GB1604777A - Catalytic reformer process - Google Patents

Description

(54) CATALYTIC REFORMER PROCESS (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 the catalytic reforming in a thermally efficient manner of hydrocarbon fractions, to improve hydrogen yields and pentane plus values in the gasoline product of a given octane quality obtained from a given feedstock. The aromatics yield significantly increases for a given severity of treatment, primarily at the expense of decreased hydrocracked paraffins in the product, so the hydrogen byproduct purity is significantly increased by the process of the invention.

The invention accomplishes this result by providing an alkylcyclohexane converter utilizing the sensible heat in the reactor effluent from the terminal reactor of a series of reactors to provide by simultaneous heat-exchange, the endothermic heat of reaction required for the dehydrogenation of the original alkylcyclohexanes in the feed to aromatics. As both the feed to the converter and the reactor effluent contain hydrogen as the principal component, the heat transfer characteristics are excellent and well suited so that heat exchange can be enacted economically with a limited surface. Because the alkylcyclohexanes dehydrogenate rapidly, only a limited catalyst volume is required and suited to the conventional heat-exchanger tube diameters available.

Catalytic reforming reactors to-day consist of fixed bed adiabatic reactors to perform all the catalytic reactions required in catalytic reforming with a platinumcontaining catalyst. These adiabatic reactors, while very suitable for carrying out the slower paraffin dehydro-cyclization reactions in forming aromatics, suffer a serious disadvantage in comparison with the process of the invention in dehydrogenating the naphthene components entering in the feed because, by their nature, they must be heated to a temperature level higher than the bed outlet temperature, to provide the endothermic heat of reaction.Higher temperatures result in decreased molar selectivity for the naphthene components to aromatics The lead adiabatic reactor usually has a sharply falling temperature drop of at least 1200F for the lower pressures made practicable by the development of promoted platinum catalytic reforming catalyst, with the addition of metals such as rhenium.

These promoted or multi-promoted platinum catalysts enhance the thermal stability and reduce the tendency of the previous best platinum on alumina catalysts to agglomerate and to form larger platinum sizes by a factor of 5 or more.

The larger the platinum size, the less effective is the catalyst for naphthene dehydrogenation or overall paraffin dehydrocyclization reactions. On the other hand, platinum size does not affect the hydrocracking activity so that platinum catalyst prior to the development of promoted platinum catalysts (first commercial tests commenced in 1968) had, with semi - regenerative adiabatic reactor arrangements, a sharp decrease in pentane plus yields and hydrogen purity as the temperatures increased to compensate for declining activity.

Platinum sintering or agglomeration of the platinum particles is a second order reaction strongly dependent upon the temperatures prevailing in the bed and, being a time-temperature function, semi-regenerative adiabatic reactor arrangements are most affected. With the promoted catalyst development, instead of a sharp yield decline when temperatures above 9700F were approached, the yield and hydrogen purity decline were gradual, even when temperatures were increased to 1,020 F.

These temperatures were permitted by another characteristic of the promoted platinum catalysts. Thus, the previous best platinum only catalyst having carbonaceous deposits up to 10 weight percent, based on the original catalyst weight, in the terminal reactor, corresponding to the highest carbonaceous deposit on the catalyst, as the catalyst is subject to the lowest hydrogen partial pressure, the highest average bed temperature, the highest concentration of aromatic and the greatest hydrocracking activity, all of which adversely affect the carbonaceous deposition.In contrast, the promoted catalysts without any sharp yield declines can routinely have carbonaceous deposits of 15 to 20 weight percent, and be regenerated (oxidation of the carbonaceous deposits, chlorination or haliding to restore original halogen content, rejuvenation or redispersion of any agglomerated platinum into smaller sizes, and reduction by hydrogen with limited C2 plus hydrocarbon content) with full restoration of their fresh catalyst activity.

This meant that lower operating pressures became practicable for even the semiregenerative adiabatic reactor arrangement. The Chevron Research Company, the original developer of promoted platinum catalyst technology, has recently reported, for example, that converted semi-regenerative units, with the reactors containing the most recentlydeveloed promoted catalyst have operated with molar hydrogen to feed ratios for the hydrogen recycle of from 2.5:1 to 3:1, and at 100 to 200 psig at the last reactor outlet, with cycles in excess of six months between regenerations.

All the major licensors of catalytic reforming technology have now either adopted or separately developed a promoted or mutli-promoted platinum catalyst for use in their licensed process, so that the adoption of lower pressures has been made practicable whatever the reactor arrangement.

The benefits of lower pressure had been previously demonstrated by the cyclic or swing reactor arrangement using previous platinum catalyst technology. In the cyclic or swing reactor arrangement, a reactor is regenerated onstream while processing with the other reactors. The regeneration is carried out by isolating the reactor previously in the terminal reactor position with separate regeneration equipment. After regeneration, including chlorination, rejuvenation (then more important) and hydrogen reduction, the regenerated reactor is placed in parallel with the lead reactor position, the parallel previous lead reactor is placed in parallel with the intermediate reactor position, the parallel previous intermediate reactor is placed in parallel with the terminal reactor position, and the previous terminal reactor is regenerated.In this cyclic reactor arrangement, all the reactors, including the swing reactor being regenerated, are of similar volume, and have double block valves permitting them to be connected to the heaters for the various reactor positions. This additional equipment enables such a reactor arrangement to process continuously while regeneration and rejuvenation are carried out onstream, but the reactors in the major commercial processes to-day are adiabatic and have essentially the same reactor characteristics. The important point is that, while the regeneration method may differ, all existing processing arrangements for catalytic reforming have the same inherent kinetic disadvantage in the case of the conversion of the original naphthene components in the feed by means of adiabatic reactors.

The other major commercial catalytic reforming process is the "continuous" reactor arrangement. This variation of the adiabatic reactor arrangement differs in the regeneration only, which is also carried out onstream while processing feedstock in the normal manner. In this arrangement, a small stream of catalyst is transported using hydrogen for pneumatic transport from the lift pot at the base of the reactor to a surge hopper above the regeneration vessel. Small batches of catalyst are periodically regenerated, including chlorination, rejuvenation, if necessary, and reduction by hydrogen with the regenerated catalyst proceeding to a surge hopper. Generally, the reactors are stacked with the lead reactor on the top and the terminal reactor on the bottom, so that with the surge hopper immediately above the lead reactor, regenerated catalyst enters the lead reactor by gravity.

Transfer between the reactors is also by gravity (the catalyst settles in the reactors in plug flow because of the small stream of catalyst being withdrawn and added). In this manner, the attrition of catalyst is limited as only the catalyst to be regenerated must be pneumatically transported. The continuous reactor arrangement, unlike the cyclic reactor arrangement, can utilize the same reactor catalyst distribution now generally employed for the semi-regenerative reactor arrangement, which consists of the smallest reactor volume vessel being in the lead reactor position and the largest reactor volume vessel being in the terminal reactor position.This reactor catalyst distribution utilizes catalyst more efficiently, as the faster reactions at the lead reactor require only a limited reactor volume before the fall in temperature sufficiently retards further desired reactions thermodynamically. In contrast, the terminal reactor position is concerned with converting the least reactive components, generally C6 and C, paraffins which have considerably slower overall paraffin dehydrocyclization rates than the higher carbon number paraffins.

Slower reaction rates require greater catalyst volume so that increased hydrocracking relative to the formation of aromatics is inevitable, but the use of higher space velocity for the faster reactions decreases the time that the entire feed can undergo hydrocracking reactions. Hydrocracking being a consumer of hydrogen is exothermic so that temperature drops in the terminal reactor are limited as these exothermic reactions act to offset the endothermic requirements resulting from the dehydrocyclization of paraffins into aromatics.

According to the present invention I provide a catalytic reformed process for the production of hydrogen and a product of enhanced aromatic content from a crude naphtha feedstock which process comprises a plurality of reactor stages in which the feedstock is subjected to naphthene conversion, wherein the heat in the effluent from the terminal reactor stage is utilised in an alkylcyclohexane converter, by means of simultaneous heat-exchange, to provide the endothermic heat of reaction required in the alkylcyclohexane converter for the dehydrogenation to aromatics of the original alkylcyclohexanes in the feed, before the said feed enters the naphthene conversion reactor stages.

The process of the invention can be used advantageously to modify any of the major commercial reactor arrangements (all of which possess an inherent kinetic deficiency) by taking full advantage of the lower operating pressures to dehydrogenate the original naphthenes in the feed more selectively into aromatics.

Reference is now made to the accompanying drawings in which: Figure 1 is a flow sheet of one embodiment of the process of the invention using a semi-regenerative reactor arrangement; Figure 2 is a flow sheet of the process of the invention applied to a cyclic or swing reactor arrangement; Figure 3 is a flow sheet showing the process of the invention applied to a continuous reactor arrangement; Figure 4 and 5 are graphs of equilibria ratios of mols of aromatics to mols of alkylcyclohexane; and Figure 6 is a flow sheet of a typical heat recovery arrangement using the process of the invention.

The apparatus shown in Figure 1 comprises an alkylcyclohexane converter 1 and an alkylcyclopentane converter 2. These are followed by a series of three reactors 3-5, namely, a lead adiabatic reactor 3, an intermediate adiabatic reactor 4 and a terminal adiabatic reactor 5. Each of these reactors is provided with a fired heater, namely, the lead heater 6, the intermediate heater 7 and the terminal heater 8. It will be appreciated that there may be more than one intermediate reactor and heater.

A feed/reactor effluent heat-exchanger is provided at 9, a separator at 11 and a hydrocarbon collector at 12.

A reactor effluent air cooler 13 is provided in the flow between the terminal adiabatic reactor 5 and the separator 11. A recontact pump 14 is provided in the flow between the separator 11 and the collector 12, and a compressor 15 is provided in the hydrogen recycle line between the latter units. The pump 14 and compressor 15 are followed in the recycle by a recontact air cooler 16 and a recontact trim cooler 17.

Hydrocarbon feed enters the system at 10 and passes through the heatexchanger 9 into the converters 1 and 2, respectively, and through the reactor arrangement 3-8. Terminal reactor effluent 21 is returned to the alkycyclohexane converter, and via the exchanger 9 and air cooler 13 to the separator 11. Hydrogen is recycled through the line 18 via the compressor 15, and hydrocarbon liquid is recycled through the line 19 via the pump 14. Hydrocarbon liquid is withdrawn at 20 from the collector 12 either to separation or to recontact with hydrogen by product. Hydrogen by-product is taken off at 22 and a hydrogen recycle to the main feed 10 is provided at 23.

Figure 2 shows the application of the process of the invention to a cyclic or swing reactor arrangement. The fragment shown in Figure 2 corresponds to the right-hand portion of the equipment shown in Figure 1 but, in this case, there are four reactors, 31 to 34, inclusive. In operation, three of these reactors are onstream whilst one is regenerating. According to one modification, the arrangement may comprise a further reactor with its fired reheater so that four reactors may be onstream whilst one is regenerating.

The reactors are provided, as in the case of the arrangement of Figure 1, with fired heaters, three such heaters, 35 to 37, inclusive, being shown. The heater 35 corresponds to the lead reactor position, the heater 36 to the intermediate reactor position and the heater 37 to the terminal reactor position. The heaters 35 to 37 are combined within a furnace 38.

The flow into the reactor arrangement enters at 24, e.g, leaving the alkylcyclopentane converter (as in Fig. 1). Terminal reactor effluent is recycled at 21 also as in Figure 1. An entering flow 41 from regeneration equipment (not shown) is shown in dotted lines and the return to the regeneration equipment is indicated, also in dotted lines, at 42. The regeneration equipment has its own heater and is assumed to be in the recirculation condition. Flue gases leave the furnace 38 to 43.

The continuous reactor arrangement shown in Figure 3 again corresponds to the right-hand portion of the apparatus of Fig. 1. In this case, the three reactors 52-54, inclusive are superposed upon one another in a single column 51. Each of the three reactors, namely the lead reactor 52, the intermediate reactor 53 and the terminal reactor 54 is provided with its fired heater.

A regeneration vessel is provided at 55 and surge hoppers 56 and 57 are included in a regeneration flow 58 which is shown in dotted lines. The regeneration vessel 55 is provided with a heater, recirculation, conditioning equipment and the necessary connections (not shown). Pneumatic transport for the stream of small catalyst particles is provided in the regeneration line 58, and a lift hopper 59 is positioned at the base of the column 51. Hydrogen enters at 61.

The feed from the alkyl cyclopentane converter 2 (Fig. 1), for example, again enters at 24, and the terminal reactor effluent, which heats the alkyl cyclohexane converter 1 (Fig. 1), leaves the base of the column 51 at 21.

The graph of Figure 4 represents the equilibria ratios of mols of aromatics to mols of alkyl-cyclohexane. These are plotted as ordinate values on a logarithmic scale in relation to abscissa values which are reciprocals of absolute temperature ("K). It should be noted that higher hydrogen partial pressures shift the individual component lines of the graph to the right.

Figure 5 is a graph of the equilibria ratios of mols of aromatics to mols of alkyl cyclohexane, and is again plotted on a logarithmic scale but in relation to hydrogen partial pressures. The latter are also on a logarithmic scale and are expressed as absolute pressures. It should be noted that higher temperatures shift the individual component lines upwards.

Figure 6 is a graph of a typical heat recovery arrangement which is usable with the process of the invention with 1,800 psig, 9500F steam turbine conditions, it appears that the catalytic reformer can supply the hydrogen recycle compressor driver requirements with significantly less fuel than is required for firing existing heaters.

A fired heater 71 comprises a lead reactor radiant section 72, an intermediate reactor radiant section 73 and a terminal reactor radiant section 74. A highpressure superheater radiant section is shown at 75.

A high pressure steam generator is shown at 76, a high pressure boiler feed heater at 77 and an air heater at 78. The flow 79 from the steam drum of the high pressure steam generator passes at 81 through the radiant section 75, and flows thereafter as high-pressure superheated steam 82 to a steam turbine driver 83 for the hydrogen recycle. An air-cooled vacuum condenser is provided at 84 and a vacuum compressor for air removal at 85. A boiler feed pump 86 recirculates at 87 to the boiler feed heater 77. Extraction steam leaves the steam cycle at 88.

A forced draft fan for the air heater 78 is shown at 89 and the flow to the radiant fired sections is indicated at 91. The flow of flue gases is shown at 92, and the alkyl cyclopentane converter is schematically indicated at 93.

As stated above, Figure 1 shows the process of the invention applied to a semiregenerative reactor arrangement, Figure 2 shows the process of the invention applied to a cyclic or swing reactor arrangement, and Figure 3 shows the process of the invention applied to a continuous reactor arrangement. however, the process of the invention may also be applied advantageously to a hybrid of these arrangements.

such as using semi-regeneration for the lead and intermediate reactor(s) and having two terminal reactors in parallel with double block valves and separate equipment for regeneration onstream similar to the cyclic regenerative arrangement. This enables long run lengths equivalent to the maintenance shutdown requirements as the lead and intermediate reactors have a considerably lower carbonaceous deposition pattern than the terminal reactor position. The lower carbonaceous rate for the lead and intermediate reactor positions can be ensured by placing maximum inlet temperature restrictions on these reactor positions. The terminal reactor position can be periodically regenerated onstream with the processing rate being lowered during the regeneration, and the full feed rate resumed upon placing the regenerated catalyst in the terminal reactor position.

The process of the invention, therefore, is a general improvement and is not restricted to a specific reactor arrangement or a specific catalyst.

Any adiabatic reactor arrangement, unmodified by the process of the invention, subjects the entire saturates in the feed to thermal cracking temperatures in the fired heater to the lead reactor, as the inside surface temperatures of the radiant heated tubes, which generally provide final heating to reactor inlet temperatures, are approximately 500 F. hotter than the bulk fluid temperature. Thermal cracking significantly increases the hydrogen consumption per mol of feed component over hydrocracking. Methane, ethane and propane are the principal by-products regardless of the saturate component cracked as any ethylene or propylene intermediate is hydrogenated in the reactor.The entering paraffins, except for any components thermally cracked, are at feed concentration, and are thus subject to significant hydrocracking losses in the lead reactor because of the high entering temperature.

Although a higher temperature increases all the reaction rates, the paraffins of given carbon number cannot take any advantage of this until the original alkylcyclohexanes of the same carbon number have been sufficiently lowered in concentration for a favourable thermodynamic equilibrium to exist as the paraffin dehydrocyclization into alkylcyclohexanes is a reversible reaction (alkylcyclohexanes can also decyclize and hydrogenate to form paraffins, and do so that the higher inlet temperatures when at maximum temperatures). The dehydrogenation of alkylcyclohexanes into the alkylcyclohexane intermediates, and further dehydrogenation into aromatics (alkylbenzenes) is also a reversible reaction which is subject to thermodynamic equilibria. The alkylcyclohexene reaction is a very fast reaction.The alkylcyclohexane dehydrogenation into an alkylcyclohexene intermediate, while fast when compared with the other reactions which occur, is slow as compared with the alkylcyclohexene dehydrogenation into aromatics. The alkylcyclohexane dehydrogenation into the alkylcyclohexene intermediate, therefore, controls the lowering of the original alkylcyclohexane concentration, and time is required. In the meanwhile, while the high original concentrations in the feed are being sufficiently reduced by conversion to aromatics, alkylcyclohexanes are being lost to the competitive reactions provided thermodynamic equilibria permit the formation of paraffins and alkylcyclopentanes, because these competitive reaction rates have been speeded up by the high temperatures prevailing at the inlet of the reactor.Thus, these original feed components can only proceed directionally towards hydrocracking as all aromatics formed must proceed through the alkylcyclohexene dehydrogenation intermediates. The alkylcyclopentanes, being initially prevented by the thermodynamic equilibria from ring isomerization into alkylcyclohexanes can only decyclize and hydrogenate to form paraffins which reinforce the hydrocracking reaction losses. This reaction is a reversible reaction in that paraffins may also dehydrocyclize into alkylcyclopentanes, but this competitive pathway is not available until the original concentration of the original alkylcyclopentanes has also been lowered to the point where a favourable thermodynamic equilibrium exists for the paraffins.

Higher temperatures unfavourably affect the molar ratios at thermodynamic equilibria which are important for the selective conversion of the original naphthenes in the feed into aromatics. The ring isomerization of the alkylcyclopentanes into alkylcyclohexanes is a good example of this effect, as this equilibrium, unlike the dehydrogenation or dehydrocyclization reactions (favoured by a lower hydrogen partial pressure because hydrogen is produced), or the reverse reactions (favoured by a higher hydrogen partial pressure because hydrogen is consumed) is little affected by hydrogen partial pressure.For example, mols of alkylcyclopentanes per mol of alkylcyclohexane at thermodynamic equilibrium are: 800"F 9000 F C6 naphthenes 9.0 11.5 C7 naphthenes 2.7 3.2 C, naphthenes 0.265 0.43 C9 naphthenes 0.087 0.175 The inlet temperatures to the lead reactor is generally never less than 900c F. For the increasing octane requirements of the future, with increasing lead additive legal restrictions being enacted, inlet temperatures for an equal reactor inlet pattern of substantially 930"F are likely for the typical hydrotreated straight-run feedstocks similar in quality to naphtha derived from Kuwait crude, even for the start of run conditions with a semi-regenerative reactor arrangement.The other reactor arrangements generally practise higher reactor inlet temperatures than 930"F as the catalyst inventory is generally smaller due to the ability to regenerate onstream in order to compensate for catalyst deactivation. Furthermore, continuous processing reactor arrangements do not normally regenerate the catalyst in the reactors at shutdown unless entering a reactor is required during the shutdown, whereas a semi-regenerative reactor arrangement starts up with freshly regenerated catalyst in all reactors as the hydrogen circulating equipment is employed for regeneration. Each semi-regeneration run, thus, commences with fresh catalyst, as in the initial start-up with new catalyst.In the case of heaters designed for an equal reactor inlet pattern, if octane requirements increase, one commences with a higher start of run inlet temperature to the adiabatic reactors in order to meet the quality demanded.

The process of the invention, in contrast, typically commences with an inlet temperature of 780"F into the alkylcyclohexane converter. This lower entering temperature together with an outlet temperature of approximately 8000F for the alkylcyclohexane converter significantly lowers the hydrocracking activity in addition to lowering the competitive reactions which cause loss of the original alkylcyclopentane and alkylcyclohexane components in the feed. One alkylcyclohexane is converted to a component having a slower reaction rate for the overall conversion into aromatics, additional hydrocracking is inevitable even through part of this component is later converted into aromatics.

As may be seen from the latter favourable thermodynamic equilibrium at 800"F as compared with 9000F, the ring isomerization of alkylcyclopentanes into alkylcyclohexanes can commence sooner if the alkylcyclohexanes are dehydrogranted at a lower temperature. This decreases the conversion of alkylcyclopentanes into paraffins, particularly for the higher carbon numbers which also have the highest reactivity.

The rising temperature gradient of the naphthene reactors of the process of the invention contrasts with the sharply falling temperature in the lead reactor of conventional catalytic reformers. Although both may have approximately the same outlet temperature of 800 F., any hydrocracking or conversion of the original alkylcyclohexanes to alkylcyclopentanes and paraffins from higher temperatures than required represents a permanent decrease in molar selectivity in converting the original alkylcyclohexanes into aromatics. Reactivity of the various components favours the rising temperature pattern as reactivity of any of the reversible reactions sharply increases with carbon numbers. The order of the reaction rates at 8000F is indicated as follows: Component Reactivity C6 Base rate C7 4 times base rate C8 12 times base rate C9 30 times base rate C,O 60 times base rate C1, 90 times base rate This increase in reactivity with increase in carbon number is well suited to a rising temperature pattern because, once a saturate is converted to an aromatic, the component can be considered thermally stable. In contrast, the thermal stability of the saturates decreases with increases in carbon number. The foregoing reactivity orders change with increase in temperature as chemisorption effects become less important.The hydrocracking reactivity of paraffins also increases with carbon number, but the orders are different from the reversible reactions, being approximately as follows: Component Reactivity C6 0.56 base rate C7 0.76 base rate C8 Base rate C9 1.28 base rate C10 1.60 base rate Because the higher carbon numbers have the most favourable thermodynamic equilibria (see Figure 4), for a given hydrogen partial pressure (140 psi), a rising temperature gradient reactor can commence at a lower temperature and be more selective in converting the components which have the least thermal stability as lower temperatures permit the porous catalyst preferentially to adsorb the highest carbon numbers, whereas higher temperatures diminish the chemisorption effects.

Because reactivities refer only to a given carbon number and each carbon number has its own thermodynamic equilibrium restraints for a given temperature and hydrogen partial pressure, the general practice of a single line representing thermodynamic equilibrium behaviour prevailing in catalytic reforming reactors is misleading. Figure 5 shows the mols ot aromatics per mol of alkylcyclohexane at thermodynamic equilibrium as in Figure 4, but shows the effect of hydrogen partial pressure upon the individual carbon numbers for a given temperature of 800"F (typical outlet temperature for the alkylcyclohexane converter). As may be seen, a lower hydrogen partial pressure favours the conversion of any alkylcyclohexane into aromatics.At 140 psia hydrogen partial pressure, even the C6 component has a favourable equilibrium at 8000F, which allows conversion of most of the original cyclohexane into benzene.

Although one generally considers the alkylcyclopentane isomerization reaction into alkylcyclohexane as slow when this is compared with the dehydrogenation rate of alkylcyclohexanes into aromatics, one must consider the carbon numbers because of increased reactivity with carbon number.

The highest carbon number ring isomerization is, therefore, comparable in rate to the lowest carbon number alkylcyclohexane dehydrogenation rate. Thus, the more preferential enhancement of the depletion of the original alkylcyclohexanes in a rising reactor means that considerable C9 and higher alkylcyclopentane conversion into aromatics should occur concurrently with the cyclohexane conversion into benzene, which may be expected to occur near the reactor outlet. Reactions for the lower carbon number alkylcyclohexanes occur later as the platinum dehydrogenation sites in the early part are preferentially occupied with the dehydrogenation of the higher carbon number alkylcyclohexanes into aromatics until their original concentration drops sufficiently to permit more competitive access.

The lower reaction rates corresponding to the lower temperatures in the alkylcyclohexane converter of the invention should be offset by employing a catalyst with a higher platinum content as the alkylcyclohexane dehydrogenation reaction into aromatics occurs only at these sites, and reaction rates at these sites have been shown to be transport-controlled. Doubling the platinum concentration of the catalyst essentially doubles the reaction rate and halves the conversion loss of cyclohexane into methyl cyclopentane while depleting the original cyclohexane.

With hourly space velocity of approximately 20 for the alkylcyclohexane converter, the expenditure for additional platinum is justified, particularly as the expenditure not only improves the kinetics of the alkylcyclohexane converter but, also lowers the temperature of the converter at the inlet below the stream entering from the feed outlet of the feed reactor effluent heat-exchanger. This is because available platinum sites limit the endothermic heat removal reactions which are afforded by the high reactivity of higher carbon number alkylcyclohexanes. The alkylcyclohexane converter is limited by the heat transfer requirement and not by the availability of alkylcyclohexanes in the feed. Accordingly, by lowering the temperature rise additional platinum sites direct more of the sensible heat available in the reactor effluent to providing endothermic reactions.The prime purpose of the alkycycldhexane converter is to reduce the original alkylcyclohexane content as soon as possible because higher temperatures can only reduce the molar selectivity for the conversion of naphthalene components into aromatics.

The alkylcyclohexane content of hydrotreated straight run feedstocks varies with the feedstock, ranging typically from an 0.34 to 0.69 fraction of the C6 naphthenes, an 0.38 to 0.79 fraction of the C7 naphthenes and an 0.55 to 0.84 fraction of the C6 naphthenes. It may be noted that the fraction of alkylcyclohexanes increases as the carbon number increases, so that the yield enhancement with the process of the invention may be expected to increase with an increase in cutpoint. As naphthene contents also generally increase with increase in cutpoint, one may expect that what is termed the alkylcyclopentane converter in this invention, using flue gas from the fired radiant heater sections (Fig. 6), applies normally to a limited naphthene content feedstock such as illustrated in the example discussed below.The first portion of this converter continues the dehydrogenation of the alkylcyclohexanes for higher naphthene content feedstocks. The alkylcyclohexane converter has limited reactor volume, and, with limited sensible heat available in the reactor effluent from the terminal reactor, can dehydrogenate only a limited quantity of aikylcyclohexanes into aromatics.

An increased amount of naphthene dehydrogenation may be expected with a higher cutpoint feedstock, as the higher carbon numbers have a higher reaction rate, but, as temperature levels also fall under such conditions, this acts to retard the reaction rate, so that the alkylcyclohexane converter outlet temperature is not more than 20"F below the design value. On the other hand, an increase in reactor effluent temperature, such as occurs in semi-regenerative reactor systems even at the end of the run, will not increase the alkylcyclohexane converter temperature by more than 20"F, because of the compensatory increase in reaction rates with increase in temperature.Thus, the alkylcyclohexane converter acts to compensate either for variations in the feedstock or for variations in the reactor effluent temperature, so that cooling duties for the reactor effluent in a semi-regenerative reactor arrangement should remain relatively constant rather than increasing as the run progresses. The fired heater duties for maintaining the necessary reactor inlet temperatures to the adiabatic reactor should also be less affected by variations in reformer feed quality and be subject to less variation from start of run to end of run for a semi-regenerative reactor arrangement.

The process of the invention provides a solution to the need for more efficient use of the increased molar sensible heat available in the reactor effluent as compared with the reactor feed. A decrease in operating pressure significantly decreases the hydrocracking by-products relative to the aromatics formed. The effects are demonstrated for the C7 saturate components entering in the feed, with methyl cyclohexane to toluene typifying the naphthene dehydrogenation reactions, and normal heptane to toluene typifying the paraffin dehydrocyclization reactions.

The reactions are: Naphthene dehydrogenation C7H,4eC,H8 (toluene)+3H2 Paraffin dehydrocyclization C,H18 < C7H8 (toluene)+4H2 At 9000F., the total molar sensible heat for the products per mol of naphthene feed converted at 87.6 Btu per "F as compared with 77.3 Btu per "F for the feed component. For each mol of paraffin converted, the equivalent figures are 95.0 Btu per OF for the product as compared with 86.0 But per OF for the feed component.

Hydrocracking reactions which occur result in less sensible heat for the products as compared with the reactants. The overall effect of increased molar selectivity in the conversion to aromatics at the expense of decreased hydrocracking is even more dramatic as the hydrocracking reaction may be typified by the following overall reactions, even if the naphthene proceeds through a paraffin intermediate and the paraffin proceeds through an olefin intermediate: : Naphthene hydrocracking C7H14+2H2.02 CH4+0.015 C2H6+ 0.965 C3H8+0.965 C4Hsso+OOl5 CsHl2+ 0.020 C6H,4 Paraffin hydrocracking C7H,8+H2~0.02 CH4+0.015 C2H6+ 0.965 C3HS+0.965 C4H10+0.015 Cosh12+ 0.020 C8H,4 For the molar sensible heat at 900"F, the products total 82.0 Btu per OF (both cases) as compared with 91.5 Btu per OF for the naphthene reactants and 93.1 Btu per "F for the paraffin reactants.Thus, overall, an improved molar selectivity in the conversion of the feed component at the expense of hydrocracking results in 5 mols of additional hydrogen per mol of aromatics formed, whether from a naphthene feed component or a paraffin feed component. The increased molar selectivity effects upon molar sensible heat are equally dramatic as the substitution of 5 mols of H2 plus toluene for the previous hydrocracked byproducts at 900OF corresponds to 102.1 Btu per "F for each additional mol of toluene produced at the expense of hydrocracking products eliminated (82.0 Btu per "F), so the molar sensible heat effect is equivalent to 1.245 that the decreased hydrocracking is converted to additional hydrogen and aromatics.Thus, the molar sensible heat for the reactor effluent compared with the reactor feed significantly increases with the lower pressures, and will increase further with the use of the process of the invention. If the approach temperature at reactor effluent entrance to the reactor feed out remains the same, the duty for cooling the reactor effluent increases with lower pressures as the temperature into the air cooler must increase, based upon the feed reactor effluent heat-exchanger providing the only heat-exchange between the terminal reactor outlet and the air cooler. Part of the excess heat can be utilized for separation but this low temperature level duty is thermodynamically better provided by extraction steam from a steam turbine.

The increased molar selectivity achieved by the process of the invention, together with the low hydrogen recycles rendered practicable without pioneering either as to mols of hydrogen per mol of hydrocarbon feed into the lead adiabatic reactor or deviating substantially from existing hydrogen partial pressures in the terminal adiabatic reactor, further increase the molar sensible heat in the reactor effluent as compared with the feed. However, the alkylcyclohexane reactor of the invention lowers the temperature into the feed reactor effluent exchanger to approximately 830OF as compared with value about 930"F which is found conventionally.

This means that, if the feed process of the invention is applied as a retrofit to existing catalytic reformers, and the feed reactor effluent exchangers are designed for low fuel costs, such a unit using the alkylcyclohexane converter of the invention will have a reactor effluent leaving temperature to the air cooler which is lower than that achieved in the recent catalytic reforming plants with true countercurrent heat exchangers, and surfaces provided for high energy costs.This should occur because of the decreased duty (corresponding to both the lower hydrogen recycle and lower reactor effluent entering temperature with the process of the invention) which would result in a close approach at reactor effluent inlet, together with the decreased cooling range for the reactor effluent, which limits the effect of the higher molar sensible heat for the reactor effluent as compared with the feed (hydrocarbon feed plus hydrogen recycle).

Based upon the terminal reactor effluent at 9300F entering the alkylcyclohexane converter, an equivalent adiabatic reactor drop of substantially 85"F may be expected. This compares with the 120 to 1300F lead adiabatic reactor temperature drops typically occurring when reactor effluent separator pressures are below 200 psig. Most of the adiabatic temperature drop corresponds to the highly endothermic alkylcyclohexane dehydrogenation into aromatics and hydrogen, with almost insignificant offsetting exothermic hydrocracking reactions.

Alkylcyclohexanes have a significantly higher endothermic heat requirement per mol of aromatics produced than any other class of saturates, as indicated below, with the heats of reaction expressed in Btu per mol of feed component converted into a mol of aromatic: To Aromatics+3H2 To Aromatics+4H2 Alkylcyclo Alkylcyclo Normal Component hexanes pentanes paraffins Isoparaffins Css 94140 87300 75060 70920 C7 92700 84600 77780 73620 C9 92160 78840 79380 75600 C9 92180 75960 81360 77760 C10 91440 73800 82800 79380 Although paraffins appear to have the same endothermic requirements as the alkylcyclopentanes, in practice, significantly more exothermic reactions occur with the paraffins than with the naphthene because of the associated hydrocracking.

with the production of aromatics. A pronounced difference between the alkylcyclopentanes and the paraffins in effective heats of reaction occurs, as shown by the following overall conversions which are typical for adiabatic reactors when operating under the same conditions as in the example discussed below:: Fraction of Fraction of feed hydro- feed to Feed component hydrocracked Component cracking aromatics to that converted to aromatics C, CH 0.028 0.960 0.029 Ce CP 0.038 0.916 0.041 nC6P 0.250 0.166 1.51 iCP 0.258 0.150 1.72 C7 CH 0.036 0.960 0.038 C7 CP 0.046 0.945 0.049 nC7P 0.320 0.580 0.552 iC7 P 0.330 0.525 0.628 C6 CH 0.039 0.960 0.041 C8 CP 0.050 0.944 0.053 nC8 P 0.245 0.745 0.328 iC9 P 0.254 0.706 0.360 C9 CH 0.042 0.958 0.044 C9 CP 0.053 0.947 0.056 nC9 P 0.208 0.791 0.263 iC9 P 0.216 0.779 0.277 C0 CH 0.043 0.956 0.045 C10 CP 0.054 0.946 0.057 nC10P 0.186 0.814 0.229 iC,o P 0.193 0.807 0.236 Although the thermal stability of platinum-based catalysts has been greatly improved with the development of better base supports and promoters, the basic chemistry of catalytic reforming remains unaltered, e.g. fresh catalyst performances in pilot plants with promoted platinum catalyst for the same feedstock, pressure, space velocity and temperature are nearly identical to those obtained with the same supported catalyst having platinum only.Moderate differences between operating catalytic reformer units can be expected for the same feedstock and reactor effluent separator pressure, as pressure drop characteristics for the equipment, catalyst inventory and distribution, and hydrogen recycle ratios to the feed differ, in addition to operating practices for the reactor inlet firing pattern. Moderate differences for the same catalytic reformer unit may also be expected, depending upon the catalyst state, but the relationship of increased hydrocracking associated with the classes remains, as this is a function of the ease of conversion into aromatics.

A given reactor is concerned only with the feed components entering and not with the manner in which the original feed components may have been changed into a less desirable class of component, unduly hydrocracked or demethylated to lower carbon number components, i.e. conversion of the original feed components more selectively is an opportunity which, if thrown away, as in the lead adiabatic reactor of existing systems for the naphthenes, can never be regained. On the other hand, once any original feed component has been converted more selectively into aromatics and hydrogen, the succeeding reactor, no matter what its type, cannot hydrocrack any feed component previously converted, because aromatics once formed may be considered thoroughly stable under fired heater and catalytic reforming catalyst conditions. At higher temperatures, aromatics entering may be subjected to limited hydrodealkylation on the catalyst (reduced further by lower hydrogen pressures), but the aromatic ring is never destroyed.

The considerably lower temperature at high space velocities in the alkylcyclohexane converter are responsible for the increased molar selectivity in converting alkylcyclohexanes in the feed into aromatics, as the residence time and temperature significantly reduce the hydrocracking losses which occur. The moderate temperature also limits the isomerization of the original normal paraffins (which are far above thermodynamic equilibrium values in most feedstocks) into isoparaffins. Normal paraffins convert much more readily into aromatics than isoparaffins, and have much better equilibrium values under the same conditions.

Isoparaffins also produce a significant quantity of methane and ethane via gemdialkyl naphthene reactions in forming lower carbon number aromatics from the C8 and higher isoparaffins, whereas normal paraffins convert directly into the same carbon number aromatics. The rising temperature pattern of the alkylcyclohexane converter further limits the hydrocracking, as the least thermally stable alkylcyclohexanes are converted into aromatics at lower temperatures, and, therfore, are not subject to the higher outlet temperature of the converter.

Hydrocracking reactions based upon residence time and temperature estimated to occur in the alkylcyclohexane converter are given below as an indication of the manner in which these hydrocracking losses compare with the foregoing adiabatic reactor values: Fraction Fraction Fraction hydrocracked hydrocracked hydrocracked Feed Alkylcyclohexane Alkylcyclopentane saturates not Component converted converted not converted C6 0.0002 conversion 0.0003 insignificant C, 0.0003 conversion 0.0004 insignificant C8 0.0004 conversion 0.0006 insignificant C9 0.0005 0.0007 0.0008 C,0 0.0005 0.0009 0.0010 The difference in hydrocracking between that for the adiabatic reactors and the alkylcyclohexanes translates directly into increased molar selectivity for the conversion of the original naphthene components into aromatics for the fraction of the feed component undergoing reaction in the alkylcyclohexane converter.

Additional hydrogen resulting from this more selective conversion of the feed components into aromatics is permanent as aromatics and hydrogen enter the succeeding reactor or heater, not the components converted. Directionally, the performance of any succeeding reactor must be improved as the unconverted saturates from the original feed have been subjected to less degradation than in an adiabatic reactor system converting the same amount of alkylcyclohexane in the feed, even though this gain can only be attained by a succeeding reactor. In contrast to the lead adiabatic or intermediate adiabatic reactor of existing catalytic reformer processes, possible conflict in requirements for the alkylcyclohexane converter catalyst does not arise.For the alkylcyclohexane converter, one would select a considerably higher platinum content catalyst with low acidity as one is only concerned with converting the original alkylcyclohexanes in the feed which require the platinum dehydrogenation sites. The low acidity reduces normal paraffin isomerization and hydrocracking.

Promoted platinum catalysts require that the feedstock be preferably hydrotreated to: less than 0.2 ppm for sulfur (desirable to prevent undue catalyst deactivation), the run length for the semi-regenerative reactor arrangement decreases with an increase in the sulfur content of the feed approximately as follows, considering 0.2 ppm sulfur content as base run length: 0.9 base run length for 0.5 ppm sulfur; 0.8 base run length for 1.0 ppm sulfur and 0.7 base run length for 2.0 ppm sulfur); less than 0.5 ppm for nitrogen (desirable in order to prevent ammonium chloride deposits in the reactor effluent coolers as well as neutralization of the acid sites); and less than 4 ppm for the water content, which is easily achieved with proper design by distillation in the hydrotreater. These feed requirements are normally met by hydrotreating the feed naphtha.The normal feed requirements for promoted platinum catalyst admirably suit the alkylcyclohexane converter because excessive water in the feed promotes hydrocracking, and sulfur in the feed attenuates the platinum sites as platinum forms a reversible sulfide which is considerably less active for dehydrogenation. The only precaution necessary for the process of the invention is that any water addition to replace water lost in the separator product stream be made after the naphthene converters of the process of the invention. Continuous water addition is generally practised in order to maintain an optimum water concentration in the recycle gas (specific to the catalyst as related to the proper hydration of the catalyst support) for maximum pentane plus yields.Continuous chlorination is also generally practised for the promoted platinum catalyst, typically adding 0.5 ppm of the feed to replace losses in the separator product streams; chloride addition is normally made at the inlet of the terminal adiabatic reactor, as acidity reactions are only governing for paraffin dehydrocyclization reactions, so no precautions are needed for this location.

It may be seen that the process of the invention before the adiabatic reactors permits a reduction in the conflicting requirements of these reactors, because the most reactive naphthenes in the original feed are efficiently converted into aromatics with comparatively few changes in the other original feed components which are to be converted by the following adiabatic reactors. Commercial feedstocks, in contrast to pure components studied in pilot plants, contain a distribution of carbon numbers and also contain considerable amounts of other classes, even in the case of stocks classified as highly paraffinic (such as the Kuwait naphtha example), highly naphthenic (such as hydrocrackates) or highly aromatic (such as high boiling hydrotreated fluid catalytic cracking gasoline).Although the alkylcyclohexane converter effluent may be used as feed to the lead adiabatic reactor after heating, this feed, even for a Kuwait naphtha feedstock, would contain a considerable amount of original naphthenes (principally alkylcyclopentanes), which can be more selectively converted into aromatics because of the limited molar sensible heat in the terminal reactor effluent stream, which can be economically recovered by the alkylcyclohexane converter having limited catalyst inventory. Therefore, in its preferred form, the process of the invention follows the alkylcyclohexane converter with an alkylcyclopentane converter with moderate temperatures but which are higher than in the alkylcyclohexane converter.Although the construction and the source of heat (preferably flue gases from the radiant fired heater sections) differs, the alkylcyclopentane converter has an ascending temperature pattern similar to that of the alkylcyclohexane converter, as the endothermic heats of reaction are provided by simultaneous heat-exchange. With an approximate entering temperature of 800"F (alkylcyclohexane converter outlet temperature), and an outlet temperature of approximately 840"F, flue gases from the fired radiant heater sections are well suited to supplying the necessary heat for the catalyst tubes with an increase in the overall thermal efficiency of the process.

The alkylcyclopentane route to the alkylcyclohexene intermediate, which is finally dehydrogenated into aromatics in overall paraffin dehydrocyclization is considered by most investigators, from pure component and mixture studies, to be at least as important as tie alkylcyclohexane route. This means that, until the original alkylcyclopentanes are reduced to a sufficiently low concentration as required by the thermodynamic equilibria, unnecessarily high temperatures increase the hydrocracking of the paraffins and the isomerization of the original normal paraffins to less convertible isoparaffins.The time period required to reduce the original alkylcyclopentanes to a sufficiently low concentration for a transition to higher temperatures to be considered (because paraffin dehydrocyclization can then proceed without undue competition) is longer than that needed for the reduction of alkylcyclohexanes even though the average temperature is higher. Catalyst volume anticipated for the alkylcyclopentane converter approximates to twice that provided for the alkylcyclohexane converter.

Because of the importance of the dehydrogenation sites for the reactions, it is recommended that a higher platinum catalyst concentration than that used in the adiabatic converters be employed. This catalyst concentration may be an intermediate concentration such as 0.45 weight per cent platinum, as compared with the 0.60 weight per cent platinum recommended for the alkylcyclohexane converter and the 0.30 weight per cent platinum typically used in the promoted catalyst of the adiabatic reactors. The additional expense for platinum is offset by the fact that a promoted catalyst which might require rhenium (an expensive metal), in an amount of substantially 0.3 weight per cent, is not required for low temperature operation, as the wintering of platinum is negligible at the temperatures of the naphthene converters in the process of the invention.

For the alkylcyclopentanes, lower temperatures than in the adiabatic reactors result in more favourable thermodynamic equilibria (greater driving force) for conversion into the alkylcyclohexene intermediate, itself converted to aromatics.

For the normal paraffins, lower temperatures reduce the reaction rate for isomerization into isoparaffins with. less driving force (lower temperatures correspond thermodynamically to a higher paraffin content), and the high space velocity reduces the residence time for the reaction to occur. This, together with the lower reaction rates for the hydrocracking of the paraffins, acts to preserve the favourable normal paraffin contents until their introduction into an adiabatic reactor at high temperature when both pathways (alkylcyclohexane and alkylcyclopentane) to the desired aromatics are operational.

The catalyst volume in the alkylcyclopentane converter is consistent with reducing the flue gases from the fired radiant heater sections at 19000F (likely temperature for the flue gases entering from a heat recovery arrangement as indicated by figure 6 of the accompanying drawings) to 1050 F. With hydrogen imparting good heat transfer characteristics inside the catalyst tube, and endothermic reactions limiting the temperature rise of the fluid in a manner similar to the vaporization of water in a high-pressure steam generator, the catalyst tubes may be safely finned for more economical heat transfer. The flue gases near atmospheric pressure have comparatively poor heat transfer characteristics as compared with hydrogen (principal component inside the catalyst tube), at approximately 150 psig, which transmits heat to the inner catalyst particles.The adiabatic equivalent temperature drop is approximately 65"F for the alkylcyclopentane converter due to the endothermic reactions overall (slight exothermic reactions occur).

Although the pressure drop (estimated at 14 psi) is significant for the alkylcyclopentane converter, this pressure drop affects only the hydrogen recycle compressor discharge pressure, and does not affect any of the adiabatic reactor outlet pressures. For comparison, a low pressure drop radiant heater, and its adiabatic reactor, (also of low pressure drop design) total approximately 10 psi pressure drop. The alkylcyclopentane converter, together with the alkylcyclohexane converter for the Kuwait naphtha example discussed below, provides a net addition of 0.876 mol of hydrogen per mol of feed. The Kuwait naphtha represents a low endpoint paraffinic feedstock which contains 2.5 liquid volume per cent C paraffins.This example may be considered representative of the lower naphthene content to be encountered in commercial practice, so that the molar hydrogen recycle per mol of feed may be 2.3 for preliminary evaluation work without pioneering in commercial adiabatic reactor practice, as over 3 mols of hydrogen per mol of feed are provided when the feed is first heated to higher temperatures to enter the lead adiabatic reactor. The reduction in hydrocracking by-products due to more selective conversion of the original naphthenes in the feed into aromatics improves the purity of the hydrogen, so that approximately 20 percent less hydrocarbon by-products are recycled with the hydrogen recycle.

This, together with less overall hydrocarbon by-products being formed during the processing of the feed, means that the hydrogen partial pressures in the terminal reactor are higher than are normally provided in conventional adiabatic reactors which operate with 3 mols of hydrogen per mol of feed at the same reactor effluent separator pressure. The higher hydrogen purity of the hydrogen recycle further reduces the volume of hydrogen recycle to be compressed in the process of the invention. Thus, the hydrogen compressor horsepower requirement is decreased for the same operating pressure level, as compared with existing catalytic reformers.Thus, the decrease in volume of the hydrogen cycle appreciatively more than offsets the hydrogen recycle increased circuit pressure drop (allowing no credit towards the decreased pressure drop in the feed reactor heat-exchanger as a result of the alkylcyclohexane converter, or decreased pressure drop in the heaters, adiabatic reactors and aircooler because of the lower hydrogen recycle).

It is recognised that adiabatic reactors require a minimum hydrogen recycle rate in order to limit the temperature drop in the lead reactor. This has been provided even though, with the process of the invention, the adiabatic temperature drops within the adiabatic reactors are considerably decreased, particularly in the lead adiabatic reactor, because the most reactive feed components have been converted into comparatively inert aromatics and hydrogen. With the process of the invention, these products are sensible heat carriers for the other reactions occurring in the lead reactor similar to the original hydrogen recycle (instead of producing endothermic reactions which quench the other reactions from ever occurring in the lead adiabatic reactor).

The process of the invention admirably suits the trend in modern catalytic reformer design of employing radiant fired heater sections exclusively to heat the feed streams to the required inlet temperatures for the respective adiabatic reactors, in order to reduce the hydrogen compressor circuit pressure drop (because compressor horsepower is related to the pressure ratio of the discharge pressure to the suction pressure in absolute terms, circuit pressure drop becomes more important the lower the operating pressure). The process of the invention provides an economical solution for the efficient use of the flue gases from these radiant fired heater sections.With an overall adiabatic temperature drop equivalent, for both the alkylcyclohexane converter and the alkylcyclopentane converter, of 150"F, heat is provided for the elimination of the hydrocracking reactions formerly occurring in the adiabatic reactors (in the conversion of the naphthenes), and is also sufficient actually to reduce the energy required for the adiabatic reactor heaters, if it is assumed that the reactor inlet temperatures remain unchanged. This is a conservative assumption as the average temperature within the adiabatic reactor bed is significantly increased for the same reactor inlet. This may be correct in a new design with lower catalyst inventories in the adiabatic reactors.The overall catalyst inventory in the process of the invention is probably decreased, particularly as the adiabatic reactors are concerned with the slower reactions and most of the catalyst volume is in the adiabatic reactors.

The foregoing should be sufficient to indicate that the process of the invention should contribute towards lowering the energy cost of existing catalytic reforming processes. However, the process of the invention also affects the kinetics of the adiabatic reactors. With feedstock costs rising, an equally important objective of catalytic reforming is the more selective production of aromatics from the feed components, whether for chemicals or for gasoline. The isomerization reactions of paraffins are unimportant contributors to the research clear octanes required currently or which may be expected in the future because of increasing restrictions on the amount of lead additive permitted in gasolines as a result of environmental concern.Typical research clear octanes for the pentane plus paraffins in 98 research clear octane reformate or above correspond to only 62 to 64 research clear octanes. These octane numbers decrease with an increase in the average temperature of the terminal reactor because the single-branched, double-branched and normal paraffins approach thermodynamic equilibria under these temperature conditions. Trimethylranched paraffins, or the isoparaffins for the C7 and higher carbon numbers with the greatest octane numbers, are present in the reformate product in low concentrations only and are far below those which might be expected on the basis of thermodynamic equilibria. Benzene has a 98 research clear octane blending number so this aromatic component may be extracted for chemical values without any effect upon the gasoline value of a 98 research octane clear reformate.Benzene is an exception to the high gasoline blending numbers of aromatics as the C7 and higher aromatics have a research clear blending number of approximately 116.

Adiabatic reactors are economically suited to the slower paraffin dehydrocyclization reactions because of the catalyst volumes required and the high temperatures which require radiant fired heater sections to provide economically the conditions which favour the paraffin dehydrocyclization reactions. Even if they were suitable for the paraffin dehydrocyclization, radiant fired heater tubes filled with catalyst would be much more expensive and have an increased pressure drop.

Furthermore, once the transition to paraffin dehydrocyclization reaction is permitted, without undue competition from the remaining alkylcyclopentanes, heating to higher inlet temperatures for the inlet into the adiabatic reactor is actually beneficial as it is done in the absence of catalyst. Higher temperatures, up to a certain temperature, increase the molar selectivity for the conversion of paraffin components into aromatics. This has been demonstrated by pilot plant studies with various pure paraffin components as the feedstocks. These studies show that the molar selectivity for the conversion of paraffins into aromatics generally increases with temperature up to a certain temperature, depending upon the component, i.e., although the hydrocracking rates increase with higher temperatures, the overall paraffin dehydrocyclization rates increase appreciably faster.

The problem with conventional adiabatic systems is that the lead adiabatic reactor and, possibly, the intermediate adiabatic reactor must also process naphthenes from the original feed for which they cannot perform efficiently. The process of the invention offers the solution to these conflicting requirements to essentially reducing the feedstock to paraffins being the principal remaining components in the feed to the adiabatic reactors for conversion into aromatics, i.e., the product from the alkylcyclopentane converter, at least for low naphthene feedstocks, resembles those involved in pure component pilot plant studies when the feedstock has been partially reacted into a mixture of paraffins and aromatics.

The alkylcyclopentane converter effluent in the Kuwait naphtha example does have a significant original concentration of alkylcyclopentanes remaining; namely, methylcyclopentane (approximately 30 percent of the original feed component remains unconverted) and C7 alkylcyclopentanes (approximately 10 percent of the original feed component). These lower carbon number alkylcyclopentanes have the slowest reaction rates and the most unfavourable thermodynamic equilibria.

Significant quantities of these components also appear in the reformate product.

Even when 99.5 percent normal hexane has been charged to a pilot plant, measurable quantities of methylcyclopentane appear in the product. Similarly, when 99.5 percent normal heptane has been charged to the pilot plant, measurable quantities of C, alkylcyclopentanes appear in the product. These only confirm the thermodynamic equilibria relationships and reaction pathways. With increased reactor volumes, one may lower these alkylcyclopentanes still further towards equilibria. This might increase the selectivity of these naphthenes but, because of the considerable amount of paraffins with different reactivities due to different carbon numbers, one would probably lose overall because of poor paraffin conversion; thus, more time would be needed for the normal paraffins to isomerize to a less convertible form.The latter indicates that, unlike pure component studies, commercial feedstocks are mixtures and require a compromise. It can be seen that, once paraffins become predominant and the naphthenes are sufficiently reduced, heating in the absence of a catalyst is actually kinetically desirable as the paraffin conversion governs. This is exactly what the adiabatic reactor does, so that the combination of the process of the invention of first reducing the original nahthene content and then following with the adiabatic reactors is kinetically desirable. The moderation of the adiabatic temperature drops in the adiabatic reactors achieved by the process of the invention can only improve the adiabatic reactor kinetic performance as the bed outlet temperatures increase.For given reactor volumes, this is inevitable as the temperature drop is related to the reactivity of the components entering, and the most reactive, i.e., those components having the highest speed of reaction are insignificant in the adiabatic reactor feed due to conversion in the naphthene converters. The remaining feed components have a higher hydrocracking exothermic reaction, which offsets their endothermic heat of reaction from the aromatic formation.

With a feed consisting entirely of a mixture of paraffins, the largest temperature drop still occurs in the lead adiabatic reactor, and the smallest temperature drop in the terminal reactor, because of the reactivities associated with the various carbon numbers and individual classes of paraffins together with the concentration differences this would alone tend to produce a smaller temperature drop in succeeding reactors even for an one component feed if reactor volumes were equal.The principal difference encountered with the process of the invention is that the sharply falling temperature decline in the lead adiabatic reactor is changed into a gradual decline, as the components with the highest activities and endothermic requirements are absent from the feed to this reactor, i.e., more similar reactivities are involved as well as significantly lowered endothermic requirements per mol of aromatic produced.

The process of the invention increases the catalyst efficiency in the lead and intermediate reactors as the temperature of the lower catalyst volume of the bed is sufficently increased for a greater paraffin dehydrocyclization reaction to occur due to the higher outlet temperatures for a given inlet temperature. This more effective use of the catalyst in the adiabatic reactor results in significantly less saturates entering the fired heater for the succeeding reactor and being subject to thermal cracking conditions. Above 975"F, the hydrocracking of the pentanes (present from earlier hydrocracking reactions) becomes significant.Above this temperature, therefore, the pentanes plus yields and hydrogen purity always decrease with further increases in temperature, as, unlike the C8 and higher paraffins, aromatics cannot be formed, and only methane through butanes can result. The process of the invention, through increasing the average bed temperature for a given reactor inlet, permits operation with a reasonable catalyst volume and not exceeding this temperature, even when higher octane reformates are required. With a semi-regenerative adiabatic reactor arrangement, the more efficient use of the catalyst in the adiabatic reactors should result in a considerable increase in run lengths because the carbonaceous deposits within the reactor are more uniform.

What is often not appreciated is that the hydrogenolysis of hexanes results in appreciable methane and ethane by-products unlike the hydrocracking of other components. Methane and ethane because of their volatility characteristics significantly affect the hydrogen purity of the recycle gas. The hexane components subject to hydrocracking originate from the hydrocracking of the C9 and higher components as they are not present generally in the original feed to the extent that they are found in the terminal reactor where the highest temperatures prevail.

Although the following table is applicable for temperatures below 840 F., it indicates the difference in by-products from the hydrocracking of the various components: By-products mol per mol of feed component hydrocracked Component CH4 C2H0 C3H8 C4Hto C5Hl2 CH,4 C7Htff C8H18 Ce 0.374 0.236 0.780 0.236 0.374 - - - C, 0.020 0.015 0.965 0.965 0.015 0.020 - - C8 0.003 0.003 0.430 1.128 0.430 0.003 0.003 C9 - - 0.215 0.779 0.779 0.215 - - C10 - - 0.134 0.547 0.678 0.547 0.134 - C11 - - 0.106 0.407 0.485 0.485 0.407 0.106 The foregoing corresponds to a hydrogen consumption of 1 mol per mol of paraffin component hydrocracked and 2 mols per mol of naphthene component hydrocracked.The hydrogen consumption per mol of component hydrocracked significantly increases with higher reactor feed inlet temperatures as more methane and ethane abstractions by hydrogenolysis are superimposed upon all components, being hydrocracked together with the higher carbon number by-products being also subjected to hydrocracking as a result of their higher adsorbability by the catalyst. Thus, it may be seen that, in view of the earlier disappearance of the remaining saturate components together with the elimination of the hydrocracking by-products from the naphthenes (due to being previously converted less selectively in the adiabatic reactors) the process of invention should further lower the production of light hydrocarbons which lowers the hydrogen purity of the hydrogen recycle in addition to increasing the hydrogen yield. This intangible has not been evaluated in the example.Higher cut point feedstocks, because they contain more C9 and higher saturates, therefore favour the process of the invention.

Because the advantages of the process of the invention in the case of feedstocks of high naphthene content and higher cut point will be self-evident, Kuwait naphtha feedstock is chosen as an example. This feedstock represents a good design consideration as most refineries must consider the possibility of processing lower naphthene content feedstocks. The example represents a typical naphtha feedstock when mid-distillate winter demands must be met from Middle East crudes.

Most catalytic reformers must process a variety of feedstocks which vary with the crude available and seasonal demands. Even though designed for the example feedstock, the catalytic reformer of the invention is capable of processing a highly naphthenic feedstock much more efficiently than existing reformers. Highly naphthenic feedstock contains a considerable amount of paraffins to be dehydrocyclized into aromatics as one would generally increase the octane value of the reformate product for incorporation into a premium gasoline.

One may adapt the adiabatic reactors to this situation by limiting the lead adiabatic reactor to a suitable lower reactor inlet temperature which would convert most of the remaining naphthenes. The naphthene converters of the process of the invention automatically adjust to the new situation in that considerably more naphthenes (and almost all the alkylcyclohexanes) will be converted into aromatics than in the Kuwait design case, so that the outlet temperature from the alkylcyclopentane converter entering the lead reactor heater will be considerably lower in value. Additional heater capacity is not required because of the lower reactor inlet temperature to the lead adiabatic reactor. The remaining adiabatic reactors are operated in a normal manner suited to accomplishing the remaining paraffin dehydrocyclization requirement for the octane demand. One may, faced with continuous processing of this feedstock, wish to carry the process of the invention further and utilize radiant fired heater tubes to continue the ascending temperature pattern, in order more selectively to convert the naphthenes for such a feedstock, and then continue processing in the adiabatic reactors for the conversion of the paraffins. Such a design, while improving the overall conversion of a high naphthene feedstock, would probably not be as efficient as the lower cost design of the example if required to process lower naphthene content feedstocks.

Kuwait Naphtha Example Kuwait naphtha, generally termed a highly paraffinic feedstock, illustrates that all commercial feedstocks contain a significant quality of the other classes. PONA analysis, as it is generally stated in liquid volume percent, considerably overstates the paraffinic contribution and understates the naphthenes and native aromatics when translated into a molar equivalent.This is illustrated for the example below: PONA, liquid volume percent P 62.8, N 28.8, A 8.4 (0.7379 specific gravity) PONA, molar percent P 58.59, N 30.61, A 10.80(91.08 molecular weight) Liquid volume Specific Weight Mol Component fraction gravity fraction fraction Benzene 0.005 0.8845 0.0060 0.0085 CZECH 0.013 0.7834 0.0138 0.0180 C6CP 0.012 0.7535 0.0122 0.0131 nC6P 0.020 0.6641 0.0180 0.0230 iC8 P 0.005 0.6640 0.0044 0.0056 Toluene 0.027 0.8719 0.0318 0.0379 C, CH 0.051 0.7740 0.0535 0.0597 C, CP 0.043 0.7579 0.0442 0.0494 nC7P 0.136 0.6882 0.1268 0.1324 iC, P 0.101 0.6870 0.0941 0.1031 C6 A 0.039 0.8720 0.0461 0.0477 C8 CH 0.067 0.7767 0.0705 0.0690 C8 CP 0.041 0.7730 0.0430 0.0421 nC8 P 0.102 0.7068 0.0977 0.0940 iC8 P 0.094 0.7061 0.0900 0.0865 C8A 0.013 0.8669 0.0153 0.0139 C9 CH 0.029 0.7775 0.0305 0.0266 C9 CP 0.019 0.7765 0.0200 0.0173 nC9 P 0.068 0.7217 0.0665 0.0568 iC8 P 0.072 0.7379 0.0720 0.0616 C,0 CII 0.006 0.7800 0.0064 0.0051 C,OCP 0.007 0.7790 0.0076 0.0058 iC,O P 0.030 0.7300 0.0297 0.0229 Total 1.000 1.000 1.000 The Kuwait naphtha feedstock is converted into 98 research clear octane reformate with typical adiabatic reactor yields when the reactor effluent separator is 115 psia, as the base case. This base case is modified with the process of the invention being used at the same reactor separator pressure with paraffin disappearences in the adiabatic reactors maintained. Selectivity to aromatics in the adiabatic reactors for the unconverted saturates are assumed to be unchanged except for the following conservative decreases in hydrocracking of the paraffins in the lead adiabatic reactor because paraffin dehydrocyclization reactions may commence immediately upon introduction of the feed (time temperature for a lead adiabatic reactor under start of run conditions);; C6 P 0.0025 fraction of feed component entering C, P 0.0032 fraction of feed component entering nC9 P 0.0044 fraction of feed component entering iC9 P 0.0043 fraction of feed component entering nC9 P 0.0062 fraction of feed component entering iC9 P 0.0058 fraction of feed component entering iC10 P 0.0093 fraction of feed component entering The decrease in hydrocracking obtained with the use of these figures is only 18 percent of the total decrease in hydrocracking produced by more selected conversion of the alkylcyclohexanes and alkylcyclopentanes in the original feed by the naphthene converters of the process of the invention.No credit has been assumed for any decrease in succeeding reactors nor has any credit been taken for the decreased hydrocracking to aromatics produced by these paraffins because of teinperature conditions (more moderate temperature drops) caused by the process of the invention. It may be that the incremental gain from the process of the invention in moderating the temperature drops of existing adiabatic reactors will be even more important than the more selective conversion of naphthenes into aromatics. This is because of the existing low adiabatic temperature outlets for the lead and intermediate reactor positions, together with the considerably higher quantity of hydrocracking by-products and the much greater potential for the improvement of selectivity to aromatics of the paraffin components.Data at 100 psig for some paraffin components in the pilot plant indicate that the incremental yield of aromatics is more appreciable in the temperature range from 885 to 932"F than for that observed above 932"F, e.g., an 1.850F temperature rise was approximately equivalent to an 10F temperature rise at 8850 F, in producing the same incremental amount of aromatics.

Comparison Summary Adiabatic reactors modified according to Adiabatic reactors the process of the before modification invention Hydrogen recycle, mol fraction 0.0886 0.906 Yields, weight fraction on feed H2 0.0317 (1530SCF/B) 0.0334 (l69lSCF/B) Ct to C4 0.1002 0.0862 C5 plus (98.0RON clear, 0.7918 (98.6 RON clear, 0.8009 lvf on feed) lvf on feed) Paraffins 0.2316 0.2266 Naphthenes 0.0130 0.0130 Aromatics 0.6235 0.6408 Subtotal 0.8681 0.8804 Total 1.0000 1.0000 Benzene 0.0181 0.0189 Toluene 0.1477 0.1530 C6 aromatic 0.2670 0.2733 C9 aromatics 0.1788 0.1825 C10aromatics 0.0119 0.0131 WHAT I CLAIM IS: 1.A catalytic reformer process for the production of hydrogen and a product of enhanced aromatic content from a crude naphtha feedstock which process comprises a plurality of reactor stages in which the feedstock is subjected to naphthene conversion, wherein the heat in the effluent from the terminal reactor stage is utilised in an alkylcyclohexane converter, by means of simultaneous heatexchange, to provide the endothermic heat of reaction required in the alkylcyclohexane converter for the dehydrogenation to aromatics of the original alkylcyclohexanes in the feed, before the said feed enters the naphthene conversion reactor stages.

2. A process as claimed in Claim 1, which further comprises an alkylcyclopentane converter stage, which also precedes the naphthene conversion stages.

3. A process as claimed in Claim 1 or Claim 2, wherein the catalyst employed is a promoted platinum catalyst.

4. A process as claimed in Claim 3, wherein the feedstock is hydrotreated to less than 0.2 ppm sulphur, less than 0.5 ppm nitrogen and less than 4 ppm water.

5. A process as claimed in any of Claims 1 to 4, wherein the hourly space velocity in the alkylcyclohexane converter is 20.

6. A process as claimed in any of Claims 1 to 5, wherein the inlet temperature in the alkylcyclohexane converter is approximately 780"F.

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

Claims (11)

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

   The decrease in hydrocracking obtained with the use of these figures is only 18 percent of the total decrease in hydrocracking produced by more selected conversion of the alkylcyclohexanes and alkylcyclopentanes in the original feed by the naphthene converters of the process of the invention. No credit has been assumed for any decrease in succeeding reactors nor has any credit been taken for the decreased hydrocracking to aromatics produced by these paraffins because of teinperature conditions (more moderate temperature drops) caused by the process of the invention.It may be that the incremental gain from the process of the invention in moderating the temperature drops of existing adiabatic reactors will be even more important than the more selective conversion of naphthenes into aromatics. This is because of the existing low adiabatic temperature outlets for the lead and intermediate reactor positions, together with the considerably higher quantity of hydrocracking by-products and the much greater potential for the improvement of selectivity to aromatics of the paraffin components.Data at 100 psig for some paraffin components in the pilot plant indicate that the incremental yield of aromatics is more appreciable in the temperature range from 885 to 932"F than for that observed above 932"F, e.g., an 1.850F temperature rise was approximately equivalent to an 10F temperature rise at 8850 F, in producing the same incremental amount of aromatics.

   Comparison Summary Adiabatic reactors modified according to Adiabatic reactors the process of the before modification invention Hydrogen recycle, mol fraction 0.0886 0.906 Yields, weight fraction on feed H2 0.0317 (1530SCF/B) 0.0334 (l69lSCF/B) Ct to C4 0.1002 0.0862 C5 plus (98.0RON clear, 0.7918 (98.6 RON clear, 0.8009 lvf on feed) lvf on feed) Paraffins 0.2316 0.2266 Naphthenes 0.0130 0.0130 Aromatics 0.6235 0.6408 Subtotal 0.8681 0.8804 Total 1.0000 1.0000 Benzene 0.0181 0.0189 Toluene 0.1477 0.1530 C6 aromatic 0.2670 0.2733 C9 aromatics 0.1788 0.1825 C10aromatics 0.0119 0.0131 WHAT I CLAIM IS: 1.A catalytic reformer process for the production of hydrogen and a product of enhanced aromatic content from a crude naphtha feedstock which process comprises a plurality of reactor stages in which the feedstock is subjected to naphthene conversion, wherein the heat in the effluent from the terminal reactor stage is utilised in an alkylcyclohexane converter, by means of simultaneous heatexchange, to provide the endothermic heat of reaction required in the alkylcyclohexane converter for the dehydrogenation to aromatics of the original alkylcyclohexanes in the feed, before the said feed enters the naphthene conversion reactor stages.
2. 2. A process as claimed in Claim 1, which further comprises an alkylcyclopentane converter stage, which also precedes the naphthene conversion stages.
3. 3. A process as claimed in Claim 1 or Claim 2, wherein the catalyst employed is a promoted platinum catalyst.
4. 4. A process as claimed in Claim 3, wherein the feedstock is hydrotreated to less than 0.2 ppm sulphur, less than 0.5 ppm nitrogen and less than 4 ppm water.
5. 5. A process as claimed in any of Claims 1 to 4, wherein the hourly space velocity in the alkylcyclohexane converter is 20.
6. 6. A process as claimed in any of Claims 1 to 5, wherein the inlet temperature in the alkylcyclohexane converter is approximately 780"F.
7. 7. A process as claimed in any of Claims 1 to 6, wherein the outlet temperature

   in the alkylcyclohexane converter is approximately 800cm.
8. 8. A process as claimed in any of Claims I to 7, wherein water addition to replace water loss in the separation of the product streams is effected after the naphthene conversion stages.
9. 9. A process as claimed in Claim 8, wherein water addition is effected continuously to maintain optimum water concentrations in the recycle gas.
10. 10. A process as claimed in any of Claims I to 9, wherein radiant fired heaters are used to heat the feed streams to the required inlet temperatures.
11. 11. A process as claimed in Claim 1, substantially as herein described with reference to the accompanying drawings and/or the specific Example.