is simply to provide those practicing engineers who have not previouslyhad reason to consider isomerization with a basic introduction to the subject. Isomerization byeither Friedel-Crafts or duals function catalysts is generally thought to entail intramolecularrearrangements of carboniumions as illustrated - for pentane:
(1) CH 3-CH-CH 2-CH 2-CH 3 CH 3-C2H3-CH 2-CH 3
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 relativeimportance under specific circumstances.
Friedel-Crafts isomerization is believed by some to require the presence of traces of olefinsor alkyl halides as carbonium ion initiators, with the reaction thereafter proceeding throughchain propagation. The initiator ion, which needs to be present in small amounts only, may beformed by the addition of HCl or HAlCI 4 to an olefin, which is present in the paraffin as animpurity or which is formed by cracking of the paraffin:
(2) RCH=CH 2 + HAlCl 4 RCHCH 3 + AlCl 4
The initiator then forms a carbonium ion from the paraffin to be isomerized:
Naturally, the same sequence could have been illustrated starting with Isopentane andending 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 finalmixture is, of course, that set by thermo-dynamic equilibrium, assuming that sufficient reactiontime has been provided. Another Friedel-Crafts route which has been suggested is direct hydrideion abstraction:
Abstraction of the hydride ion is energetically favored by the fact that the aluminum atomcan 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:
Some chemists feel uncomfortable with the above because of the required postulation ofhydrogen formation. The dual-function hydro-isomerization catalysts are thought by some tooperate through an olefin intermediate whose formation is catalyzed by the metallic component,assumed for illustration purposes to be platinum: -
This reaction is, of course, reversible and, since these catalysts are employed undersubstantial hydrogen pressure, the equilibrium is far to the left. However, the acid function of thecatalyst consumes the olefin by formation of a carbonium ion and thus permits more olefin toform despite the unfavorable equilibrium. This step is entirely analogous to Reaction (2) shownfor Friedel- Crafts, except that it is better to denote the acid function by a more general.
(11) CH 3-CH 2-CH 2 — CH=CH 2 + H + A - CH 3-CH 2-CH 2-CH-CH 3 + A -
The usual rearrangement ensues:
(12) CH 3-CH 2-CH 2-CH-CH 2 CH 3-CH 2CCH 3-CH 3
The is olefin is then formed by the reverse analogue of (11):
(13) CH 3-CH 2C CH 3-CH 3 + A - CH 3-CH 2-C=CH 2 + H + A -
The iso-paraffin is finally created by hydrogenation:
Those dual — functi0nal hydro-isomerization catalysts which operate at very lowtemperatures have stronger acid sites than those which require higher temperatures. In this case it
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 readilyliberated for further reaction. This may account, at least in part, for the higher activity of suchdual functional catalysts.
Equilibrium limits the maximum conversion possible at any given set of conditions. Thismaximum 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 thepentane 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 tooccur 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 thatshown in Fig. 5.1. The methylpentanes have been combined because they have nearly the sameoctane rating. The methylpentane content in the C 6-paraffin fraction remains nearly constant overthe entire temperature range. Similarly, the fraction of 2,3-dimethylbutane is almost constant atabout 9 mol % of the C 6 paraffins. Theoretically, as the temperature is reduced, 2,2-dimethylbutane can be formed at the expense of normal hexane. This reaction is highly desirablebecause nC
6 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 plottedagainst the temperature characteristic of that equilibrium for a typical Chargestock. Both the C 5 and the C 6 paraffins show an increase in octane ratings as the temperature isreduced.
Because equilibrium imposes a definite upper limit on the amount of desirable branchedisomers that can exist in the reactor product, operating temperatures are thought to provide asimple basis for catalyst comparison or classification. However, temperature is only anapproximate comparison that at best can discard a catalyst whose activity is so low that it mightbe operated at an unfavorably high temperature. Further, two catalysts that operate in the samegeneral low-temperature range may differ in the closeness with which they can approachequilibrium in the presence of reasonable amounts of catalyst.
B. ALUMINIUM CHLORIDE
The isomerization catalysts employed during World War Il were all of the Friedel-Craftstype. Those which contained aluminum chloride only were either a hydrocarbon/aluminumchloride complex (the so-called sludge process) or they were manufactured in deposition onto asupport such as alumina or bauxite. They were intended to operate at very low temperatures(120-265 0F) and to approach the very favorable equilibrium composition characteristic of thesetemperatures.
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 speedthe 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 ofaluminum chloride in hydrocarbon, to its relatively high volatility and to the difficulty ofremoving it from the product by caustic washing. Aluminum chloride deposition in and pluggingof reboiler tubes was not uncommon. The process faced problems in sludge disposal which were
FIGURE 5.3 Unleaded RON Ratings Of Equilibrium Fractions.
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 typeisomerization plants, the advent of catalytic reforming which not only made hydrogen generallyavailable in refineries but also demonstrated the practicality of using noble metal containingcatalysts on a large scale, and the octane number race which postwar high compression enginesinitiated all combined in the 1950‘s to spawn a spate of hydro-isomerization processes. Thesecatalysts generally contained a noble metal and some halide, operated at temperatures betweenabout 560°F and temperatures approaching those characteristic of catalytic reforming, employedrecycle hydrogen to prevent catalyst carbonization and utilized either no promoter or traces atmost. In general, they did not require an especially dry feedstock but did benefit from a lowsulfur content feedstock. Most achieved a close approach to the equilibrium characteristic oftheir particular operating temperature. Because of their high operating temperatures and their
necessarily low conversions to iso-paraffins, these high temperature catalysts were quicklyreplaced 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 puanything below 390°F operating temperature. Typically these are fixed bed catalysts containing asupported noble metal and a component to provide acidity in the catalytic sense. They operate ina hydrogen atmosphere and may employ a catalyst promoter whose concentration in the reactormay 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 andmolecular weights of hydrocarbons. Hydrocracking to light gases is generally slight, so liquidproduct yields are high. The type of catalyst used in the Penex unit is of this type. Apart from theparaffin isomerization reactions which were discussed in detail in the proceeding pages, there areseveral other important reactions including:
1. Naphthenes Ring Opening:
Penex feeds can contain up to 30% naphthene rings. The three Naphthenes which aretypically present in Penex feed are cyclopentane (CP), methyl cyclopentane (MCP) andcyclohexane (CH). The naphthene rings will hydrogenate to form paraffins. This ring openingreaction increases with increasing reactor temperature. At typical Penex reactor conditions, theconversion 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 shifttowards MCP production as temperature is increased.
Penex feeds can contain up to 4% benzene. The catalyst will saturate benzene tocyclohexane. The catalyst-will saturate benzene to cyclohexane. This reaction proceeds veryquickly 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 producesheat. 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 thefeed quality and severity of operation. Large molecules such as C 7‘s tend to hydrocrack moreeasily than smaller molecules. C 5 and C 5 paraffins will also hydrocrack to a certain extent. AsC5 /C 7 paraffin isomerization approaches equilibrium, the extent of hydrocracking increases. Ifisomerization is pushed too hard, hydrocracking will reduce the liquid yield and increase heatproduction. 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 2 METHYL PENTANECH 3
CH 3-CH 2-CH 2-CH 2-CH 2-CH 3 CH 3-CH-CH 2-CH 2-CH 3 24.8 ron 0 73.4 ron
3 METHYL PENTANECH 3
CH 3-CH 2-CH 2-CH 2-CH 2-CH 3 CH 3-CH 2-CH-CH 2-CH 3 24.8 ron 74.5 ron
In the normal operation of a Penex unit having once set the operating pressure, fresh feedrate and the recycle hydrogen and makeup hydrogen flows, it is usually only necessary to adjustthe reactor inlet temperatures. Nevertheless, it is to the operator‘s advantage that he hathorough understanding of the influence process variables will have on performance of the unitand the life of the catalyst. Once the catalyst has been loaded into the unit, the manner in whichthe catalyst is placed in service and the treatment it receives when in service will to a large extentinfluence 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 primeconsideration for it can be regarded as the heart of the operation on which the quality of theresults 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 naphthawhich is usually processed in a catalytic reformer. The prime consideration in splitting thefeedstock is to maximize the C 5 /C 6 paraffins to the Penex unit keeping the C 7+ material in thenaphtha 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 alsopresent in the splitter feed and adecision on which unit to send them to must be made. Ideally,the C 6 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 componentboiling points and volatilities. It is observed that the benzene and normal hexane have essentiallythe same volatility and, hence, cannot be separated by fractionation. Therefore, the decision mustbe 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 ofnormal hexane in both processes, it is seen that a catalytic reformer leaves a significant level ofnormal hexane in the C 6 product. The question then becomes which is better to leave, low octanen-C 6 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-C 6 and benzene in the Penexfeedstock., Most of the cyclohexane should be fractionated out of the Penex feed and sent to thereformer. Hence, the Penex feed should include all of the C 6 paraffins plus benzene and someMCP and the reformer feed should contain most of the cyclohexane and C 7+. This feedpreparation philosophy may shift as the percentage levels of MCP plus benzene increase in theC6 fraction and as the reforming selectivity in converting n-C 6 to benzene increases.For those operations that process reformates in an aromatics recovery unit such as Sulfolane or
Udex, the C 6 cyclics and n-C 6 should be included in the reformer feedstock. The unconverted C 6 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 percen
of isopentane to total C 5 aliphatic paraffins, the percentage ratio of 2,3 dimethylbutane in thetotal C6 aliphatic paraffins, or the percentage .ratio of normal hexane in the C 6 aliphatic paraffinsin the stabilizer bottoms stream, The terms ―liquid feed," ―reactor charge" and "comrefer to the C 5 /C 6 charge to the liquid feed dryers of the unit, the effluent from the liquid feeddryers and the reactor charge plus makeup hydrogen and recycle hydrogen gas, respectively.
In general, reactor temperature is the main process control. A definite upper limit existsfor 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 afteran infinite length of time, i.e. with an infinitely large reactor. in practice, therefore, the productwill contain slightly less iso-paraffins than this equilibrium concentration. As the reactortemperature is raised to increase the rate of isomerization, the equilibrium composition will beapproached more closely. At excessively high temperatures, the concentration of iso-paraffins inthe product will actually decrease because of the downward shift in the equilibrium curve, eventhough the high temperature gives a higher reaction rate.
In the case of the Hydrogen Once Through Penex designs, where the hydrogen contentof the reactor charge is so much lower than for recycle gas Penex, the reactor feed will contain aconsiderable 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 lowerGibbs free energies for the components in the liquid phase. The attached figures illustrate theequilibrium for both the vapor and liquid phases for isopentane and 2,2 DMB. The use oftemperatures higher than necessary to achieve a reasonable close approach to equilibriumaccomplishes 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, thecarbon forming propensity of the catalyst is inherently so low that excessive hydrocrackingwould normally be encountered before carbon formation problems would develop. It isrecommended, however, that UOP be consulted before temperatures above about 380°F areemployed.
A typical C5 /C
6 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 ishydrogenated in the first reactor, even when the catalyst therein has lost its activity with respectto paraffin isomerization. Some conversion of cyclohexane and methyl cyclopentane to hexanesalso occurs, as does some hydrocracking of C 7 to C 3 and C 4. These three reactions (benzenehydrogenation, naphthene conversion to hexane, and C 7 hydrocracking) are exothermic and, for atypical feedstock, contribute more to the temperature rise in the first reactor than does paraffinisomerization, which is also exothermic.
Normally, the first reactor will be operated at such a temperature as to maximize theconcentration of isopentane and 2,2 dimethyl butane in its effluent. The concentrations attainableand-the required outlet temperature will be influenced by the amount of active catalyst presentand by the amount of C 6 cyclic and C 7 components present in the feed, higher temperatures beingrequired with high concentrations of these components in the feed. By this procedure, therequired operating temperature for the second reactor is reduced and it is possible to operateunder conditions where the equilibrium is more favorable.
The optimum reactor temperatures are determined in the field by establishing a "base‖ seof 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 eacheffluent. Reactor effluent or product octane is not used for this determination since it is alsodependent on the relative amounts of C 5 and C 6 in the feedstock.
Another interaction may occur with feeds which are rich in C 5 cyclics. Since thesematerials tend to reduce the rate of paraffin isomerization, it may be beneficial with very richfeed stocks to choose the first reactor system temperature to control the amount of` cyclics whichenter the second reactor system. By raising the first reactor system temperature, more of thecyclics be converted to hexanes and the rate of isomerization in the second reactor systemthereby 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. Thedesign LHSV for C 5 /C 5 Penex operation is- normally 1 to 2 and increasing the LHSV beyondthis 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 molshydrogen at the reactor outlet per mol of reactor charge passing over the catalyst and is specifiedat 0.05 mols H 2 /mol H:C. The primary purpose of maintaining the ratio at or above the design isto avoid carbon deposition on the catalyst and maintain enough H 2 for the reactions to proceed. lfnecessary, the reactor charge rate is to be reduced to maintain the design hydrogen tohydrocarbon ratio. The H
2 /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 willgenerally increase the amount of liquid phase material in the reactor and may lead to flowdistribution.
PRESSURE
C5 /C 5 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 ofisomerization reactions. Higher pressure helps to offset this effect of the C 6 cyclic compounds.Lowering the unit pressure or operating at a slightly lower level would not affect the catalyst lifebut the extent of isomerization would be influenced.
To sustain catalyst activity, the addition of chloride is necessary. At no time should theplant be operated for longer than six hours without the injection of chloride. Whenever there is-acatalyst chloride decency, the product isomer ratios will decrease (although not necessarily
instantaneously), other things being equal. Restarting the injection of chloride will tend to returnthe activity of the catalyst to its previous level, but it is possible that full activity will not berestored if a decline in activity, as a result no chloride injection has been observed. Carbontetrachloride and specific grades of perchloroethylene are the only approved sources of chloride.
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 inaddition had corrosive properties. Furthermore, chlorine addition during the reaction is necessaryto guarantee catalyst stability.
In the Hysomer process zeolite based catalysts were used which had the major advantageof 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 arerequired for this technology.
The isomerization of hydrocarbons < C 6 is currently carried out very successfully using
bifunctional supported platinum catalysts. However, difficulties are encountered withhydrocarbons larger than hexane since the cracking reactions become more significant overplatinum catalysts as the chain length increases. Catalysts used in state of the art isomerization-cracking reactors are bifunctional. They have a metal function providing de-hydrogenation andhydrogen 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 theactivity and stability of the catalyst. However they have a low resistance to poisoning by sulfurand 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 themetal 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 contaminantsand that they are regenerable. The chlorinated alumina catalysts are very sensitive tocontaminants such as water, carbon oxides, oxygenate, and sulfur. Thus, feeds and hydrogenmust be hydrotreated and dried to remove water and sulfur. Furthermore, the chlorinated aluminacatalysts 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 removesuch contamination.
The UOP BenSat process uses a commercially proven noble metal catalyst, which hasbeen used for many years for the production of petrochemical-grade cyclohexane. The catalyst isselective and has no measurable side reactions. Because no cracking occurs, no appreciable cokeforms on the catalyst to reduce activity. Sulfur contamination in the feed reduces catalystactivity, but the effect is not permanent. Catalyst activity recovers when the sulfur is removedfrom the system.
Alumina or aluminum oxide (Al 2O3) is a chemical compound with melting point of about2000°C and sp. gr. of about 4.0. It is insoluble in water and organic liquids and very slightlysoluble 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 iscomposed of minute colorless cubic crystals with sp. gr. of about 3.6 that are transformed to thealpha form at high temperatures. Figure 8.1 shows the shape of (Al 2O3).
Identifiers Aluminium oxide
The most common form of crystalline alumina, α-aluminium oxide, is known ascorundum. If a trace of the element is present it appears red, it is known as ruby, but all othercolorations 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 withaluminium 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 theycontain silicon, aluminium and oxygen in their framework and cations, water and/or othermolecules 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 applicationswith a global market of several million tonnes per annum. In the western world, major uses are inpetrochemical cracking, ion-exchange (water softening and purification), and in the separationand removal of gases and solvents. Other applications are in agriculture, animal husbandry andconstruction. They are often also referred to as molecular sieves.
Zeolites have the ability to act as catalysts for chemical reactions which take place withinthe internal cavities. An important class of reactions is that catalysed by hydrogen-exchangedzeolites, whose framework-bound protons give rise to very high acidity. This is exploited inmany organic reactions, including crude oil cracking, isomerisation and fuel synthesis.
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 shape-selective catalysts. Increasingly, attention has focused on fine-tuning the properties of zeolitecatalysts 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 ion-exchangers:
FIGURE 8.2. Structures and dimensions of different types of zeolite
1. Well-defined crystalline structure.2. High internal surface areas (>600 m 2 /g).3. Uniform pores with one or more discrete sizes.4. Good thermal stability.5. Highly acidic sites when ion is exchanged with protons.
6.
Ability to sorb and concentrate hydrocarbons.
The tetrahedral arrangements of [SiO 4]-4 and [AlO 4]-5 coordination polyhedra createnumerous lattices where the oxygen atoms are shared with another unit cell. The net negativecharge is then balanced by cations (e.g. K + or NH 4
+). Small recurring units can be defined forzeolites named, ‗secondary building units.
The primary building blocks of all zeolites are silicon Si +4 and aluminum Al +3 cationsthat are surrounded by four oxygen anions O -2. This occurs in a way that periodic threedimensional framework structures are formed, with net neutral SiO 2 and negatively chargedAlO 2.
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 ofFigure (8.2) the structure of a faujasite type zeolite is shown. The secondary building block ofthis zeolite is a sodalite cage, which consists of 24 tetrahedral in the geometrical form of a cubo-octahedron. The sodalite cages are linked to each other via a hexagonal prism.
The main aim of the present study is to analyze the kinetics of hydroisomerizationprocess by assessing the effect of reaction time and reaction temperature on the performance ofthe 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 bydehydrogenation to n- olefins.
2- Desorption of n- olefin from the dehydrogenation sites and diffusion to a skeletal rearrangedsite, 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 bifunctionalscheme shown below:
The hydroisomerization process of light naphtha is regarded as one of the complexchemical reactions network, where such types of reactions take on a metal and acid sites ofcatalysts.
Therefore, the mathematical modeling of the hydroisomerization process is a veryimportant tool in petroleum refining industries. It translates experimental data intoparameters used as the basis of commercial reactor process optimization.
In the hydroisomerization of alkanes it is supposed that the alkane is dehydrogenated toan 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 iongives the proton back to the acid site, the resulting branched alkene is hydrogenated on themetallic site. The branched alkane is formed, and can be desorbed from the catalyst surface. Thereaction mechanism scheme is shown in Figure (4.1)
The eneral reactions mechanism for isomerization of n-alkane
The catalytic hydroisomerization kinetic the following assumptions are taken intoaccount:
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 withoutpressure 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 thehydroisomerization process. Therefore, the main aim of the present work is to estimate thereaction 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 forlight naphtha (n-paraffin) can be considered by the following scheme depending on the presentmodel assumptions which can be formulated to the following equations:
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:
The Su ested Reactions Of Li ht Na htha Isomerization
To develop a reaction model for an integral reactor, a material balance is made over thecross section of a very short segment of the tubular catalyst bed, as shown in Figure (.2):
The tubular-flow reactor is one in which there is no mixing in the direction of flow andcomplete 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 .Weconclude that the rate of reaction will vary with reaction length.
A steady- state mole balance on reactant gives: – + =
Therefore, the volume element in the mole balance must be differential in length, butextend 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 steady-state operation applied to the volume element ∆V, becomes
FA│V - FA│V+∆V+ rA∆V = 0
FA│V - FA│V+∆V= - r A∆V Divide by ∆V both the sides
FA│V - FA│V+∆V= - r A
∆V applying limit ∆V 0
lim F A│V - F A│V+∆V= - r A∆V 0 ∆V
dF A= - r A
dV
For a flow system, F A has previously been given in terms of the entering molarflow rate F A and the conversion X:
FA = F AO - F A
OX
Now for amount of mole convert in differential form ,
ε =Voidage is the proportion of unoccupied volume (that is, gaps or empty spaces) in avolume of some material. The term voidage is normally used to refer to the tiny spaces betweenparticles in a powder or granulated material like sand
