US3389075A - Process for producing aromatic hydrocarbons and liquefied petroleum gas - Google Patents

Description

June 18, 1968 G. E. ADDISON Filed May 11, 1966 AND LIQUEFIED PETROLEUM Ci; arge Sto c/r GAS 2 Sheets-Sheet 1 Figure First Reaction Zone \M I /5 t f 4 l3 7'3 First Separation Zone/ Aromatics Separation Zone I Aromatics Product 20 7 rl8 Second Reaction a Zone a Hydrogen Containing Gas Product 9% l9' Second Separation Zone l /0'- Liauit'ied Petroleum Gas Product l l \zll I N VE N TOR l I George E. Addison Third Reaction Zone Wye-d2 VA TTORNEYS s. E. ADDISON 3,389,075 PROCESS FOR PRODUCING AROMATIC HYDROCARBONS I June 18, 1968 AND LIQUEFIED PETROLEUM GAS 2 Sheets-Sheet 2 Filed May 11, 1966 N Fiat //V V EN TOR: George E. A ddison A TTORNEXS United States Patent 3,389,075 PROCESS FOR PRODUCING AROMATIC HY- DROCARBONS AND LIQUEFIED PETRO- LEUM GAS George E. Addison, Mount Prospect, 111., assignor to Universal Oil Products Company, Des Plaines, 111., a corporation of Delaware Filed May 11, 1966, Ser. No. 549,287 6 Claims. (Cl. 208-66) ABSTRACT OF THE DISCLOSURE Process for production of high-purity aromatics and liquified petroleum gas by catalytic reforming a naphtha in a first reaction zone, extracting aromatics from the reformate, hydrocracking the resulting paraifinic raffinate to produce LPG in a second reaction zone, and recycling the C -lfraction of the hydrocracking effluent to a third reaction zone for dehydrocyclization to aromatics. In addition, C and G, from the various fractionating columns within the process, and all C aromatics from an arcmatics fractionation train which produces benzene, toluene, and C aromatics, may be recycled to the hydrocracking zone for additional synthesis of LPG. By operating the initial catalytic reformer of the first reaction zone to maximize aromatic synthesis, suflicient hydrogen is produced to supply the requirement of the hydrocracking zone. In the preferred embodiment, the initial catalytic reformer contains multiple-stage catalytic contact sections and the third reaction zone wherein the 0 fraction of the hydrocracking efiluent is dehydrocyclized, comprises the last of the series of catalytic contact sec tions.

This invention relates to a combination process for producing aromatic hydrocarbons and liquefied petroleum gas (LPG) from a petroleum fraction while minimizing the production of gasoline. In particular, this invention relates to a combination process for effecting the production of aromatic hydrocarbons and the production of LPG with flexibility in effecting distribution between the two desired product streams in accordance with the fluctuating needs for supplying the markets. More particularly, this invention affords a combination process which includes a hydrogen producing zone and a hydrogen consuming zone operated in combination with other processing techniques, for the production of increased yields of aromatic hydrocarbons and LPG, while simultaneously functioning in hydrogen balance such that no external hydrogen source is required.

Such a combination process has its primary applicability in the manufacture of petrochemicals by chemical companies who do not have the marketing facilities which the necessary for the disposal of a gasoline product or by petroleum refiners who have sufficient gasoline product available from other sources and who wish to enhance the profitability of their operations through expansion into petrochemicals. Chemical companies owning crude oil production with little or no refining capacity of their own, must sell their crude at market price or at contract price to acquire some return on their investment. Since the hydrocarbons available in crude oil provide the basic building blocks for organic chemical synthesis, maximum profitability for the crude lies in upgrading it to final chemical products. With an awareness of this fact, many chemical companies now purchase hydrocarbon fractions from petroleum refiners for the specific purpose of converting them to organic chemicals, plastics, rubber, etc. In many cases, these companies are burdened with the disposal of byproduct petroleum fractions. Some domestic pertoleum refiners and, more particularly, some foreign 3,389,075 Patented June 18, 1968 ownership is not sufficiently prevalent to afford a substantial gasoline market, and the petroleum refining industry in such nations is oriented to the maximum production of fuel oil, LPG, and petrochemicals.

In view of the profitability advantage in expanding into petrochemicals, the hydrocarbon processing industry has seen a world-wide trend away from emphasis upon the production of high octane gasoline by catalytic reforming, and a shift to maximize aromatics production by catalytic reforming. With a similar economic objective, petroleum refiners operating a foreign nations or domes tic areas which have little or no natural gas have shown an increasing trend to install high pressure catalytic reforming in order to decrease high octane gasoline yield while increasing LPG production. The present invention affords a method of achieving both of these refining ends simultaneously.

In one embodiment, these ends are achieved by introducing a hydrocarbon charge stock in admixture With hydrogen into a reforming zone containing a reforming catalyst which is maintained substantially at dehydrogenation and dehydrocyclization conditions; separating the resulting effluent into a hydrocarbon liquid and a hydrogen rich gas, which gas is further separated, into one portion which is recycled to said catalytic reforming zone and one portion which is sent to a hydrocracking zone as hereinafter set forth; separating said hydrocarbon liquid from the reforming zone into at least two fractions in the first fractionation zone, one of which fractions comprises pentane and lighter hydrocarbons and one of which comprises hexane and heavier hydrocarbons; introducing said hexane and heavier hydrocarbon stream into an aromatics separation zone wherein substantially pure aromatic hydrocarbons are recovered as the desired products and substantially pure parafiinic hydrocarbons are recovered as a raffinate stream; introducing said parafiinic raffinate stream into the hydrocracking zone in admixture with the hydrogen rich gas from the catalytic reforming zone in the presence of a hydrocracking catalyst; separating the effluent from the hydrocracking zone into a hydrocarbon liquid and a hydrogen containing gas, which gas is further separated into at least a portion which is recycled back to the hydrocracking zone and a portion which is discharged as an excess hydrogen containing gas; introducing the hydrocarbon liquid from the hydrocracking zone into a second fractionation zone wherein the hydrocarbon stream is separated into at least a propane-butane fraction comprising the desired LPG product, a light gasoline fraction which is recycled back to the hydrocracking zone, and a heavy gasoline fraction which is recycled back to the reforming zone for dehydrocyclization. In a modified operation, the pentane and lighter hydrocarbon liquid from the first fractionation zone will be admixed with the efliuent from the hydrocracking zone and the combined stream will be sent to the second fractionation zone wherein the separation above described occurs.

In a modified embodiment of this invention, these desired ends are achieved by introducing a hydrocarbon charge stock in the presence of hydrogen into a catalytic reforming zone comprising a plurality of reactor sections which contain reforming catalyst and which are maintained substantially at dehydrogenation and dehydrocyclization conditions. The processing procedure is then identical to the embodiment described hereinbefore with the exception that the heavy gasoline fraction from the second fractionation zone is not recycled back to the entire reforming zone but is recycled only to the last of the plurality of reactors containing the reforming catalyst, wherein the catalyst in the last reactor section is maintained substantially a dehydrocyclization conditions.

In another embodiment of this invention, these desired ends are achieved by introducing a hydrocarbon charge to a first reforming zone in admixture with hydrogen in the presence of a reforming catalyst which is maintained at substantially dehydrogenation and dehydrocyclization conditions; separating the resulting efiluent into a hydrocarbon liquid and a hydrogen rich gas, which gas is further separated into at least a portion which is recycled to the first reforming zone and a portion which is sent to a second reforming zone; separating said hydrocarbon liquid from the first reforming zone into at least a pentane and lighter hydrocarbon fraction and a hexane and heavier hydrocarbon fraction in a first fractionation zone; introducing said hexane and heavier fraction into an aromatics separation zone wherein substantially pure aromatic hydrocarbons are recovered as desired products and substantially pure par-affinic hydrocarbons are recovered as a raflinate stream; introducing said paraffinic raffinate stream into a hydrocracking zone in admixture with hydrogen in the presence of a hydrocracking catalyst; separating the efiluent from the hydrocracking zone into a hydrocarbon liquid and a hydrogen containing gas, which gas is further separated into a portion which is recycled back to the hydrocracking zone and a portion which is discharged as an excess hydrogen containing gas; introducing the hydrocarbon liquid from the hydrocracking zone in admixture with the pentane and lighter hydrocarbon fraction from the first fractionation zone into a second fractionating zone wherein the mixed hydrocarbon stream is separated at least into a propane-butane fraction comprising the desired LPG product, a light gasoline fraction which is recycled back to the hydrocracking zone, and a heavy gasoline fraction; introducing said heavy gasoline fraction in admixture with the hydrogen rich gas from the first reforming zone into a second reforming zone which contains a reforming catalyst which is maintained substantially at dehydrocyclization conditions; separating the efiluent from the second re forming zone into a hydrocarbon liquid and a hydrogen rich gas, which gas is further separated into a portion which is recycled to the second reforming zone and a portion which is introduced int-o the hydrocracking zone; introducing the hydrocarbon liquid from the second reforming zone into the first fractionating zone in admixture with the hydrocarbon liquid from the first reforming zone.

The charge stocks which may be reformed in accordance with the present invention comprise gasoline boiling range hydrocarbons containing naphthenes, paraffins, and aromatics with only minor amounts of olefins being present. Suitable hydrocarbon charge may be a straight run gasoline or a natural gasoline, or it may be a refined gasoline such as a thermal cracked gasoline or a hydrocracked gasoline, etc., or it may comprise any combination thereof. The gasoline may be a full boiling range gasoline fraction having an initial boiling point of from about 50 F. to about 100 F. and an end boiling point of from about 375 F. to about 425 F., or it may be a selected fraction thereof. The usual fraction to be processed in the subject invention will be a naphtha fraction having an initial boiling point of from about 150 F. to about 250 F. and an end boiling point of from about 350 F. to about 425 F. Where such fraction comprises substantial sulfur or nitrogen containing hydrocarbons or where substantial olefinic hydrocarbons are included, hydrogen pretreatment will be necessary in order to protect the reforming catalyst from loss of stability due to premature excessive carbonization.

Hydrogen pretreatment of contaminated hydrocarbon charge stocks is well known in the art of hydrocarbon processing, and a preferred method is shown in US. Letters Patent No. 2,878,180. However, this preference should in no way be construed to limit the present invention. Any hydrocarbon charge stock containing more than about 10 parts per million (p.p.m.) by weight of sulfur and/or more than about 1.0 ppm. of nitrogen and/or more than about 1.0 volume percent of olefinic hydrocarbons should be so treated. Such pretreatment will also serve to remove trace quantities of arsenic, lead, copper, nickel, vanadium, tungsten, and other metals which may be present in untreated hydrocarbon fractions and which may be detrimental to noble metal reforming catalysts. Such purification is effected by passing the hydrocarbon charge in admixture with hydrogen across a suitable catalyst at a pressure of from about p.s.i.g. to about 1000 p.s.i.g., the operating pressure being dependent upon the composition or type of charge stock being processed. The hydrogen not only serves as a reactant in effecting the purification of the hydrocarbons, but it also affords a method for protecting the catalyst against excessive carbonization and it provides a thermal sink for the exothermic heat of reaction. Hydrogen is, therefore, normally present at a concentration of from about 100 standard cubic feet per barrel (s.c.f.b.) of hydrocarbon charge to about 3000 s.c.f.b., the amount again being dependent upon the type of charge stock being processed. The temperature of the hydrogen pretreating reaction zone is maintained in the range of from about 500 F. to about 900 F. The actual temperature required will necessarily vary in accord with the degree of contamination, the type of stock being processed, and with the activity level of the catalyst. The hydrocarbon is normally processed at a liquid hourly space velocity in the range of about 1.0 to about 10.0. A suitable catalyst for such hydrogen treating of hydrocarbons comprises alumina, silica and a Group VIII metal 'or a Group VI-B metal or any combination of metals thereof. The metals of Groups VI-B and VIII are intended to include those indicated in the Periodic Chart of the Elements, Fischer Scientific Company, 1953. A more preferred catalyst is comprised of alumina, silica, a nickel compound, a molybdenum compound, and a cobalt compound wherein the nickel, molybdenum and cobalt compounds are more specific ally the oxides or sulfides.

The efiluent hydrocarbon from the hydrogen pretreating unit is subsequently fractionated to remove dissolved hydrogen, hydrogen sulfide, ammonia, and paraifin gases comprising methane and ethane. In the present invention gaseous propane is maintained in the liquid hydrocarbon stream which is then stabilized to provide a light liquid hydrocarbon comprising propane, butane, and pentane, and a heavy liquid comprising hexane and heavier hydrocarbons, which heavy liquid is then charged to the catalytic reforming zone.

The art of catalytic reforming has seen world wide acceptance with recent trands toward diversification of purpose. The basic processing technique and a preferred catalyst are indicated in U.S. Patent Nos. 2,479,109 and 2,479,110, issued to Vladimir Haensel, wherein the catalyst comprises alumina, platinum, and halogen. Reforming is undertaken at a temperature in the range of from about 600 F. to about 1100 F.; at a pressure in the range of from about 50 p.s.i.g. to about 1000 p.s.i.g., but more normally at about 400 to about 500 p.s.i.g.; at a liquid hourly space velocity in the range of from about 0.5 l./hour to about 10.0 l./hour; and in the presence of from about 0.5 to about 10.0 moles of hydrogen per mole of hydrocarbon. As understanding of the reaction mechanisms occurring Within the reforming zone has increased, it has become poss ble to adjust operating techniques to enhance the specific reaction desired. Thus it was the primary purpose of catalytic reforming to subject a substantially sulfur, nitrogen, oxygen, olefin, and metal free gasoline boilrng range or naphtha boiling range charge t k to hi h temperature and pressure in the presence of hydrogen in order to enhance the anti-knock properties of the hydrocarbons contained therein. It was determined that such enhancement, resulting in a high octane gasoline product, was derived from four specific chemical reactions; (1) the dehydrogenation of naphthenic hydrocarbons to produce the corresponding aromatic derivative, (2) the dehydrocyclization of parafiinic hydrocarbons to produce corresponding aromatic hydrocarbons, (3) the hydrocracking of high molecular weight hydrocarbons to produce lower molecular weight hydrocarbons, and (4) the isomerization of normal paraffinic hydrocarbons to produce branched chain isomers of equal molecular weight. Each of these four reaction mechanisms upgrade low octane hydrocarbons to high octane hydrocarbons, but as the automotive manufacturers have increased engine compression ratios it has become necessary to adjust operating techniques in order to control the reaction mechanisms selectively to maximize octane with minimum loss of liquid product yield and minimum production of parafiinic gas (methane, ethane, and propane). It has thus been determined that the dehydrogenation of naphthenes to aromatics is promoted by operating at lower pressure levels; that dehydrocyclization of parafiins to aromatics is promoted by low pressure and high temperature; that hydrocracking of parafiins is promoted by high pressure, high temperature, and high residence time of the charge stock on the catalyst; and that isomerization of paraflins is promoted by intermediate temperature, and a catalyst comprising a much higher halogen content than normally employed. Since aromatic hydrocarbons have higher octane ratings than other hydrocarbons of equivalent molecular weight, catalytic reforming showed a tendency to operate at higher temperatures and lower pressures in order to enhance the resulting gasoline octane rating by increasing the aromatic hydrocarbon content of the gasoline. Such high octane aromatic containing gasoline initiated the diversification of catalytic reforming into the field of petrochemicals production. As aromatic hydrocarbons became high in price due to in creased chemical consumption without an increase of aromatics production from the By-Product Coke Industry, low pressure catalytic reforming in conjunction with solvent extraction became the primary source of aromatic hydrocarbons for chemical use.

A recent development in catalytic reforming, has been its utilization in the production of LPG in conjunction with high octane gasoline production. In such installations, the catalyst is maintained at operating conditions which enhance the hydrocracking characteristics of the catalyst. Thus the catalytic reformer is operated at a temperature in the range of about 850 F. to about 1000 F. but the pressure is maintained at about 600 p.s.i.g. to about 800 p.s.i.g. This higher pressure not only results in lower liquid yields due to the formation of parafiin gases, 'but it also results in a lower aromatic content in the resulting gasoline product. This decreased aromatic production is due to three sources. First, although naphthenes continue to dehydrogenate to produce aromatics, the higher pressure allows a greater percentage of naphthene to pass across the catalyst without being dehydrogenated to form aromatics. Secondly, the higher pressure enhances the hydrocracking of the naphthene ring to produce paraffinic hydrocarbon which thus reduces potential aromatic production. Finally, at elevated pressure the paraffinic hydrocarbons are not preferentially dehydrocyclicized to form aromatics but are preferentially hydrocracked to lower molecular weight paraffins. Thus it may be seen that although high octane gasoline may be produce or aromatic hydrocarbons may be produced in conjunction with LPG, it is not possible to maximize these liquid products while producing LPG. It is therefore the purpose of the present invention to maximize the production of LPG without loss of potential aromatics production.

The catalytic reforming unit of the present invention is maintined at operating conditions to enhance the dehydrogenation of naphthenes and the dehydrocyclization of paraffins in order to maximize the production of both aromatic and hydrogen, maximum hydrogen being desired since it is consumed elsewhere in this combination process. The production of aromatic hydrocarbons is enhanced by catalytic reforming at a temperature in the range of from about 850 F. to about 1050 F. and at a pressure in the range of from about p.s.i.g. to about 400 p.s.i.g. when the end boiling point of the charge stock is about 350 F., but when the end point of the charge stock is about 400 F. or more, the preferred pressure is about 500 p.s.i.g. in order to maintain catalyst stability. This higher pressure is preferred on the heavier charge stock since the resulting higher partial pressure of hydrogen will minimize the formation of coke, the cracking of the heavier hydrocarbons to olefins, and the polymerization of olefins to form polynuclear aromatics.

The efiiuent hydrocarbon from the catalytic reforming zone is subsequently fractionated to remove dissolved hydrogen, other inorganic gases, and paraffin gases comprising methane and ethane. In the present invention gaseous propane is maintained in the liquid hydrocarbon stream which is then stabilized to provide a light liquid hydrocarbon comprising propane, butane, and pentane, and a stabilized aromatics-containing reformate comprising hexane and heavier hydrocarbons. The aromatic-containing reformate may be further fractionated to provide a light reformate comprising hydrocarbons having about six to about eight carbon atoms per molecule and a heavy reformate comprising hydrocarbons having about nine or more carbon atoms per molecule. The light reformate having an end boiling point of about 350 F. may then be charged to an aromatics recovery unit wherein benzene, toluene, xylenes, and ethylbenzene are recovered as desired products while the heavy reformate having an initial boiling point of about 350 F. containing heavier aromatics may be recovered as a high octane heavy gasoline blending component or it may be charged to the hydrocracking reaction zone. In the present invention, however, the stabilized reformate need not be split into light and heavy fractions but may be charged directly to the aromatics recovery unit as a full boiling range reformate comprised of hexane and heavier hydrocarbons wherein maximum aromatics may be recovered as products.

The aromatics separation zone may be characterized as a solvent extraction technique, or an aromatics solid adsorption technique, or an extractive distillation technique, or a fractional crystallization technique. A preferred separation method is described by US. Letters Patent No. 2,730,558, but this preferred solvent extraction process should in no way be construed to limit the present invention. A particularly preferred solvent for separating aromatic hydrocarbons from non-aromatic hydrocarbons is a mixture of water and one or more hydrophilic organic solvents. Such a combination solvent may have its solubility regulated by varying the water content. Thus, by adding more Water to the solvent, the solubility of all components in the hydrocarbon mixture is reduced, but the solubility difference between components (selectivity) is increased. The net effect is to decrease the number of contacting stages required to achieve a given purity of product, or to increase the resulting purity of product where the number of contacting stages is held constant. Because of the resulting reduction in solubility due to the increased water content, the throughput of the combination solvent must be increased in order to dissolve the solute at the same production rate. Suitable hydrophilic organic solvents for this process include alcohols, glycols, aldehydes, glycerine, phenol, etc. Particularly preferred solvents are diethylene glycol, triethylene glycol, dipropylene glycol, tripropylene glycol, and mixtures thereof containing from about 2% to about 30% by weight of water. In classifying hydrocarbon and hydrocarbon type compounds according to increasing solubility in such mixed solvent, it is found that parafi'ins are least soluble followed in increasing order of solubility by naphthenes,

olefins, diolefins, acetylenes, sulfur containing hydrocarbons, nitrogen containing hydrocarbons, oxygen containing hydrocarbons, and aromatic hydrocarbons. It may thus be seen that the ideal charge to such a solvent extraction process is one consisting essentially of paraifins and aromatics, and since catalytic reformates contain only minor amounts of naphthenes and olefins, they are well suited to such an aromatics extraction procedure.

The aromatic hydrocarbons which are so recovered.

from the reformate may be further separated by proper fractionation to produce high purity benzene, toluene, xylene isomers, and ethylbenzene. These aromatics are produced in purities which not only meet, but exceed the standard of nitration grade, and they therefore are a preferred source of aromatics for subsequent chemical proccssing. A bottoms product will also be produced which comprises aromatics containing about nine or more carbon atoms per molecule. This heavy aromatics product is normally used as a solvent or as a high octane blending component for the gasoline pool in the refinery, but in the present invention this stream may be combined with the raffinate from the aromatics separation zone for further processing. The rafiinate from the aromatics extraction unit will comprise paraflinic hydrocarbons with slight amounts of naphthenes, olefins, and aromatics. The amount of naphthenes, olefins and aromatics contained within the paraflinic rafiinate stream will depend upon the composition of the catalytic reformate charged to the extraction zone, the solvent circulation rate within the extraction zone, the number of equilibrium stages within the extraction zone, etc. Thus it is normal to charge a reformate having an end boiling point of about 350 F. in order to maximize recovery of benzene, toluene, and xylenes. If the end point of the reformate is about 400 F., the resulting rafiinate will contain a greater amount of benzene, toluene, and xylene as well as the heavier aromatics which boil in the range of about 350 F. to about 400 F. Because of the highly paraffinic nature of the rafiinate stream, it is commonly used as a source of fuel for jet aircraft. In the present invention, however, this paraflinic stream is charged to a catalytic hydrocracking zone for further processing.

Hydrocracking is known in the art of hydrocarbon processing as a means of converting high molecular weight fractions into lower molecular weight fractions in the presence of hydrogen wherein the products are essentially non-olefinic, and less coke is produced than would be experienced if mere thermal cracking were employed. Hydrocracking processes are commonly employed in the conversion of coals, coal hydrogenation liquid products, shale oil, cr-ude liquid products, petroleum crude oils, heavy residual oils, tars, heavy vacuum gas oils, etc., to produce increased yields of liquid hydrocarbons products having improved chemical and physical characteristics. Thus, a typical charge stock having an initial boiling point of about 650 F. and an end boiling point of about 1000 F. will be hydrocracked into lower boiling fractions which may subsequently be more readily refined by other techniques. Such a high boiling range stock may be cracked to produce a gasoline fraction or a naphtha fraction which may be subjected to catalytic reforming, a kerosense fraction which may be a final product or which may be further hydrocracked, a light gas oil or a heavy gas oil which may be subjected to catalytic cracking, fuel oils of various grades as final products, etc., or to produce any combination of such products which may or may not be further processed by any combination of refining techniques as dictated by the economics of the particular commercial refining installation. It may be seen that in order to hydrocrack to maximum economic advantage, the cracking must occur in a manner to maximize the amount of hydrocarbon liquid produced. Thus, a non-selective hydrocracking will split a high molecular weight hydrocarbon randomly and methane, ethane, and propane gases will be produced as well as the desired lower molecular weight liquids. The amount of such undesired mixed hydrocarbon gas which is produced is minimized by a proper catalyst selection and by judicious choice of operating conditions, as determined by the physical and chemical characteristics of the hydrocarbon charge stock. Another disadvantage of non-selective hydrocracking is that increased quantities of coke and hydrocarbonaceous materiais are produced, and, when the hydrocracking is promoted by a catalyst, these materials become deposited thereon resulting in rapid loss of catalytic activity and stability. It has also been determined that the presence of aromatic hydrocarbons within the high boiling charge stock have a detrimental influence upon the hydrocracking catalyst. In addition to high-boiling alkyl-substituted m-ononuclear aromatic hydrocarbons, such aromatics include naphthene, anthracene, phenanthrene, pyrene, etc. and their alkyl-substituted derivatives. The deleterious effect of the presence of aromatic compounds within the charge stock is exhibited in a two-fold manner. First, certain condensed aromatics appear to be adsorbed on the catalyst surface without being either hydrogenated or hydrocracked and the catalytically active sites become effectively shielded from the material being processed. The net result is that catalyst activity and selectivity are prematurely destroyed. Secondly, the presence of aromatic hydrocarbons results in their saturation with hydrogen and the resultant high exothermic heat of reaction promotes premature hydrocracking of the ring as well as the alkyl chain, thus raising the temperature still further and causing localized temperature run-away on the catalyst. Such localized hot-spots in the catalyst bed result in the unabated conversion of normally liquid hydrocarbons into excessively large amounts of light parafiinic hydrocarbon gas, accompanied by the virtually immediate deposition of coke and other hydrocarbonaceous material. Such localized hot-spots may not be eliminated or controlled by merely lowering the operating temperature since an initial triggering temperature is required to initiate hydrogenation and hydrocracking and the exothermicity of reaction is instantaneous. Temperature control is normally achieved by providing the hydrogen in the reaction zones in quantities not only sufficient for the required chemical reactions, but in an excess to provide a thermal quench for the heat of reaction. Thus the exothermic heat of reaction in saturating the aromatic ring is quickly dissipated in raising the sensible heat of the adjacent hydrogen rich gas and the danger of localized temperature run-away is minimized. As an additional safeguard against temperature run-away, the hydrocracking art has seen a recent trend to provide a separate reaction zone specific to the saturation of aromatics over a hydrogenation catalyst such that little or no cracking is possible. Again the hydrogen is provided in an excess sufiicient to quench the exothermicity and hold a tolerable temperature level. The saturated hydrocarbon stream with the hydrogen is then passed into the hydrocracking zone wherein the catalyst selectively hydrocracks to provide liquid hydrocarbons with minimum temperature rise and minor production of parafi'in gases. Such dual-catalyst systems may be established as separate reaction vessels but the preferred method is to provide separate catalyst beds within the same reaction vessel.

It will be noted that the catalytic hydrocracking system which is embodied in the present invention dilfers from the described art in many respects. Whereas propane is an undesired by-product in most cases, it is a desired product in the present invention and higher cracking temperatures may be tolerated. Thus, hydrocracking for liquid production is normally undertaken commercially at a temperature level of from about 750 F. to about 850 F. Whereas hydrocracking to maximize LPG production may be undertaken at about 900 F. with no detriment to catalyst activity or stability. Whereas aromatic hydrocarbons require special precautions as indicated, the charge stock to the hydrocracking zone in the present invention is substantially free of aromatic hydrocarbons. Thus the need for a separate hydrogenation catalyst zone is minimized or even eliminated, and the amount of hydrogen gas being circulated for purposes of thermal quench is substantially reduced with the result that equipment demands, capital expense, and operating expense are thereby reduced. Since aromatics are present in only minor quantities in the parafiinic raffinate charge, there is a considerable reduction in the chemical consumption of hydrogen, little aromatic ring saturation occurring, and the available hydrogen from catalytic reforming is thus more effectively utilized in hydro'cracking to produce the desired LPG product thus eliminating or at least minimizing any need of a hydrogen source external to the present invention. The catalytic hydrocracking to produce LPG may be undertaken, in the presence of hydrogen, at a pressure of from about 1000 p.s.i.g. to about 2000 p.s.i.g. and normally at a pressure of about 1500 p.s.i.g. Whereas the hydrogen circulation for production of liquid hydrocarbons would be in the range of from 3000 s.c.f.b. to about 10,000 s.c.f.b., the circulation rate of the present invention may be in the range of from about 3,000 s.c.f.b. to about 6 ,000 s.c.f.b. The raffinate charge may be at a Liquid Hourly Space Velocity of from about 0.5 to about 10.0 and preferably from about 1.0 to about 4.0. The catalyst for the hydrocracking may comprise silica, alumina, and a Group VIII metal, and more specifically the Group VIII metal may be nickel, platinum, or palladium but the composition of such catalyst should in no way be construed to so limit the present invention.

The final fractionation zone will comprise whatever number of distillation columns is necessary to effect an acceptable separation of the combined parafiinic hydrocarbons within the efiluent from the catalytic hydrocracking zone, the stabilizer light liquid hydrocarbon from the catalytic reforming zone, and the stabilizer light liquid hydrocarbon from the catalytic hydrogen pretreating zone. The first fraction removed will be the dissolved gas comprising hydrogen, methane, and ethane, and this gas stream will normally be utilized .as a fuel gas. The second fraction produced will be primarily propane and butane, and these hydrocarbons will be withdrawn in the liquid state under a slight pressure to constitute the desired LPG product. The next fraction to be removed is rich in the pentane isomers and the low boiling isomers of hexane, including 2,2-dimethylbutane and 2,3-dimethylbutane; these are removed from the process as a light gasoline fraction. If it is the prime purpose to completely eliminate gasoline product in the present process, this light gasoline fraction should not be recycled to the catalytic reforming zone since the chemical structure of the included hydrocarbons will prevent dehydrocyclization to romatics, but this fraction may be recycled to the hydrocracking zone for further cracking to produce LPG. Unfortunately, while the included pentane and hexane isomers may be further cracked, they may not be selectively hydrocracked to produce LPG. Due to the chemical structure of the included hydrocarbons, the prevailing products which result comprise methane, and ethane. Since these parafiinic gases have a minimum economic value and since their presence reduces the hydrogen purity on the hydrocracking catalyst bed the preferred embodiment of this invention is to remove this light gasoline fraction entirely from the process as a final by-product, but this preference should not be construed to so limit the present invention. This light gasoline will have a high volatility and a high octane number and will be most advantageously used as a preferred blending component in the refinery gasoline pool. The use of such a high volatility blending component not only enhances the octane of the gasoline pool, but it has the further advantage of decreasing the amount of butane which .must be blended in the gasoline pool for volatility purposes, and thus makes such displaced blending butane available for increasing the LPG pool. The next fraction to be distilled in the final fractionation zone is an intermediate gasoline com'prised of the heavier hexane isomers and of the heptanes. Due to the chemical structure of these hydrocarbons they are not effectively dehydrocyclicized in the catalytic reforming zone but they may be effectively hydrocracked to selectively produce propane and butanes. This intermediate gasoline fraction is therefore recycled to the hydrocracking zone in the present invention. The last fraction produced is a heavy gasoline fraction containing octane and heavier hydrocarbons. This heavy gasoline fraction may be readily hydrocracked to produce propane and butane, but experience teaches that the octane and heavier hydrocarbons contained therein are most readily dehydrocyclicized. It is therefore the intent of the present invention to recycle the heavy gasoline fraction to the catalytic reforming zone.

The present process may be more clearly understood by reference to the accompanying drawings which illustrate various embodiments of the invention. Various pumps, heat exchangers, valves, control instruments, knockout pots and minor vessels, fractionator reflux and reboiler circulating lines, etc. have been eliminated or greatly reduced in order to clarify the drawing and thus implement the complete understanding of the present process. The utilization of these and other miscellaneous appurtenances will immediately be recognized by one skilled in the art of hydrocarbon processing, and it is not intended that such ommissions in the drawings or in the following discussion will unduly limit the present invention to the particular embodiments contained therein.

Referring now to FIGURE 1, wherein a simplified embodiment of the present invention is illustrated, the hydrocarbon charge stock is mixed in line 1' with hydrogen containing gas entering via line 14' and the resulting stream enters a first reaction zone 2' comprising catalytic reforming wherein substantial dehydrogenation and dehydrocyclization occurs to form an effluent rich in arcmatic hydrocarbons. Said effluent leaves the first reaction zone via line 3' and is combined with an eflluent from line 13', to be described hereinbelow. The combined stream enters a first separation zone 4' via line 3 wherein gas and liquid phases are separated into at least a hydrogen rich gas leaving via line 14 and an aromatics rich liquid leaving via line 5'. The hydrogen rich gas in line 14' is recycled to line -1' for further circulation through the first reaction zone 2' and at least a portion of the hydrogen which is produced in reaction zone 2' leaves line 14' via line 15'. The aromatics rich liquid in line 5' enters an aromatics separation zone 6- wherein substantially pure aromatic hydrocarbons are separated and recovered via line 16' while a substantially paraffinic hydrocarbon stream is discharged via line 7'. Said parafiinic stream is further combined with a light paraflinic hydrocarbon which enters line 7' via line 20', to be described hereinbelow, and with hydrogen entering via lines 17' and 18'. The hydrogen in line 17 is derived from line 15" and comprises at least a portion of the hydrogen rich gas which was produced by catalytic reforming in reaction zone 2', and the hydrogen in line 18 is derived from a second separation zone 10 to be described hereinbelow. The combined hydrogen and hydrocarbon stream in line 7' enters .a second reaction zone 8' wherein substantial hydrocracking occurs to produce hydrocarbons of lower molecular Weight. The efiluent from reaction zone 8' enters the second separation zone 10' wherein it is separated into at least a hydrogen containing gas, an LPG product which is recovered via line 21', a light gasoline comprising heavy isomers of hexane and isomers of heptane, and a heavy gasoline comprising octanes and heavier hydrocarbons. The hydrogen containing gas leaves via line 18' wherein it is separated into at least a portion which is recycled back to line 7' and a portion which is discharged via line 19' as the net gas make from the combination process. The light gasoline fraction leaves via line 20 and is recycled back to line 7' for further hydrocracking in reaction zone 8'. The heavy gasoline fraction leaves via line 1'1 wherein it is mixed with at least a portion of 1 1 the hydrogen from the first reaction zone 2 which enters via line .15 and the combined stream enters a third reaction zone '12 wherein dehydrocyclization of the hydrocarbon occurs. The resulting effluent leaves via line 13' and is recycled to line 3 for separation in zone 4'.

'Referring now to the more detailed FIGURE 2 of the drawings, there is shown a hydrocarbon charge stock entering the process system through line 1 where it is admixed with hydrogen which enters from line 9. To illustrate a particularly preferred embodiment of the present invention, the charge stock in line 1 is assumed to be a straight-11m naphtha contaminated by a substantial amount of sulfur and nitrogen but low in olefinic hydrocarbons. The charge stock is further characterized as having an initial boiling point of about 120 F. and an end boiling point of about 400 F., as determined by stadard ASTM distillation. Such a boiling range indicates that the charge stock contains hydrocarbons in the range of from about four carbon atoms per molecule to about twelve carbon atoms per molecule. The hydrocarbon and hydrogen mixture is charged into heater 2, leaves through line 3 at a temperature from about 500 F. to about 750 F., and is introduced into reactor 4. Reactor 4 normally operates at a pressure of from about 250 p.s.i.g. to about 600 p.s.i.g. on such a straight-run naphtha. Reactor 4 comprises at least one reactor vessel containing a hydrogen treating catalyst which saturates olefinic hydrocarbons and removes sulfur, nitrogen, and oxygen from the hydrocarbon molecules. Such catalyst normally is characterized as being comprised of silica, alumina, a nickel component, a cobalt component, and a molybdenum component. The hydrocarbon and hydrogen are then cooled in passing through line 5 and are introduced into separator 6 which serves to remove hydrogen and other gases, such as some hydrogen sulfide and ammonia, as Well as light hydrocarbon gases such as methane, ethane, and some propane. The gas stream leaves separator 6 via line 7 where it is admixed with a hydrogen rich stream which enters via line 34, to be defined below. This combined hydrogen rich gas is circulated back to line 1 via line 9 by means of recycle gas compressor 8. Because the hydrogen being supplied by line 34 is far in excess of what is being consumed in the hydrocarbon conversion which occurs in reactor 4, excess hydrogen rich gas is discharged through line 10 to compressor 11 where it is increased in pressure and sent via line 53 to the hydrocracking zone which will be described hereinafter.

The liquid hydrocarbon stream, which contains dissolved gases, leave-s separator 6 by way of line 12 and enters fractoinator 13 which is shown as a single column but may comprise more than one distillation column as required to obtain the desired separation. Essentially noncondensible gases, comprising hydrogen, hydrogen sulfide, ammonia, methane, and ethane are discharged through an upper line 14. This gas stream is normally further treated to remove and recover the elemental sulfur contained therein and the remaining gases may be used for further chemical processing or may be consumed as fuel gas. Because of trace amounts of water in the hydrocarbon stream leaving separator 6, provision is made for discharging such Water from column 13 through line 63. The hydrocarbon liquid is fractionated to effect a separation between pentanes and hexanes, and substantially all propane, butanes, and pentanes are removed as a common stream in line 62 for further processing as will be subsequently set forth.

The hexanes and heavier hydrocarbons leave fractionator 13 via line 15 and are mixed with hydrogen rich gas entering at line 31. In addition a further hydrocarbon stream obtained as hereinafter described may be introduced into line 15 via line 61. The combined stream enters heater 16 where it is raised to a temperature from about 900 F. to about 1050 F. and from which it is sent to reactor 16 via line 17. The etfiuent leaves reactor 18 by line 19 and enters heater 20 where the combined stream is again raised to a temperature from about 900 F. to about 1050" F. The stream is sent via line 21 into reactor 22 and is discharged therefrom by way of line 23. The combined hydrocarbon and hydrogen gas stream may be further combined in line 23 with a C and heavier hydrocarbon stream introduced by line 60 and obtained as will be subsequently described hereinafter. The combined stream enters heater 24 where it is again raised to a temperature from about 900 F. to about 1050" F., and it is then discharged via line 25 into reactor 26.

Reaction chambers 18, 22, and 26 are maintained at a pressure in the range of from about 200 p.s.i.g. to about 400 p.s.i.g. and contain reforming catalyst which promotes dehydrogenation and dehydrocyclization as well as some hydrocracking of the hydrocarbon molecules. Such catalysts are well known in the hydrocarbon processing art and are characterized as comprising a Group VIII noble metal and alumina. More specifically such catalysts are comprised of platinum on alumina substrate, or of platinum and halogen on alumina. Because dehydrogenation of naphthenic hydrocarbons to produce aromatic hydrocarbons is a highly endothermic reaction, the hydrocarbon and hydrogen mixed stream must be intermittently reheated in order to maintain the mixture at effective reaction temperatures. It is for this reason that the drawing indicates that the catalyst is provided in three reactors 18, 22, and 26 and that the process stream is reheated in heaters 20 and 24. Since the amount of endothermic reaction will vary in accordance with the concentration of naphthenes in the hydrocarbon charge stream, the number of reheating applications may vary. Thus, it is well known in the art of catalytic reforming of hydrocarbons, that charge stocks having naphthene contents in excess of 45 volume percent will require at least three reheatings to maintain adequate temperature on the catalyst and thus at least four reactors must be provided. It must therefor be realized that although the drawing indicates three reactors, the embodiment is not so as to be limited.

The final reformed efiiuent leaves reactor 26 through line 27 and upon cooling enters separator 28 where gas and liquid are separated. The gas comprising hydrogen, methane, ethane, and some propane leaves separator 28 by line 29 and is circulated by means of compressor 30 through line 31 back to line 15 for admixing with the hydrocarbon stream from fractionator 13. The excess hydrogen rich gas which is produced is discharged from the catalytic reformer through line 32 where compressor 33 increases its pressure to send it via line 34 into the catalytic hydrogen treater at line 7 as previously indicated. The liquid hydrocarbon which contains dissolved gases leaves separator 28 by way of line 36 and enters fractionator 37, which may of course be comprised of more than one distillation column where required to obtain the desired separation. Essentially non-condensible gases comprised of hydrogen, methane, and ethane are discharged through line 38. Such gases may be used for further chemical processing or may be consumed as a fuel. A light hydrocarbon stream comprised of propane, butanes, and pentanes leaves fractionator 37 via line 39' for further processing in the system, while a hydrocarbon stream comprising hydrocarbons containing from about six to about nine carbon atoms per molecule leaves via line 40 and a heavy reformate stream containing about nine or more carbon atoms per hydrocarbon molecule leaves via line 64.

Said hydrocarbon stream containing six to nine carbon atoms per molecule enters an aromatic hydrocarbon separation system 41 by Way of line 40. Such aromatic separation system may consist of a solvent extraction, or an adsorption, or an extractive distillation system together with any necessary auxiliary fractionators, separators, pumps, etc. One preferred method for recovering aromatic hydrocarbons is by solvent extraction using a glycolwater solvent or an equivalent solvent, but such preference must not be construed to so limit the present invention. This extraction process is well known in the hydrocarbon processing art. The recovered aromatic hydrocarbons are discharged through line 42, and they may subsequently be separated into their specific components in fractionator 65. Fractionator 65 is merely diagrammatic since it will normally be comprised of several distillation columns in order to effect the desired purity of separation and it is operated to result in the recovery of benzene via line 66, toluene via line 67, combined xylenes and ethylbenzene via line 68, and a heavy aromatics stream comprised of nine or more carbon atoms per molecule via line 69. A raffinate stream, consisting primarily of pa'raffinic hydrocarbons, but containing some naphthenes and a trace of aromatics, leaves the aromatic separation system 41 via line 43 wherein a hydrogen rich gasin introduced from line 52, the heavy reformate is introduced via line 64, the heavy aromatics stream is introduced via line 69, and a light paraflinic hydrocarbon stream, to be defined hereinbelow, is introduced from line 59. This combined stream is charged through line 43 and into heater 44 where it is heated to a temperature range of from about 750 F. to about 900 F. The heated hydrocarbon-gas mixture leaves heater 44 via line 45 and enters reactor 46 where hydrocracking of the paratlinic hydrocarbons occurs. The hydrocracking reaction may be undertaken at a pressure in the range of about 1000 p.s.i.g. to about 2000 p.s.i.g. but preferably to about 1500 p.s.i.g. Such reaction is promoted in the presence of a hydrocracking catalyst comprised of at least one metallic component selected from the metals of Group VI-B and VIII of the Periodic Table and mixtures thereof composited with silica-alumina substrate. More specifically the catalyst may be comprised of a Group VIII noble metal on a silica-alumina substrate and such Group VIII noble metal is preferably platinum, but such catalyst composition is in no way to be construed to so limit the present invention.

The resulting efiluent leaves reactor 46 via line 47 and upon cooling enters separator 48 wherein the gas and liquid phases are separated. The gas comprising hydrogen, methane, ethane and some propane is taken through line 49 by means of compressor 51 and discharged via line 52 into line 43 where it is mixed with the hydrocarbon stream being charged to the hydrocracking reaction system. Since hydrogen is consumed in reactor 46, make-up hydrogen is introduced into the system at line 52 from line 53 as the hydrogen rich excess gas from the hydrogen treating system which has been previously discussed. Because the gas supplied from line 53 is normally in excess of what is consumed in the hydrocracking reaction, provision is made for maintaining system pressure by discharging the excees gas from the system by way of line 50. The discharged gas which is rich in hydrogen and includes methane, ethane, and propane may be utilized in subsequent processing elsewhere or it may be consumed as a fuel gas.

The light hydrocarbon stream in line 62 is combined with the aforedescribed light hydrocarbon stream in line 39 and the resulting hydrocarbon mixture of propane, butane, and pentanes 'in line 39 is introduced into line 54 where it is combined with the liquid efil-uent leaving separator 48. The resulting hydrocarbon mixture in line 54 is introduced into fractionator system 55, which again may be comprised of more than one distillation column in order to effect the required hydrocarbon separations therein. The essentially nonacondensi-ble gases comprising hydrogen, methane, and ethane are discharged from the process through line 56. Such gas subsequently may be used in processing elsewhere or it may be consumed as a fuel. The desired LPG product comprised of propane and butanes is sent to storage via line 57. A light gasoline fraction comprised of pentanes and light hexanes, including 2,2-dimethylbutane and 2,3-dimethylbutane, is recovered as a by-product and sent to storage via line 58. Experience has shown that such parafiinic hydrocarbons cannot be dehydrocyclicized in a catalytic reformer, and that upon hydrocracking they produce excessive methane and ethane and are thus detrimental to hydrogen purity in the hydrocracking reactor 46. These hydrocarbons are therefor not recycled in the inventive process, but they may subsequently be utilized as a high octane gasoline blending component, or they may be used in other processing as in ethylene production, or they may be consumed as fuel. An intermediate gasoline fraction, comprised of the heavy hexanes and of the heptanes, leaves fractionator 55 via line 59. Experience has shown that these paraflinic hydrocarbons are not dehydrocyclicized in the catalytic reformer to produce aromatic hydrocarbons, but that they may be effectively subjected to hydrocracking to produce propane and butanes. Thus, in accordance with the present improved system, said intermediate gasoline fraction in line 59 is, therefor, recycled back to line 43 where it is mixed with hydrogen and said rafiiuate hydrocarbons for further hydrocracking. A heavy gasoline fraction, comprising octanes and heavier hydrocarbons, leaves fractionator 55 via line 60. Existing art teaches that .such hydrocarbons may be recycled to the catalytic hydrocracker for further cracking into propane, butanes, and pentanes. However, since it has been found that said heavy gasoline fraction, comprised substantially of paraflins, is readily dehydrocyclicized to form aromatic hydrocarbons when subjected to catalytic reforming, it is a feature of the present improved combined system to have said heavy gasoline fraction recycled back to the catalytic reformer via line 60.

In one mode of operation for the present invention, the heavy gasoline fraction may be charged through all reaction chambers by leaving line 60 by way of line 61 and entering line 15 wherein it is mixed with hydrogen containing gas from line 31 and the hexane and heavier hydrocarbons from fractionator 13 as discussed above. Because of the endothermicity of the first reactors due to the dehydrogenation of naphthenes to form aromatics, a temperature drop is experienced through the catalyst bed and the average catalyst temperature is not sufficient to effect substantial dehydrocyclization of paraifins. Upon reaching the last reactor, the naphthcnes will have been most fully dehydrogenated, little or no temperature drop will be experienced, the average catalyst temperature will be highest, and dehydrocyclization will effectively occur. The preferred embodiment of this invention is to, therefor, close-01f line 61 with sufficient valving 61' and to charge the heavy gasoline fraction from line 60 into line 23 and thus subject the recycled paraffinic hydrocarbons therein to primarily dehydrocyclization in the last reactor 26. In so doing, the required capacities of lines 15, 17', 19, and 21 and of heaters 16 and 20, and of reactors 18 and 22 are reduced and the corresponding capital expense is thereby reduced.

The present process is not restricted in that only high sulfur or high nitrogen hydrocarbon charge stocks may be refined therein. Where the hydrocarbon charge stock contains less than 10 ppm. by weight of sulfur and less than 1 ppm. by weight of nitrogen, as is the case for a gasoline or naphtha fraction which results from the hydrogen treating and hydrocracking of a heavy hydrocarbon to reduce molecular size, the hydrocarbon charge stream may be introduced into the process of the present invention via line 12, thus completely eliminating the hydrogen treating zone indicated by reactor 4. As an alternative operation, where such hydrocarbon charge stock is further characterized as having an initial boiling point of F. so that it primarily consists of hexane and heavier hydrocarbons, the charge may be introduced into the process of the present invention at line 15 thus elimimating the hydrocarbon component separation step in fractionator 13. Since the hydrogen treating zone has been eliminated, or merely by-passed, the excess hydrogen containing gas produced in the catalytic reforming zone leaves line 32 via line 35 and enters line 10 to be increased in pressure with compressor 11 for transfer via line 53 to the hydrocracking zone. Except for these modifications of process flow, the remaining steps for catalytic reforming, aromatic hydrocarbon separation, catalytic hydrocracking, fractionation, and recycling of the various gas and hydrocarbon liquid streams remains the same as detailed hereinbefore. Also, it should be noted that in operations Where the charge stocks will be low in naphthene content, such stocks may produce hydrogen in the catalytic reforming zone which is insuflicient in quantity to adequately supply the hydrogen consumption of the hydrocracking zone. In such instances, additional hydrogen must be supplied by an external hydrogen source and this additional hydrogen may be introduced into the process of the present invention at line 10 or line 52 or line 53.

The foregoing description clearly indicates the method of the present invention and the various benefits to be derived thereof. The combination process affords maximum production of LPG, maximum production of aromatics, and minimum production of gasoline while requiring little or no external source of hydrogen. By judicious adjustment of operating conditions within the process or by the withdrawal or the by-passing of at least a part of various internal streams it is possible to upgrade the gasoline or naphtha charge stock to the maximum profitability by varying the balance between the amount of each desired product which is produced.

I claim as my invention:

t1. Process for the production of aromatic hydrocarbons and liquefied petroleum gas which comprises;

'(a) reacting a substantially sulfur and nitrogen free hydrocarbon charge stock with hydrogen in a first reaction zone, said zone comprising a plurality of catalytic reactors maintained at dehydrogenation and dehydrocyclization conditions;

'(b) passing the total efiluent from said first reaction zone into a first separation zone under conditions sufficient to produce a hydrogen-rich gaseous stream, and a liquid stream containing aromatic and paraflinic hydrocarbons;

(c) subjecting at least a part of said liquid stream to solvent extraction to produce a substantially pure aromatic hydrocarbon product stream and a paradinrich raifinate stream;

(d) reacting said railinate stream in a second reaction zone with hydrogen at conditions selected to convert said raflinate stream into lower-boiling hydrocarbon products;

(e) separating the product effluent from said second reaction zone in a second separation zone in a manner 'suflicient to provide at least a hydrogen-rich gaseous phase, a first hydrocarbon fraction comprising principally hydrocarbons having from three .to four carbon atoms per molecule, a second hydrocarbon fraction comprising pentanes and relatively light hexanes, a third hydrocarbon fraction comprising relatively heavy hexanes and heptanes, and a fourth hydrocarbon fraction containing hydrocarbons having t least eight carbon atoms per molecule;

(f) returning said third hydrocarbon fraction to said second reaction zone for conversion therein into lower-boiling hydrocarbon products;

(g) returning said fourth hydrocarbon fraction to the last of said reactors in said first reaction zone; and,

(h) recovering said substantially pure aromatic hydrocarbon product stream, recovering said first hydrocarbon fraction, and recovering said second hydrocarbon fraction.

2. Process of claim 1, wherein the hydrogen-rich gaseous stream from said first reaction zone is supplied to said second reaction zone in an amount sufiicient to provide all of the hydrogen reacting in the second reaction zone.

3. Process of claim 1 wherein the total eflluent from said first reaction zone is separated to provide a nonaromatic hydrocarbon stream containing from three to five carbon toms per molecule and said non-aromatic stream is passed to said second separation zone with the liquid eflluent from said second reaction zone.

4. Process of claim 1 wherein the total efliuent from said first reaction zone is separated to provide a first portion comprising hydrocarbons having from three to five carbon atoms per molecule, a second portion comprising hydrocarbons having from about six to about eight carbon atoms per molecule, and a third portion comprising hydrocarbons having about nine or more carbon atoms per molecule, and that said second portion is solvent extracted to pl'OduCe said substantially pure aromatic hydrocarbon product stream while said third portion is charged to the second reaction zone.

5. Process of claim 1 wherein said recovered aroma-tic hydrocarbon product stream is further separated to provide a first aromatic fraction comprising benzene, a second aromatic fraction comprising toluene, at third aromatic fraction comprising xylenes and ethylbenzene and a fourth aromatic fraction comprising aromatics containing at least nine carbon atoms per molecule and said fourth aromatic fraction is charged to the second reaction zone in combination with said rafiinate stream.

6. Process of claim 1 wherein said second hydrocarbon fraction contains 2,2-dimethylbutane and 2,3-dimethylbutane.

References Cited UNITED STATES PATENTS 2,915,455 12/1959 Donaldson 208- 2,93 8,853 5/ 1960 Ammer et a1 208-6 5 3,265,510 8/1966 Lavergne et al. 208-59 DEIJBE'RT E. GANTZ, Primary Examiner. ABRAHAM RIMENS, Examiner.

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