GB1604900A - Preparation of a pentane/hexane feedstock for isomerisation and the isomerisation process including such preparation - Google Patents

Description

(54) THE PREPARATION OF A PENTANE/HEXANE FEEDSTOCK FOR ISOMERISATION AND THE ISOMERISATION PROCESS INCLUDING SUCH PREPARATION (71) I, ROBERT LEARD IRVINE, a citizen of the United States of America of Rob Nes, Pyle Hill, Woking, Surrey, do hereby declare the invention, for which I pray that a patent may be granted to me, and the method by which it is to be performed, to be particularly described in and by the following statement: This invention relates to an improved process for preparing a pentane/hexane feedstock for isomerisation, together with an improved liquid phase low temperature isomerisation process, which has significant utility and yield advantages over existing isomerisation processes. The low temperature process of the invention converts any C6 naphthenes entering in the feedstock into a 95 percent purity cyclohexane product. This narrow boiling range product is free of aromatics and suitable as a solvent.

The importance of feed preparation for isomerisation is well known. All noble metal hydroisomerisation feeds must have a limited sulphur and water content whereas the lower temperature processes require even more stringent specifications as the catalyst is more sensitive the lower the reactor temperature. Any of the existing isomerisation processes operating below 320"F in the reactor must also preferably have any benzene in the C6 paraffin fraction reduced to low concentration as benzene inhibits isomerisation activity as well as increasing catalyst deactivation of the isomerisation catalyst.

Although hydroisomerisation can tolerate limited C7 components, these components are hydrocracked into propane and butanes and, therefore, decrease the yields. Furthermore, the C7 components which are present in C5/C6 fraction feedstocks are dimethylpentanes which have excellent octane values. Existing hydroisomerisation processes have another yield drawback in that a significant part of the higher octane cyclics, namely, cyclopentane, methylcyclopentane and cyclohexane, are either hydrocracked or converted by decyclization into paraffins. The disappearance is dependent upon space velocity but approximately 28 percent disappearance may be considered typical of present commercial isomerisation operations.

What is generally overlooked is that, becaue of the cyclic and the dimethylpentanes present in the usual dehexanized straight run feedstock, the octane values of such feedstocks when denormalized by molecular sieve separation are generally higher than those resulting from existing commercial isomerisation processes, without any of their yield losses. Furthermore, the dehexanizer separation on the preparation of the C5/C6 fraction for the feed to molecular sieve separation is easier than for preparing a C5/C6 fraction for feed to isomerisation.

The octane values for the two gasoline products will be illustrated by comparing the following typical Middle East hydrotreated light straight run C5/C6 feedstock when denormalized by molecular sieve separation with the reported octane values for the isomerisation product obtained in existing commercial isomerisation processes. The effects of the cyclic components may be considered to be conservative for the octane values of the denormalized product as most straight run feedstocks have considerably higher cyclic and dimethylpentanes content than those of the Middle East crudes. The following table also indicates the properties of the components of commercial interest in typical C5/C6 fractions.

It may be seen that only normal pentane and normal hexane have research or motor clear octane values below 73.4. Molecular sieve separation can routinely and economically reduce the normal paraffins in the denormalized product even without the improvement described in this invention. However reducing the normal paraffin in the gasoline product by distillation to such a low concentration would be economically prohibitive because of the additional utility which would be required apart from investment considerations.

Component properties Composition, weight Specific 1) Middle East fraction Component NBP,"F gravity Octane clear feed Denorma RONcl MONcl lized pro duct nC4 31.10 0.5844 93.8 89.6 .0057 .0006 iC5 82.13 0.6248 92.3 90.3 .1743 .3420 nC5 96.93 0.6312 61.7 61.9 .2812 .0110 CP 120.67 0.7505 101.7 85.0 .0175 .0343 22 DMB 121.53 0.6540 91.8 93.4 .0044 .0086 23 DMB 136.38 0.6664 103.6 94.3 .0197 .0387 2 MP 140.49 0.6579 73.4 73.5 .1243 .2439 3 MP 145.91 0.6690 74.5 74.3 .0972 .1907 nC6 155.73 0.6640 24.8 26.0 .2115 .0041 MCP 161.26 0.7535 91.3 80.0 .0326 0.640 22 DMP 174.55 0.6783 92.8 95.6 .0067 .0132 Benzene 176.18 0.8845 98.0 94.0 .0111 .0138 24 DMP 176.90 0.6772 83.1 83.8 .0050 .0098 CH 177.33 0.7834 83.0 77.2 .0088 .0173 1.0000 1.0000 1) API Project 45 values except for the research of cyclpentane (CP) and the octane values of benzene which were from M.W. Kellogg.

The calculated properties of the denormalized product are 84.2 RONcl, 81.9 MONcl, and 0.662 specific gravity. This compares with 67.5 RONcl for the feed. The foregoing feed, when tested in Shell's Hyisomer pilot plant at 260"C (500"F), produced a research octane clear of 79.2 with the pentane plus yield an 0.965 weight fraction based on the pentane plus feed. The propane and lighter yield correspond to 0.0142 weight fraction of the total feed.

Shell has obtained better octane values for the isomerate product commercially (80.5 to 82.1 RONcl) at higher yields (0.972 to 0.974 weight fraction C5 plus basis). Typical of most commercial C5/C6 feedstocks, dimethylpentanes were not present in the feed so that the higher yields could be expected. The higher octane values could be attributed to the considerably lower cyclic content of the commercial feedstock as too high a cyclic content inhibits isomerization activity. Shell's commercial octane values do not differ significantly from the other commercial isomerisation processes charging a similar C5/C6 feedstock: 81.5 to 83.1 RONcl for the BP isomerisation process; and 83.4 to 83.8 RONcl for the UOP isomerisation process (Penex). It can be seen that the octane values of these commercial isomerate products are lower than the dernormalized product octane values achieved by molecular sieve separation All the aforegoing commercial processes produce isomerate on a once-through basis, as recycle to produce a higher octane value can seldom be justified commercially. The octane value of the denormalized product would be even higher than for the foregoing feedstock from a paraffinic crude with naphthenic crudes (higher cyclopentane and methylcyclopentane contents) as well as from isoparaffinic crudes (higher dimethylpentanes), so that the octane advantage of the denormalized product over the product of existing isomerisation processes should normally be even higher. Even the isoparaffinic crudes have generally over 30 percent of normal paraffins in the C5/C6 fractions, so that the utility expenditure by the molecular sieve separation is justified by the octane improvement of the denormalized product over the feed, as the entire feed must be vaporized and heated to 6000F for separation by the pressure swing molecular sieve spearation process. This means that molecular sieve separation of the isomerisation feed should be considered for maximum yield and octane value, particularly if the molecular sieve separation can be made more thermally efficient and can process an increased amount of throughput for the same molecular sieve inventory in producing a C5/C6 feedstock for isomeristion.

According to the present invention, I provide in a process of isomerisation, the improved preparation of a pentane/hexane feedstock for isomerization, which comprises passing a C5/C6 feed to an splitting stage to obtain a denormalised product stream and a normal paraffin feed stream; and passing the latter stream to a low temperature liquid phase isomerisation stage using a Friedel-Crafts halide catalyst suitable for liquid hydrocarbon phase contacting and promoted with hydrogen halide recirculating stream to the catalytic reactor, the isomerate splitter bottom liquid containing dissolved Friedel-Crafts catalyst in the liquid phase from the isomerization reactor being returned as recycle with the C6 paraffins not converted into desired C6 isomeric product, together with sufficient naphthene to suppress, in conjunction with hydrogen any disproportionation and hydrocracking reactions.

Although the low temperature liquid phase isomerisation process described herein can produce much higher once-through octane values than the existing commercial processes, isomerisation from an investment and utility viewpoint is best performed on the normal paraffin components (nC5, nC), gas these are the components that detract from the gasoline pool octane. Furthermore, the investment per unit of fresh feed together with the utilities and the product are practically the same whether a mixture of paraffins or normal paraffins alone constitutes the feedstock to isomersiation.

The high octane values produced by the isomerisation process described in the invention are due to its considerably lower reactor temperatures (approximately 100"F for the terminal reactor), which increase the equilibrium value of the paraffins having the highest octane number, e.g. isopentane in the C5 paraffin fraction and that of the 22dimethylbutane in the C6 paraffin fraction. Because of the 22-dimethylbutane concentration of the C6 fraction in the process, there is because of lower reactor temperatures an approximately 0.50 fraction of the C6 paraffins in the reactor product as compared with an approximately 0.32 to 0.34 fraction in the BP of Penex isomerisation products, which give the best values of the commercial processes. All the C6 paraffins are normally converted to over 95 percent dimethyl-butanes in the isomeric product because of the ease of separation (being considerably easier than deisopentanization). The high octane of the resulting isomerate product (typically 89.4 RONcl, 88.8 MONcl) of the invention described herein further removes the necessity for any more expensive octane upgrading process than the molecular sieve separation unit as the feed preparation unit for the isomerisation process.

The isomerisation process would normally suffer from an octane viewpoint as regards methylcyclopentane in the C6 naphthene fraction because the methylcyclopentane content in the C6 naphthene fraction is lowered, the methylcyclopentane being converted to cyclohexane. Thus, the methylcyclopentane fraction in the C6 naphthenes leaving the reactor is only approximately 0.14. This means that the C6 naphthenes from a typical straight run on a catalytic reformer C > /C6 fraction would undergo an octane loss due to isomerisation as the octane number of the cyclohexane product resulting from the conversion of the methylcyclopentane entering would correspond to an octane decrease of 18.7 RONcl. This effect is avoided (by molecular sieve) as only a low portion of the methylcyclopentane in the original feed enters. Furthermore, C6 naphthenes are not part of the isomerate product as all C6 naphthenes entering as feed are ultimately removed as a more valuable cyclohexane solvent product. The isomerisation process of this invention limits the benzene concentration of the feed preferably below 50 parts per million and the C7 components entering in the feed preferably below 500 parts per million. Although higher levels of these contaminants can be tolerated with increased deactivation rates and increased light hydrocarbon formation (propane and lighter), the effect on catalyst life is much more serious than with existing isomerisation processes. However, these specifications can economically be easily met by a modification of the molecular sieve feed separation unit, so that the benzene and the dimethylpentanes (C7 plus components) are reduced to lower concentrations than those specified, simultaneously with the molecular sieve separation unit being improved overall with respect to investment and thermal efficiency because of decreased compression costs, condensing requirements, vapourization requirements, heating requirements, and better use of the molecular sieve capacity. British Petroleum, who have developed a pressure swing process for molecular sieve separation after originally developing and commercializing a displacement process for molecular sieve separation, have stated that the capital costs are lower for the pressure swing process (85 percent of the displacement process), whereas the operating costs are significantly lower (65 percent of the displacement process). The modification described herein modifies the more economical pressure swing molecular sieve separation process to further increase its advantages when preparing a C5/C6 feedstock for an isomerisation process. In the existing pressure swing molecular sieve separating processes, part of the normal paraffins are desorbed to serve as a purging medium, to assist in the removal of the non-normal paraffin feed components in the sieve voids and the easily removed components adsorbed on the surface of the sieve. The desorbed material from this purge operation is condensed (desorption is carried out at lower pressures than the adsorption), segregated from the other desorbed product, and recycled back to the feed wherein material must be revapourized and heated to absorption temperature for normal paraffin extraction. The quantity of normal paraffins lost in this purge operation as compared with the total normal paraffins desorbed in the desorption step varies with the required purity of the normal product. Typically, the normal paraffins lost to purge are substantially 15 percent of the total normal paraffins desorbed during the entire desorption step for segregating an isomerisation feedstock corresponding to 95 percent normal paraffins. Benzene and the dimethylpentanes are more strongly absorbed on the surface of the material so that these components tend to desorb later with the normal paraffin product itself. These components appear as an impurity in the normal paraffin product in considerably increased ratios over that which may be expected based upon its ratio in the normal paraffin-free portion of the feed.

Reference is now made to the accompanying drawings, in which: Figure 1 is a flow sheet of purge feed preparation; Figure 2 is a flow sheet of typical reactors in a commercial plant; Figure 2a is a flow sheet illustrating the separation step only; and Figure 3 is a flow sheet of an alternative arrangement for a large-scale refinery.

The apparatus illustrated in Figure 1 comprises five molecular sieve vessels 1-5 inclusive.

A crude unit overhead hydrotreater 6 receives incoming feed 7, consisting of a depropanised crude unit overhead fraction. Hydrogen is fed in at 8 and propanes, butanes and catalytic reformer feedstock are extracted at 9.9A and 10, respectively. A C5/C6 hydrotreated straight-run stream leaves the hydrotreater 6 at 11, to constitute the main feed stream. The main feed may optionally, be supplemented by means of a C5/C6 fraction from a catalytic reformer stage, as indicated by a dotted line at 12. A side-stream 13 is branched off from the main stream 11 at 14, and treated as described below.

The apparatus is illustrated in its purge phase. Accordingly, the flow stream 15 branched off at 16 is shown as a dotted line but corresponds to normal flow other than when the apparatus is in its purge phase.

In the illustrated purge phase, the main feed is fed through a heat-exchanger 17, in which it exchanges heat with denormalised product 18 taken off at 18 A from the vessels 1 to 3, inclusive, and then passes through a main feedstream heater 19 before entering the molecular sieve vessels 1 and 2, which are in their absorbing phase. In the illustrated phase, the sieve vessel 3 receives a purge stream resulting from the treatment of the sidestream 13, as discussed below.

The normal paraffin feed to isomerization 21 is withdrawn from the vessels 4 and 5 in their desorbing phase via low pressure desorption compressor stages 22 and 23, respectively. A stream 24 corresponding to desorption at atmospheric pressure is shown in dotted lines connected to the vessels 1-2.

The side-stream 13 is fed to a purge feed preparation stage, which comprises a benzene hydrogenation stage 25, and a two-stage nC6 - 2,4-DMP splitter 26. A bottom stream 27 is withdrawn from the second stage of the splitter 26 as reformer or molecular sieve separation feed. A propane plus stream is withdrawn from the first stage as overhead 27A. Hydrogen is supplied to the benzene hydrogenation stage 25 at 20 and withdrawn at 20A.

A C6-naphthene make-up feed to isomerization is withdrawn from the recycle stream 28 at 29. The remaining stream 28 passes through a purge feed heater 30 and is returned to the vessel 3 as stated above.

Various valve connections allowing the apparatus to be operated in purge feed and normal phases alternately are indicated, in addition to conventional equipment such as pumps and heat-exchangers. It will be noted in this connection that a number of valved branch lines are connected to the vessels 1 to 5, inclusive, allowing variations in the feeds to, and withdrawals from, these vessels.

The feed stream 15 may be used during the purge feed phase to make full use of the purge feed exchanger and heater and to maintain temperatures. Switching the feed pump suction is simpler than operating the purge feed heater discontinuously.

In order that the feed purge stream may be discontinued, the corresponding valve is closed and the valve allowing purge feed to join the main feed stream is opened. Suction is switched while closing the denormalised feed outlet which is open to the atmospheric desorption line.

Figure 2 shows a typical commercial plant.

The fresh feed from molecular sieve separation enters 20 at 31. Four reactor stages are indicated at 32 to 35, inclusive. The single reactor 32 is in lead reactor position and the reaction 33 in second reactor position. The two-column reactor 34 is in third reactor position and the three-column reactor 35 in terminal reactor position. An HC1 adsorber is provided at 36 and a by-pass stream at 37. The merged stream 38 passes through the reactor stages 32 to 35, inclusive of the isomerization reactor section. At the outlet of the reactor stages 32 to 35, there are provided hydrogen recovery and hydrogen chloride recovery towers 39 and 40, respectively. The latter is followed by a hydrogen chloride stripper 41, which feeds bottom product to the middle of an isomerate splitter 42 fom which isomerate product is recovered at 43. An overhead flow 44 from a cyclohexane solvent product stripper 45 is fed into the bottom end of the isomerate stripper 42, and a sidestream lean oil stream 46 is fed from the stripper 42 to an isopentane plus adsorber 47. Hydrogen and hydrogen chloride are recycled at 48 from the head of the absorber 47 to the bottom of the hydrogen chloride adsorber 36.

A cyclohexane product 50 is withdrawn at 49 from the bottom of the stripper 45. C6 naphthene make-up and isomerate splitter bottom liquid are recycled at 51. C6 naphthene make-up feed is added at 52 through a purge adsorber 53, and overhead recycle from the stripper 45 passes through a purge adsorber 54.

A sidestream 55 (shown in dotted lines) from the recycle stream 51 can be passed through a swing vessel 56, shown in its regeneration phase, and is recycled at 57 into the bottom of the HCI absorber 36. Overhead from the latter comprises at 58 chlorinating agent for a catalytic reformer and at 59 excess to caustic wash.

Hydrogen and hydrogen chloride make-up feeds are indicated at 60 and 61, respectively, entering the recycle sidestream 62, which is split off from the main recycle 48. Cooling by means of chilled water is indicated at 63 in the overhead stream to the HC1 adsorber 36 and 64 in the lean oil sidestream 46 fed from the splitter 42 to the isopentane plus absorber 47.

Further cooling is shown at 65 in the overhead stream from the hydrogen recovery tower 39 to the isopentane plus adsorber 47. Finally, stream heating and cooling devices 67 and 68, respectively, are included in the bypass 69 from the feed 55 to the swing vessel 56.

The diagram of Figure 2a is concerned with the separation step only. Figure 2a accordingly corresponds to the right-hand portion only of Figure 2 and again comprises a hydrogen recovery tower 39 an isopentane plus recovery stage 47, from which hydrogen chloride and hydrogen are recycled at 48, hydrogen chloride recovery and stripper stages 40 and 41, respectively, an isomerate splitter stage 42 and a cyclohexane solvent stripper stage 45. The isopentane plus recovery stage 47 is modified, as diagrammatically indicated, in that the feed 46 from the isomerate splitter 42 to the adsorber 47 is omitted. The recycle arrangement of the hydrogen chloride recovery tower is also modified as indicated.

The modifications indicated allow ready removal of inactive solid aluminium chloride complex formed from impurities in the C6 naphthene make-up feed, i.e. feed stream provided by benzene hydrogenation product separation; C5-C6 for the combined Isomeration unit and the C6 fraction for the hexane isomerisation unit.

The butanes formed as a by-product of disproportionation reactions are sufficient to allow the elimination of the lean oil sidestream 46 referred to above. AT 365 PSIA, 40 "F, the hydrogen and hydrogen chloride product will approximate to the following composition: Component Mols/Mols C5 plus feed entering lead reactor H2 0.020 HCI 0.120 C4H1o 0.006 Total 0.146 The recycle is accordingly less than 0.40 that of the normal make with either Cg, C6 or C5/C6 feedstocks. This limits the primary disproportionation reaction, namely, butane formation, so that disproportionation reactions are further limited without any inhibiting effect upon isomerization activity. Butanes recirculated are small in relation to the total feed and do not affect the C5 plus components entering the lead reactor. The separation fully uses the benefits of the higher operating pressure for the reactor system, and also provides the desired hydrogen partial pressure in the reactors to decrease catalyst deactivation.

Figure 3 shows an alternative arrangement for large-scale refineries.

The feedstream at 71 represents debutanised straight-run hydrotreater product. This is fed successively to a dehexaniser 72, a depentaniser 73, a molecular sieve separation unit 74 (see Figure 1 discussed above), a common HC1 absorber 75, an hexane isomerization unit 76, which is followed by a hydrogen chloride stripper as in Figure 2 and 2A, respectively, and an isomerate splitter 77 similar to the corresponding isomerate splitter 42 of Figures 2 and 2A.

The apparatus comprises a benzene hydrogenation unit, the construction and operation of which are broadly conventional, so that it is only indicated in highly diagrammatic form.

Catalytic reformer (C/C6) feed enters at 81 and hydroge make-up at 82. A propane plus overhead stream 83 includes hydrogen sulfide. The stream passes successively through a dimethyl pentane splitter 84 and a depentaniser stage 85. Recycle to the catalytic reformer is shown at 80. The hexane-containing stream from the catalytic reformer, leaving the benzene hydrogenation unit as bottom liquid 86 from the depentaniser column 87, passes through a purge absorber 88 before entering the hexane isomerization unit 76. The feedstream 89 to the latter unit 76 has passed through a static mixer 90 and filters 91.

A cyclohexane solvent product stripper 45 (as in Figures 2 and 2A) is again provided from which cyclohexane solvent product 50 is again withdrawn as bottom liquid. Bottom liquid 92 from the isomerate splitter 77 is fed as a drag stream through a further purge adsorber 93 into the column 45, and overhead is recycled at 94 to the splitter 77.

A catalytic reformer pentane stream 95 is fed as overhead from the depentaniser 87 to a deisopentaniser stage 98. Isopentane product is withdrawn at 99 from the desiopentaniser column 102. A feed stream 101, including liquid recycle, is connected between the base of the column 102 and a pentane isomerization unit 103, which includes a hydrochloric acid stripper as in Figures 2 and 2A. The outlet stream from the HCI stripper (not shown) enters the column 102 at 104.

An overhead stream from the top of the depentaniser column 73 and comprising straight-run pentane is also fed at 105 to the column 102. Filters 106 remove solid aluminium chloride complex formed by impurities in fresh pentane feed. Although they react at lower temperatures, they have residence in the lower portion of the deisopentaniser to enable most of the reactions to be completed.

Purge adsorbers are provided at 107 for the pentane isomerization unit 103. Purge from the pentane isomerization is fed into the depentaniser 87 through the line 108.

Catalytic reformer feedstock is withdrawn as bottom liquid 109 from the base of the dehexaniser column 72. Straight-run hexanes are fed from the base of the depentaniser column 73 as main feed 110 to the molecular sieve separation unit 74 (see Figure 1). A denormalised gasoline product is withdrawn at 111 from the latter unit 74, and normal paraffins (primarily normal hexane) are fed at 112 to the HCI adsorber 75. A purge feed to the unit 74 is provided at 113.

Finally, dimethyl butane product is shown at 120 as overhead product leaving the isomerate splitter 77.

It is to be noted that all the fresh feed contacts the liquid recycle containing disolved aluminium chloride so that most of the inactive solid aluminium chloride complex formed by impurities in fresh hexane feed can be filtered and prevented from entering the reactor vessels themselves. Contact of the fresh feed with the liquid recycle is provided by the static mixer 90, as indicated above, to provide thorough mixing and residence before filtering any resulting solid complexes formed.

Figure 1 illustrates the molecular sieve pressure swing step. The operation is as follows.

The loss of normal paraffin in the initial disorption purge together with the associated utilities is eliminated by substituting a feed stream for adsorption in place of the feed stream containing benzene and C, plus components and possibly other impurities such as olefines.

The main feed is discontinued from entering the vessel being prepared for desorption and enters other molecular sieve vessels in the adsorption phase. This substituted feed stream serves as a urge medium even though the bed continues to absorb normal paraffins to produce a denormalized product. The purge feed stream substituted for the main feed stream is free of aromatics, Cl plus hydrocarbons and olefins. Continuing the same flow pattern with the purge feed stream as the main feed stream enables a more ready displacement of the main feed stream non-normal components in the voids of the sieves and on the surface of the sieves. The purge feedstock should preferably contain normal hexane as this component deposits more readily and displaces any normal pentane previously deposited forward towards the top of the bed in a normal manner. A purged medium not having normal hexane would require a greater quantity for displacement of the main stream components, and would also tend to serve as a stripping medium for the normal hexane adsorbed component to introduce more normal hexane into the denormalized product. This is undesirable as normal hexane has a considerably lower octane value than normal pentane. It is also desirable that the normal hexane be concentrated towards the feed end of the molecular sieve vessel as normal hexane is less readily desorbed than normal pentane.

In desorption, the feed end has the lowest pressure (pressure drop consideration) and also benefits from the normal pentane being absorbed higher up in the molecular sieve bed during desorption to serve as a sweep medium for the normal hexane (lowers normal hexane partial pressure). As primarily non-normal paraffin component, the purge medium of the process of the invention would be filling the voids of the sieves when desorption commences. This is also beneficial as these non-normal components also serve as a sweep medium for the normal paraffins to be desorbed.

Although other feedstocks may be employed as a feedstock for the preparation of the purge medium, it is assumed for the purpose of illustration only that part of the main feed stream feestock to the molecular sieve separation unit (such as 5 to 10 percent of the total stream to be processed) is subjected to treatment for the removal of benzene and the C7 plus paraffins in preparing the purge feed stream. The following treatment may also be employed as an alternative to molecular sieve separation of the normal paraffins for preparing isomerisation feed to the i because of the closeness to normal hexane (more volatile key component). Thus, the overhead product, in addition to supplying the small purge feed stream requirements for the molecular sieve separation unit, also supplies the C6 naphthene requirements for the isomerisation process later described. The cyclohexane component leaves with the bottom product. The bottom product can either be charged to a catalytic reformer if BTX is desired (cyclohexane is converted almost quantitatively into benzene) or to the molecular sieve main stream feed for recovery of tegasoline values by the removal of normal hexane.

Using this feed stream which is free from benzene, C7 plus components and olefins as a purge, the main feed stream flow is shut off and the purge stream introduced so that the adsorber continues to function as a normal adsorber as long as the purge stream is continued. After sufficient purge, the molecular sieve vessel is then desorbed. Because all the material in the purge stream is fully satisfactory as an isomerisation charge, i.e., sufficiently free of aromatics, C7 plus components and olefins for all the requirements of the isomerisation process of the invention, which are generally lower than those of the commercial isomerisation processes to be met, the void space material may be included in the normal paraffin product. This modification of the pressure swing molecular sieve in producing a normal paraffin feedstock for isomerisation means that the entire desorbed material becomes part of the isomerisation feedstock with no less of normal paraffins.

Desorption is the controlling step in molecular sieve separation so that approximately a 20 percent foverall gain in molecular sieve processing capacity, i.e., net feedstock processed per given quantity of molecular sieves, may be obtained over existing pressure swing molecular sieve separation processes. Eliminating the initial depressurization to purge pressure makes the process more continuous and less subject to contamination.

The pressure swing molecular sieve separation process of the invention which is preferably used to prepare the main feed to the isomerisation unit has a desorption equipment design which is more compatible with the condensing and compression operations.

It is possible to have different vessels operating at difference pressures so that the minimum compression equipment is used continuously. Furthermore, it is not necessary to segregate part of the desorbed product for recycle to adsorption. For example, Figure 1 shows three molecular sieve vessels in various stages of desorption. The highest pressure desorption is at or near atmospheric pressure levels so that the stream can enter the final condenser without compression. This atmospheric or near atmospheric pressure desorption occurs after the purge phase in adsorption is completed and the process flow is discontinued. The purge feed phase requires only a short part of a normal cycle which limits the purge feed stream material required. The desorption vessel is then isolated from the denormalized product stream by means of mechanically or pneumatically operated valves and the valve to the atmospheric condenser is opened. Each pressure service in desorption has its own air cooler as the molecular sieves operate at approximately 600"F in the vapour phase, both in adsorption and in desorption, so that considerable cooling is required before condensing ever occurs.

Two other pressure levels are shown in the diagram for desorption, the intermediate pressure level corresporiding to the next cycle during which the vessel undergoes purging and atmospheric desorption can proceed. The intermediate pressure desorption vessel then proceeds to the final desorption pressure cycle. After reaching the lowest pressure in the final desorption cycle, the vessel proceeds to the adsorption cycle. It will be noted that, by having three pressure levels for desorption, the swing in pressure which a given compressor must handle is less. Compressors more suited to the volume can be provided. The horsepower required to handle a given desorption duty is less as interstage cooling is automatically provided between pressure levels for any material not condensed. The use of several pressure levels is conducive to more efficient condensing as normal pentane is more easily desorbed than normal hexane. Therefore, normal pentane having the lowest condensing temperature is primarily desorbed at the neat atmospheric condensing level whereas normal hexane is primarily desorbed at the lowest pressure level (the higher molecular weight is of advantage in increasing the gas density and polytropic compressor efficiency).

The pressure swing molecular sieve separation process of the invention, with more defined pressure levels in desorption, assists more efficient heat recovery by heat-exchange in place of cooling in desorption as now practised almost exclusively in the existing processes i.e. heat is recovered more efficiently instead of air cooling similar to that shown in Figure 1 with the denormalized effluent product. This further reduces the heating requirements.

As any refinery interested in octane values would normally have a catalytic reformer, the more efficient pressure swing molecular sieve separation process of this invention ensures that, in addition to the hydrotreated straight run, the catalytic reformer C5/C6 fraction is also processed as part of the main stream. Most of the lower octane non-aromatics are contained in the C5/C6 fraction, and these fractions only have an octane number of approximately 65 RONcl. This C5/C6 fraction from the catalytic reformer may be easily separated in a manner similar to the hydrotreated straight-run fractions. Unlike the hydrotreater straight-run feed, the catalytic reformer-stabilized reformate contains practically no sulphur or water. However, some olefins are present so that it would be undesirable to charge this as a feed stream directly to the isomerisation process of the invention. Small quantities of olefins are formed by the nature of catalytic forming as intermediate products.

Fortunately, olefins remain on the molecular sieve during desorption of the normal paraffins. After long periods on stream, absorbed olefins are removed (as carbonaceous deposits) in regeneration. Regeneration consists of oxidation with controlled quantities of oxygen in a recirculating nitrogen stream similar to that practised in catalytic reforming.

Benzene in reformate at the 95 RONcl level is always considerably in excess of any possible azeotrope that may be formed with the components in the C5/C6 fraction. The dehexanizer would be separating a normal hexane/benzene azeotrope volatile component from benzene itself as the less volatile key component. This provides a reasonable fractionation requirement for removing 95 percent of the normal hexane from the entering reformate feed. Any methylcyclopentane entering (always present in significant quantities) is a distributed component with approximately 80 percent removed overhead as a methycyclopentane/benzene azeotrope. Most of the benzene would remain with the dehexanizer bottom product which would have an RONcl value in excess of 102 even for the moderate severity levels required at the reformer in producing reformate for regular gasoline blending. The product thus not only lends itself to future chemical recovery of the B-T-X aromatics, if desirable, but also simplifies scheduling the catalytic reformer operations with improved yields overall in producing reformate of the necessary octane for blending into premium gasoline. Upgrading the normal pentane and normal hexane components from the straight run and the catalytic reformer in the manner described herein preserves the saturates so that the octane levels of today may be manufactured even if required to be lead free. Furthermore, the aromatic and olefin contents of the gasoline pool would if anything actually decrease, unlike more severe catalytic reforming which increases the aromatic content of the gasoline pool.

The isomerisation process described herein operates at near ambient reactor temperature levels in the liquid phase so that no heating or vaporization is required for heating the isomerisation feed to reactor temperature levels. Cooling only is required for the reaction system on process but this is normally effected by cooling tower water. Apart from cooling the feed, cooling water is required corresponding to a total of approximately 50"F overall to allow for the heat of reaction in the conversion of normal pentane in the feed to isopentane, and the conversion of normal hexane and other hexanes entering primarily into dimethylbutanes, as this constitutes the normal C6 paraffin product. A small amount of methylcyclopentane in the feed of the naphthene makeup stream (usually corresponding to less than 5 percent of the total isomerisation fresh feed) is converted to cyclohexane as all C6 naphthenes are usually withdrawn as a cyclohexane solvent product. Reference should now be made to Figure 2, which is illustrative only; if e.g., the cyclohexane solvent is not desired as a product, the drag stream may be returned to the normal hexane/ dimethylpentane splitter of the benzene hydrogenation unit which prepares the C6 naphthene makeup stream. Variations in the product separation are not of major significance as the normal pentane may, for example, be recycled for conversion into isopentane, but this separation operation is seldom justified.

The catalyst employed in the process of the present invention may be any Friedel-Crafts halide catalyst suitable for liquid hydrocarbon phase contacting, and promoted. with a hydrogen halide recirculating stream to the reactor. The preferred catalysts are either aluminium chloride or zirconium chloride, with a suitable support such as activated alumina having a discrete particle size which is preferably below 1/16", and promoted with recirculating hydrogen chloride.

These catalysts possess good isomerisation activity and have suitable melting points at reactor temperature levels so that they remain as solids. The use of hydrogen chloride as a promoter to enhance catalytic activity for chlolride catalysts of the Friedel-Crafts type is so well established fo aluminium chloride, that H Al C14 is generally considered the active component, although, to the best of my knowledge, the latter compound has not been isolated. Friedel-Craft catalysts have a small but significant amount which dissolves in the hydrocarbon phase. Zirconium chloride dissolves less than aluminium chloride. However, while zirconium chloride significantly lowers the solubility of the catalyst in the hydrocarbon liquid phase product leaving the reactor, it is much more expensive than aluminium chloride and is not significantly more active at this state of the art. Furthermore, the process of the invention provides a solution to the solubility problem by returning most of the catalyst to the reactor in the naphthene and unconverted C6 paraffin liquid recycle, which is the bottom liquid stream from the isomerate splitter.

Aluminium chloride is assumed to be the active Friedel-Crafts type catalytic agent because a considerable quantity is required in the reactors, and aluminium chloride is relatively inexpensive and readily obtainable commercially in high purities. The large reactor volume is one of the disadvantages of the process of the invention. However, the reactors are constructed of carbon steel and designed for moderate temperature (maximum operating temperature is less than 325"F - required in regeneration only).

Unlike the noble metal-containing hydroisomerisation catalyst, catalyst cost per unit volume is comparatively low. The catalyst inventory required for the process of the invention based upon aluminium chloride on a bauxite support would be less than for the platinum cost alone for existing hydroisomerisation processes. The cost of the additional reactor volume for the process of the invention should therefore be offset by the cost of the more expensive base support for the platinum catalyst together with the heating system required for vaporizing the feed and heating the reactants to the temperature levels required in conventional hydroisomerisation processes.

Carbon steel is an acceptable material of construction for most Friedel-Crafts chloride catalysts provided water is prevented from entering the reactor system whether the system is promoted with hydrogen chloride or not. This is because dry hydrogen chloride has a corrosion rate of less than 0.0002 inch per year at moderate temperatures with carbon steel.

BP's isomerisation reactor data indicate that this order of magnitude holds for temperatures up to approximately 320 F. Even if hydrogen chloride is not present, the corrosive agent aluminium chloride when water vapour is present, continues to be moist hydrogen chloride because formation of hydrogen chloride with aluminium oxide byproduct occurs overall if any water vapour is present. Anhydrous aluminium chloride is extremely hydroscopic and precautions must be taken in handling this material. However, this chemical is widely employed commercially for other commercial purposes so that both the equipment and the technique are well established.

Kinetically, the presence of any water in the feed is extremely undesirable as the surface of the anhydrous aluminium chloride becomes glazed with a coating of aluminium chloride hydrate. The deactivation rate is dependent upon the contaminants entering in the feed. As the feed preparation outlined herem should result in lower contaminants than previously used in pilot planting the aluminium chloride process, the deactivation rates may decrease in commercial service. Another advantage of the upflow reactor proposed as a preferred form for the process of the invention is that it may utilize smaller support particles for the support of the aluminium chloride catalyst, and this should further increase the isomerisation performance as, kinetically, the reaction is controlled by access to active aluminium chloride or other Friedel-Crafts type catalyst. Furthermore, the smaller catalyst particles should result in a more gradual deactivation decline in performance.

Figure 2 shows a typical arrangement of the reactors anticipated for a commerical plant.

The reactors are preferably operated upflow and are of equal reactor volume with each reactor corresponding to approximately 1 LHSV based upon feed. Each reactor vessel is valved so that it may serve in any process operating state (lead position, second position, third position and terminal position) as well as in the regeneration position. The preferred reactor arrangement is for one reactor vessel normally to serve in the lead reactor position, one reactor vessel in the second reactor position, two reactor vessels in the third reactor position and three reactor vessels in the terminal reactor position. Such an arrangement suits the recommended descending temperature pattern and the cooling of the isomerisation reaction heat. The lead position operates with a typical reactor outlet of 118"F (the highest reaction temperature). After saturation in the aluminium chloride saturators, the reaction mixture is further cooled to enter the second position reactor with a typical 114"F reactor outlet temperature. After further cooling, the reactant mixture enters the third reactor position which operates with a typical 108"F reactor outlet temperature. Finally, after further cooling, the reactant mixture enters the terminal reactors where the reactor product leaves typically at 1000F. Reactor outlet pressure is, typically, 300 psig. Reactor outlet pressure is set primarily by that convenient to the desired hydrogen partial pressure and hydrogen chloride recirculation as a hydrocarbon liquid phase can be maintained at a lower pressure. Reactor vessel design requirements are determined by the conditions required in the regeneration procedure, not by the operating pressure of the process.

The process of the invention utilizes a combination of recirculating hydrogen (typically.

0.02 mol of hydrogen enters the lead reactor with each mol of C5 and heavier hydrocarbon feed entering), together with a limited content of naphthene for feedstocks containing hexanes, corresponding typically to 5 to 6 weight percent of the C5 + hydrocarbon feed entering the lead reactor, so as to suppress disproportionation reactions and cracking of the higher carbon components, which become prominent in Friedel-Crafts halide catalysis with Cs and C6 paraffin feeds unless a suppressant is present in the reaction zone.

The use of hydrogen as an inhibitor to reduce disproportionation reactions is well established. Shell reported the successful use of hydrogen as a suppressant for pentanes (commercially demonstrated in the Shell Curacoa refinery). With this inhibitor for their liquid phase isomerisation process (operating at higher temperature, 176 to 2120F, with aluminium chloride dissolved in molten antimony trichloride), Shell reduced the total make of propane and lighter hydrocarbons to less than 0.01 weight per cent of the feed. The latter "make" value is an indication of the higher carbon number disproportionation products formed as C7 and higher carbon number hydrocarbons crack readily; isobutane, which is the primary disproportionating reaction product has negligible carbonaceous deposits associated with its formation, but the secondary reactions such as C4 carbonium ions combining with a C5 activated species, or another olefin, result in Cx hydrocarbons where x represents the number of carbon atoms which undergo further cracking reactions to form olefins) to less than 0.1 weight percent of the feed. The beneficial effect of C6 naphthenes in suppressing disproportionation reactions for a C6 feed on a laboratory and pilot plant scale has also been reported for Friedel-Crafts halide catalysts with their respective H2 promoters, and with the hydrocarbon in the reaction zone in the liquid phase. Naphthenes, without any H2 present in the reaction zone, but at considerably higher naphthene concentrations than those proposed, have been successful in suppressing cracking and disproportionation reactions with feedstocks containing C6 paraffins over a long period.

The mechanism of suppressing disproportionation reactions is only conjecture at this moment, but the operating mechanism for all Friedel-Crafts type catalysts is based on carbonium ions. Isoparaffins are generated from normal paraffins by a combination of hydride abstraction and hydride ion transfer reactions with skeletal rearrangement reactions in the absorbed state. The combination of hydrogen and naphthenes supplements the deficiencies of the individual suppressants alone. For example, the C6 naphthenes are always in the liquid phase and in contact with the catalyst because of their surface adsorption preference. The C6 naphthenes provide better thermodynamic equilibria stability as the catalysts are also active for isomerisation from rings into chains and vice versa but the rate is very slow at low temperatures. Hydrogen is unable to supply the foregoing benefits which C6 naphthenes provide. On the other hand, hydrogen can suppress carbonaceous deposits by saturating activated species or olefins from the reactions which occur so that these do not condense further into polymers which ultimately form carbonaceous deposits on the catalyst. Hydrogen also ensures an amply supply of hydride ions. These benefits which hydrogen is able to provide C6 naphthenes cannot supply.

The combination of hydrogen with a limited quantity of C6 naphthenes has kinetic benefits apart from suppressing disproportionation reactions. This consists of enhanced isomerisation activity for the pentanes, which should lower the normal pentane in the isomerate product further than projected as thermodynamic equilibria permit considerably lower values. Higher concentrations of C6 naphthenes appear to have the greatest depressing effect upon the pentane isomerisation performance in mixed C3 / C6 feedstocks as compared with equivalent results for the same space velocity and reactor temperature when processing pentanes alone with hydrogen as the suppressant. Reactor volume is better employed because the higher concentration of C6 naphthenes which would be necessary if used alone displaces paraffin feed which could utilize the volume occupied by additional naphthenes for isomerisation reactions, Furthermore, naphthenes at concentrations higher than 6 weight percent of the feed appear to inhibit C6 paraffin isomerisations unduly, unlike hydrogen. This discovery appears to be universal and noticeable in hydroisomerisation, and appears to explain the difference between varying feedstocks in their performances under equivalent reactor conditions.

The process of the invention provides an economical solution to the solubility of Friedel-Crafts type catalysts in liquid phase processes. Shell's liquid phase process after using the kinetic advantage of the hydrocarbon liquid phase in the reactors then vaporizes the entire reactor hydrocarbon effluent in order to recover the catalyst. This is an expensive solution as it requires again a condensation of the vaporized product before proceeding to hydrogen chloride separation.

Although the reactors operate with the hydrocarbon in the liquid phase, Shell's liquid phase process possesses the same utility disadvantage that vapour phase isomerisation processes possess, namely, the vaporization of the entire feed.

The terminal reactor of the process has the lowest reactor outlet temperature, which reduces the amount of Friedel-Crafts catalyst leaving with the reactor product because solubility decreases with temperature. The more moderate liquid phase upflow velocity in the terminal reactor position, due to the three reactors in parallel, also prevents any entrainment of catalyst.

The lower temperature and higher reactor volume of the terminal reactor position kinetically enhances the isomerisation product octane values because of the effect of lower temperatures on the thermodynamic equilibria. Operating the reactors upflow tends to permit any inactive aluminium chloride solid complex formed to be suspended in the liquid phase and thereby to cause less deactivation as the reactants have a freer access to the active aluminium chloride catalyst. Impurities in the fresh feed react with dissolved active aluminium chloride in the hydrocarbon liquid phase when they contact the recycled liquid stream containing dissolved aluminium chloride. The impurities react much faster than the isomerisation reactions. The process of the invention, by operating upflow in the reactors, has the same pressure drop at the begining and at the conclusion of a run when shut down for periodic maintenance. In contrast, in downflow operation, these solid complexes would cause an increase in operating pressure with time, by filling the void spaces at the bottom of the reactor vessel used for the passage of the liquid phase, and could cause flow maldistribution within the reactors, which lowers catalytic contacting efficiency. The inactive aluminium chloride complexes would then be compressed against catalyst particles with active aluminium chloride. By preventing access of the reactants in the liquid phase to the active sites occluded, catalyst deactivation would be increased by more than the amount of solid inactive aluminium chloride complexes formed with impurities.

The process of the invention preferably limits reboiling temperatures below 250"F. This enables the main process heating requirements to be supplied with low level heat, such as 35 psig steam, which is generally in excess or could be economically recovered in most refineries. The limited reboiling temperatures ensure that catalyst is not subject to undue deactivation before evantual return to the reactor system. In Figure 2, this is accomplished by staged flashing to recover hydrogen and hydrogen chloride with the hydrogen chloride stripper operating at approximately 75 psia, which is considerably lower than usual.

Operating the hydrogen chloride stripper at low pressures enables hydrogen chloride to be economically and satisfactorily reduced to low concentrations in the isomerate product similar to those now tolerated in reformate streams from catalytic reformers, for which there have been no reported undesirable effects.

Hydrogen chloride is recirculated to the lead reactor in a ratio of 0.12 mol of hydrogen chloride per mol of Cs plus hydrocarbon feed entering the lead reactor. Flashing in stages saves compression horsepower as all the hydrogen is stripped and removed at terminal reactor outlet pressure. Use is made of the lower pressure material when compressed to serve as a stripping medium, which reduces the hydrogen chloride entering the hydrogen chloride stripper, so that a comparatively small amount must be compressed from the lowest pressure level.

A small part is bled as necessary from the hydrogen and hydrogen chloride recycle gas stream to the hydrogen chloride absorber, in order to prevent any undue concentration of methane or ethane in the reactor system. The hydrogen chloride absorber uses part of the fresh feed as a lean oil to recover the hydrogen chloride. Purged gas streams from regeneration are also routed to this absorber to recover the hydrogen chloride, but these streams are only periodic. The entire main fresh feed is used as a lean oil when these streams are purged. The normal hydrogen chloride absorber offgas is used as a catalyst chlorinating agent (generally required for promoted catalyst) in the catalytic reformer. Any excess quantity of offgas (anticipated during regenerations only) is caustic washed to remove traces of hydrogen chloride. Anhydrous hydrogen chloride make-up requirements are low. The anhydrous hydrogen chloride has any carbon dioxide removed before entering as a make-up to the reactor system. Most commercial anhydrous hydrogen chloride typically contains 0.005 mol fraction carbon dioxide which increases catalyst deactivation.

The foregoing enables elimination of the conventional caustic and water wash for removing the dissolved aluminium chloride in the hydrogen chloride stripper bottoms. This eliminates the waste product associated with this conventional messy operation. Environmental considerations are now often more important than the savings in the additional aluminium chloride, anhydrous hydrogen chloride, and caustic required with the conventional solution. The elimination of the conventional washing of the hydrogen chloride product stream provides insurance against water being introduced into the reactor system with the hydrocarbon liquid recycle stream as wel as eliminating the drying equipment and operation conventionally required. Dissolved aluminium chloride is returned to the reactor system in the liquid hydrocarbon bottom product stream from the isomerate splitter, which is recycled to the lead reactor.

Precipitation of aluminium chloride in the separation system before any pumps is unlikely because the solubility of the aluminium chloride increases with an increase in temperature.

Aluminium chloride has a low volatility and will therefore always proceed to the tower bottoms. The higher temperature of the tower bottoms compensates for the decreasing amount of hydrocarbon liquid phase in which the catalyst leaving the reactor is dissolved as the isomerate splitter bottoms contains the methylpentanes and higher boiling components in the terminal reactor product. All the feed streams are mixed with the liquid recycle being returned to the reactor before cooling, in order to permit most of the impurities in the fresh feed to react with the dissolved aluminium chloride present in the liquid recycle. These reactions and filtration in the liquid phase are preferably complete before the introduction of the gas streams and final cooling to a temperature which corresponds to the lead reactor outlet temperature being at 118 F. The filtering removes any solid aluminium chloride complexes formed by impurities introduced in the fresh feed and reacting with the dissolved aluminium chloride in the liquid recycle. In this manner, a significant part of the inactive aluminium chloride is prevented from entering the reator vessels, so that the periods between regenerations and evantual catalyst replacements are prolonged.

A drag stream is withdrawn from the isomerate splitter bottoms and passed upwards through an adsorber bed containing an adsorbe contacting tower, where it is withdrawn.

The process of the invention utilises filters to remove solid aluminium chloride after sufficient mixing and residence has been provided for the dissolved active catalyst in the liquid recycle to contact the fresh feed where the trace impurities are reacted, to form a solid aluminium chloride complex. As a major part of these impurities react, and the resulting inactive solids are removed, outside the reactor vessels themselves, the catalyst deactivation within the reactor vessels is significantly lower than in the Friedel-Crafts halide isomerisation catalyst systems which introduce feed directly to the reactor (a characteristic of fixed bed isomerization systems up to the present).

Reycle of catalyst, according to the present invention will improve the low temperature process based upon aluminium bromide (just discussed), since aluminium bromide is considerably more expensive, and has a higher solubility, than aluminium chloride, catalyst recovery is even more important.

Three to four years between plant turnarounds for mechanical maintenance may be considered usual for the process of the invention as the plant has no fired heaters, operates at low temperature, and Is mechanically simple. Long onstream periods are convenient for the reactor system as the fewer the shutdowns, the lower the probability of moisture being introduced accidentally into the reactor vessels. Catalyst regeneration and catalyst replacement, when required, are performed onstream. Reactor vessels are individually regenerated onstream in the swing vessel position with an estimated 15 days between regenerations of the reactor vessel. The regeneration procedure requires approximately 4 to 5 days. Catalyst deactivation is gradual with this regeneration frequency. An estimated 0.005 fraction of decrease of the 2,2-dimethylbutane concentration in the total C6 paraffins leaving the terminal reactor may be expected per month. One may commence periodically to replace spent catalyst in a vessel with fresh catalyst after approximately 1 year onstream from initial startup (the only startup where fresh catalyst would be present in all reactor vessels), in order for the catalyst inventory to approach a steady state isomerisation activity.

This may be done by replacing the catalyst in one of the reactors with fresh catalyst every 1.5 to 2 months instead of regenerating. This should offset the gradual decline in isomerisation activity with onstream time that cannot be fully restored by the regeneration procedure. Having a separate vessel that serves as a swing vessel in regeneration service permits periodic replacement of spent catalyst with fresh catalyst without interfering with onsteam processing of feedstock. Actual catalyst replacement depends upon actual commercial performance in terms of once through concentration of the isopentane in the total C5 paraffins and the concentrations of the 2,2-dimethylbutane in the total C6 paraffins leaving the terminal reactor. The desirability of maintaining isomerisation activity at a high level depends upon the amount of feedstock to be processed and the octane requirements of the refinery. With the feed preparation proposed according to the invention, the actual performance may be expected to be better than that projected, as pilot plant background data were obtained with much higher contaminant levels.

The isomerate splitter, together with the other separation equipment including the liquid hydrocarbon recycle stream to the reactor and the reactor vessels, is designed on the basis of the 2,2-dimethylbutane being an 0.40 weight fraction of the C6 paraffins leaving in the terminal reactor. This design flexibility for separating the C6 paraffins into a high octane component permits ample time to determine the commercial deactivation rates applicable to a particular plant because catalyst deactivation is largely dependant upon the impurities introduced by the fresh feed components together with any in the hydrogen and hydrogen chloride makeup streams. Having an isomerate splitter that separates the C6 paraffins into an overhead consisting primarily of 2,2-dimethylbutane results in a product with a motor clear octane rating higher than the research clear octane rating. Today's gasolines are generally limited by motor octane ratings in obtaining satisfactory road octane perform ance. The 2,2-dimethylbutane component has a higher octane rating for the motor method than for the research method so that full advantage should be taken of the process's capability of upgrading the C6 paraffins into this component economically, with comparatively low utility cost, as recycle to the reactor is never vaporised as in the existing hydroisomerisation processes. The separation in the isomerate splitter is much easier than in an isopentane/normal pentane splitter, with 2,3-dimethylbutane being a distributed component. The recycle of the methylpentanes assists in returning the dissolved catalyst to the reactor system without any precipitation. The conservative design for the isomerate splitter enables the production of a constant octane product in spite of the fluctuations resulting from the freshly regenerated reactor being placed on line and of the fact that, immediately before regeneraton, rearrangement is scheduled, together with the gradual onstream decline in once through performance for the 2,2-dimethylbutane fraction. With C6 paraffins produced primarily as dimethylbutanes, the octane (both research and motor) of the isomerate product should never decrease below 87.5 clear, providing isomerisation activity is maintained at the design level with planned replacement with fresh catalyst. The principal octane variation is the ratio of the pentanes to the hexanes introduced with the feedstock.

All C5 fractions entering the reactor appear in the terminal reactor product and in the isomerate product. Pentanes are the lower octane paraffins in the isomerate product but should still have a research octane in excess of 86.8 clear under steady state isomerisation activity. Any cyclopentane entering the isomerisation reactor appears unchanged in the terminal reactor product and forms part of the isomerate product. This contrasts with hydroisomerisation wherein the naphthenes are subject to losses. Pentanes, provided they meet the feedstock specifications, may thus be charged to the process of the invention with full recovery of the octane value of the cyclopentane component.

The fact that naphthenes are not destroyed in the isomerisation process of the present invention, in contrast to existing hydroisomerisation processes, means that alternative arrangements for large-scale refineries become practical. Many variations are possible, but Figure 3 of the accompanying drawings shows a typical illustration which enables a large-scale refinery to increase the isomerisation product octane at a modest overall incremental expenditure, so that the cost per octane-barrel of product incrementally provided becomes competitive.

The apparatus illustrated employs two isomerisation units, one for the pentanes fractions, a pentane isomerisation unit, and one for the hexanes, including the primary normal hexane feed from the molecular sieve separation unit. Having a separate pentanes isomerisation unit enables upgrading of most of the normal pentane in both the straight run and catalytic reformer sources into an isopentane product containing less than 2 percent normal pentanes (91.7 RONCI, 89.7 MONCL), as compared with the combined pentanes/hexanes isomerisation unit having the isopentane in the isomerate product with 17 to 18 percent normal pentanes or from 4.5 to 4.8 octane numbers lower. Having a separate pentanes isomerisation unit enables an increase in the once-through conversion per pass to where the isopentane in the C5 paraffins leaving the terminal reactor is increased to approximately 0.86. This high reactor effluent can thus be economically increased in a deisopentaniser to a less than 2 percent normal pentanes overhead product, because of the required reflux ratio. Having such a deisopentaniser enables separation of the pentanes (easily separated from the hydrotreater straight run feed and from the benzene hydrotreater product), so that the original isopentane is reduced to approximately 8 percent in the bottoms, and this corresponds to the comparatively moderate reflux ratio for upgrading the reactor feed product. The expensive utility fraction would be to decrease the bottom isopentane content to a low value, but with an incremental conversion per pass of 0.76, as a fraction of the total pentanes feed, this is not required. This means that the heat compressor, which supplies all of the heat required (some common surface as most of the condensing surface), together with a trim cooler, is thermally efficient, particularly when high flux tubing for both sides of the reboiler are employed. The trim cooler supplies the heat balance for the system. A higher incremental conversion per pass results with pentanes, even for the same space velocity, because a pentanes feed can employ hydrogen by itself to suppress the cracking and disproportionation reactions. This eliminates the inhibition due to naphthenes, which are required for the hexanes isomerisation unit, or for a mixed pentanes/hexanes feed.

The introduction of the pentanes feed into the deisopentaniser overcomes the low liquid recycle which would result in the absence of this feed, causing precipitation of the dissolved solids from the reactor, as the bottoms temperature roughly corresponds to the terminal reactor outlet temperature. Introducing the pentanes feed into the deisopentaniser means that the liquid recycle is practically the same as the terminal reactor liquid phase which contains the dissolved catalyst, so that precipitation is avoided. Furthermore, separation by distillation for the pentanes fraction, is overall economically competitive with molecular sieve separation. for the separation of normal hexanes from the cyclics and dimethylpentanes, the molecular sieve separation has a distinct economic advantage. The pentanes can be readily separated from cyclopentane, which has approximately the same boiling point as neohexane or 2,2-dimethylbutane. On the other hand, the significant cyclic content is associated with the straight-run hexane fractions, whereas the catalytic reformer C5/C6 fraction has only a limited cyclic and dimethylpentane content associated therewith, because of the nature of catalytic reforming.

The catalytic reformer C5/C4 feedstock, unlike the straight-run hydrotreated feedstock, has a significant olefin content, typicaly in the range of from 0.5 to 1.0 weight percent of the feed. As liquid phase benzene hydrogenation requires comparatively little energy and also saturates the olefins, the arrangement shown in Figure 3 treats the whole catalytic reformer C5/C6 fraction, serving to hydrogenate both the benzene and the olefins to less than 10 ppm in the product. This ensures a significant reduction in the cost of the molecular sieve separation unit itself, and also ensures that the hydrotreater hexane fraction (main feedstreams) should afford a molecular sieve life compatible with the depreciable life of the molecular sieve separation plant itself. The molecular sieve inventory of a molecular sieve separation unit whether treating the cgc6 feed or the C6 feed itself is generally at least 20 percent of the battery limit investment. In general, the molecular sieve life experienced with C5 to C7 feedstocks is approximately 8 years, but the life depends upon the frequency of regeneration. This itself depends upon certain sulphur compounds being present in the feedstock and the olefin content of the feedstock. Olefins in straight run feedstocks can be considered negligible in the future as almost all catalytic reformers may be expected to replace their catalyst with promoted catalytic reforming catalyst because of its appreciable yield advantages and longer onstream times with semi-regenerative reactor arrangements.

As the promoted catalytic reformer licensors generally require below 0.5 ppm sulphur for their catalyst life guarantees, not only will the straight run sulphur be negligible, but the olefin content will also be negligible. This is because, if a hydrotreater is operated at too high a temperature, olefin traces appear. These traces react, and combine with the hydrogen sulphide present in the reaction zone to form mercaptans. At 0.5 ppm sulphur content, not many mercaptans can form without exceeding the sulphur specification. The bulk of Friedel-Crafts halide isomerisation catalyst experimental work was performed when 0.002 weight percent sulphur represented a well prepared feed. As Friedel-Craft catalysts react almost quantitatively with feedstock impurities, high catalyst make-up rates are normally associated with the use of these catalysts for isomerisation. The arrangement shown in Figure 3, for large scale refineries, whose volume can justify two separate isomerisation units, provides a better adaptation of the limitations of the respective units to the properties of the feedstocks with a vlew to increasing the pool octane. The molecular sieve inventory for the separation unit, as shown, has an expected life equivalent to the economic life of the plant itself. Thus, sulphur and olefin contents, the factors which cause a gradual decline in the adsorptive properties of molcular sieves, until regeneration is required, are such that long period are permitted between regenerations, the regeneration or replacement periods being in excess of the plant's economic life. Although both molecular sieves and the catalyst of the process of the present invention benefit from the lower contamination, with longer onstream lives between regeneration and evantual replacement, a higher levels of contaminant can be tolerated without affecting immediate performance, but the periods between regeneration and evantual replacement are then shortened.

Regeneration equipment may be of conventional type, but each process unit has its own swing vessel because of the valving requirements. The regenerations typically occur alternately every 7.5 days, but, if necessary, the frequency may be increased to once every 5 days so that a sufficient margin should exist to recover from anticipated contamination upsets, such as a temporary increase in dimethylpentane content. The hydrogen chloride absorber is located at the hexane isomerisation unit to handle purges from that unit, and also to cope with recovery of hydrogen chloride from regeneration purges as only one reaction vessel is regenerated at one time. This enables advantage to be taken of the higher boiling normal hexane feed as a lean oil, which, unlike the pentanes feed, would never be introduced into the isomerate splitter. This is because any normal pentane in the feed would contaminate the isomerate product. The hexane isomerisation unit of Figure 3 has a higher octane for the isomerate product as the ratio of pentanes to hexanes in the feed is comparatively low, unlike a combined C5/C6 isomerisation unit.

Figure 3 illustrates that alternative arrangements are possible and that the basic concepts of the invention disclosed herein can be successfully adapted to other individual situations and requirements.

The benefits of low temperature isomerisation with respect to easing the separation requirements in a recycle operation are well known. Low temperatures are the reason for which once-through paraffin distributions are considerably better than the best perform ance to date, or which may reasonably be expected, from the existing high isomerisation processes which operate at higher temperatures in the reactor. Even with an infinite reactor volume, once can only approach the paraffin distribution which corresponds to the thermodynamic equilibrium at the particular operating reactor temperature.

The following isomerisation performances are representative of the process of the invention after the initial catalyst activity, which is considerably higher, has stabilised.

6 months on stream 12 months on stream Upon regenerated reactor Immediately before Weight basis being placed on line regeneration rearrangement ices 0.839 0.824 nC5 fraction 0.161 0.176 Total C5 paraffins 1.000 1.000 22 DMB fraction 0.508 0.466 23 DMB fraction 0.098 0.098 2 MP fraction 0.231 0.246 3 MP fraction 0.095 0.110 nC6 fraction 0.068 0.080 Total C6 paraffins 1.000 1.000 Although the distribution for isopentane in the total C5 paraffins is an improvement over existing hydroisomerisation processes, the more favourable distribution of the 2,2dimethylbutane in the total C6 paraffins is significantly better, and is the key to economic conversion of the C6 paraffins to a high octane product, as this component determines the separation energy and the required quantity of recycle, the purpose of the isomerisation being, of course, octane improvement. With the utility advantages of the process of the invention, the conversion of C6 paraffins to dimethylbutanes in the isomerate product is relatively inexpensive for the enhanced octane values obtained. The yield of the isomerate product is unaffected, in contrast to catalytic reforming, which suffers a yield decrease for every improvement in octane increase. Furthermore, more severe catalytic reforming adds to the front end octane problem (delta octane number) of European gasolines rather than solving it, as is accomplished by the process of the invention. Also, unlike catalytic reformate, the octane blending numbers of the isomerate product of the process of the invention are 5 to 7 numbers higher, being similar to alkylate with the same characteristic.

The yield of liquid products, i.e., isomerate product and cyclohexane solvent product in the process of the invention is 0.999 weight fraction of the hydrocarbon feed streams entering the isomerisation process, as less than 0.001 weight fraction of the entering feed is converted to propane and lighter, or to deposit on the catalyst. A significant amount of isobutane is formed in the reactor, namely, approximately 0.012 weight fraction of the feed.

Isobutane byproduct is beneficial because the component has 102.0 RONcl, 97.1 MONcl with a .5631 specific gravity which becomes part of the isomerate product.

The volumetric yield of the process of the invention is excellent with 1.012 liquid volume fraction of isomerate product being produced from each 1.000 liquid volume fraction of the C5 paraffins, cyclopentane, and C6 paraffins entering the lead reactor as fresh feed. These higher volumetric yields than existing hydroisomerization processes are economically important as gasoline is generally sold on a volumetric basis. The specific gravity of the isomerate product is 0.636.

WHAT I CLAIM IS.

1. In a process of isomerization, the improved preparation of a pentane/hexane feedstock for isomerisation, which comprises passing a C5/C6 feed to a splitting stage to obtain a denormalised product stream and a normal paraffin feed stream; and passing the latter stream to a low temperature liquid phase isomerisation stage using a Friedel-Crafts halide catalyst suitable for a liquid hydrocarbon phase contacting the promoted with hydrogen halide recirculating stream to the catalytic reactor, the isomerate splitter bottom liquid containing dissolved Friedel-Crafts catalyst in the liquid phase from the isomerization reactor being returned as recycle with the C6 paraffins not converted into desired C6

**WARNING** end of DESC field may overlap start of CLMS **.

Claims (19)

**WARNING** start of CLMS field may overlap end of DESC **. The following isomerisation performances are representative of the process of the invention after the initial catalyst activity, which is considerably higher, has stabilised. 6 months on stream 12 months on stream Upon regenerated reactor Immediately before Weight basis being placed on line regeneration rearrangement ices 0.839 0.824 nC5 fraction 0.161 0.176 Total C5 paraffins 1.000 1.000 22 DMB fraction 0.508 0.466 23 DMB fraction 0.098 0.098 2 MP fraction 0.231 0.246 3 MP fraction 0.095 0.110 nC6 fraction 0.068 0.080 Total C6 paraffins 1.000 1.000 Although the distribution for isopentane in the total C5 paraffins is an improvement over existing hydroisomerisation processes, the more favourable distribution of the 2,2dimethylbutane in the total C6 paraffins is significantly better, and is the key to economic conversion of the C6 paraffins to a high octane product, as this component determines the separation energy and the required quantity of recycle, the purpose of the isomerisation being, of course, octane improvement. With the utility advantages of the process of the invention, the conversion of C6 paraffins to dimethylbutanes in the isomerate product is relatively inexpensive for the enhanced octane values obtained. The yield of the isomerate product is unaffected, in contrast to catalytic reforming, which suffers a yield decrease for every improvement in octane increase. Furthermore, more severe catalytic reforming adds to the front end octane problem (delta octane number) of European gasolines rather than solving it, as is accomplished by the process of the invention. Also, unlike catalytic reformate, the octane blending numbers of the isomerate product of the process of the invention are 5 to 7 numbers higher, being similar to alkylate with the same characteristic. The yield of liquid products, i.e., isomerate product and cyclohexane solvent product in the process of the invention is 0.999 weight fraction of the hydrocarbon feed streams entering the isomerisation process, as less than 0.001 weight fraction of the entering feed is converted to propane and lighter, or to deposit on the catalyst. A significant amount of isobutane is formed in the reactor, namely, approximately 0.012 weight fraction of the feed. Isobutane byproduct is beneficial because the component has 102.0 RONcl, 97.1 MONcl with a .5631 specific gravity which becomes part of the isomerate product. The volumetric yield of the process of the invention is excellent with 1.012 liquid volume fraction of isomerate product being produced from each 1.000 liquid volume fraction of the C5 paraffins, cyclopentane, and C6 paraffins entering the lead reactor as fresh feed. These higher volumetric yields than existing hydroisomerization processes are economically important as gasoline is generally sold on a volumetric basis. The specific gravity of the isomerate product is 0.636. WHAT I CLAIM IS.

1. In a process of isomerization, the improved preparation of a pentane/hexane feedstock for isomerisation, which comprises passing a C5/C6 feed to a splitting stage to obtain a denormalised product stream and a normal paraffin feed stream; and passing the latter stream to a low temperature liquid phase isomerisation stage using a Friedel-Crafts halide catalyst suitable for a liquid hydrocarbon phase contacting the promoted with hydrogen halide recirculating stream to the catalytic reactor, the isomerate splitter bottom liquid containing dissolved Friedel-Crafts catalyst in the liquid phase from the isomerization reactor being returned as recycle with the C6 paraffins not converted into desired C6

isomerate product, together with sufficient naphthene to suppress, in conjunction with hydrogen any disproportionation and hydrocracking reactions.

2. A process as claimed in Claim 1, wherein the Friedel-Crafts halide catalyst is aluminium chloride or zirconium chloride, and the hydrogen halide promoter is hydrogen chloride.

3. A process as claimed in Claim 2, wherein the low temperature isomerization stage has provision for a feed of HCI makeup.

4. A process as claimed in any of Claims 1 to 3, wherein the splitter comprises molecular sieve vessels.

5. A process as claimed in any of Claims 1 to 4, in which the pentane/hexane feedstock is prepared for isomerization by subjecting said feedstock to a pressure swing molecular sieve adsorption treatment, the purge stream of which is derived by benzene hydrogenation, followed, after light end removal, by distillation to remove C7 plus components, whereby all the material within the voids of the molecular sieves, upon desorption, is acceptable as an isomerization feed.

6. A process as claimed in any of Claims 1 to 5 in which a pentane/hexane feedstock is prepared for isomerization by means of a pressure swing molecular sieve adsorption treatment wherein the desorption of the normal paraffins is carried out at three pressure levels.

7. A process as claimed in any of Claims 1 to 6, in which pentane/hexane isomerization is effected by means of hydrogen halide promoted Friedel-Crafts catalyst operating at ambient temperatures in the liquid phase except for a stream of circulating hydrogen in the gas phase, which is recovered together with separated volatile hydrogen halide and returned as recycle to the reactor, the liquid product being freed of hydrogen halide at a low pressure which permits a moderate temperature in the reboiler, and separated into the desired more volatile isomerate products by distillation.

8. A process as claimed in claim 7, in which a cyclohexane stripper, replacing part of the isomerate splitter reboiling, is fed with part of the isomerate splitter bottoms, freed by adsorption of dissolved Friedel-Crafts catalyst, part of the fresh feed being used to regenerate the adsorbent so as to return any active Freidel-Crafts catalyst to the isomerization reactor.

9. A process as claimed in claim 8, in which two absorbent beds are employed, alternating in service between one depositing Friedel-Crafts catalyst on the absorbent, and the other removing the active Friedel-Craft catalyst deposited by means of part of the fresh feed.

10. A process as claimed in claim 8 or 9, in which the C6 naphthene drag stream from the cyclohexane stream is a 95 percent cyclohexane solvent.

11. A process as claimed in claim 7 wherein the hydrogen halide-free liquid product from the isomerization is separated into an isopentane overhead product; a normal pentane heartcut part of which is used as a purge stream and the remainder of which is recycled to isomerization for conversion into isopentane; and a neohexane product sidestream, and the remaining C6 paraffins in the bottom stream are recycled to extinction to return any dissolved Friedel-Craft catalyst withdrawn from the isomerization reactor back to the isomerization reactor system as well as to return all C6 naphthenes used within the reactor as a suppressant for further use as a suppressant, the onginal inventory required for the circulating C6 naphthenes being supplied from other sources at startup.

12. A process as claimed in claim 7, wherein pentanes with a negligible cyclopentane plus content are fed as fresh feed to a common isopentane/normal pentane splitter, and separated this into an isopentane isomerate product and a normal pentane bottoms, containing dissolved Friedel-Crafts catalyst, which is returned direct as recycle without drying to the isomerization reactor for isomerization of the normal pentane into isopentane, the pentane feed being introduced into the isopentane-normal pentane splitter below the isomerization liquid reactor product feed, so that the isopentane component is preferentially separated overhead to form part of the isopentane product, whereas the fresh normal pentane feed component provides sufficient bottom product for any Friedel-Crafts catalyst dissolved in the isomerization reactor product to be always below the solubility limit of the liquid phase present at any point.

13. A process as claimed in claim 12, wherein part of the bottoms from the isopentane-normal pentane splitter is withdrawn as a purge stream for use in the adsorption treatment by passing over an adsorbent on which dissolved Friedel-Craft catalyst has been deposited, the catalyst being returned to the system by using part of the pentanes feed for dissolving active catalyst deposited on the absorbent, before the latter is returning to its absorbent mode; the purge stream serving as a drag stream to limit the cyclopentane plus content in the recycle.

14. A process as claimed in claim 13, in which part of the normal pentane is stripped in a cyclopentane plus product stripper and returned to the base of the isopentane-normal pentane splitter, to allow further concentration of cyclopentane plus component to be tolerated in the pentane feedstock without excessive concentrations in the isomerization reactor feed.

15. A process as claimed in claim 11, wherein the facilities for the regeneration of the Friedel-Crafts catalyst in the reactor are shared between a plurality of reactors.

16. A process of isomerization as claimed in claim 7 or claim 11, wherein the Friedel-Crafts catalyst is aluminium chloride supported on activated alumina particles less than one millimeter in diameter, and the liquid phase is passed upflow in the isomerization reactor with circulating hydrogen chloride as the promotor.

17. A process for isomerization as claimed in claim 7 or claim 11, wherein the Friedel-Crafts catalyst is zirconium chloride with catalyst particles less than one millimeter in diameter, and the liquid phase is passed upflow in the isomerization reactor with circulating hydrogen chloride as the promoter.

18. A process for isomerization as claimed in claim 7 or 11 wherein a plurality of reactors are used with valved connections such that they may serve in sequence in the various onstream reactor positions and be periodically regenerated, using hydrogen at higher pressure and higher temperatures, and activated with circulating hydrogen halide, before isomerization reactor service.

19. A process according to claim 1 substantially as herein described with reference to the accompanying drawings.