US3787313A

US3787313A - Production of high-octane, unleaded motor fuel

Info

Publication number: US3787313A
Application number: US00308352A
Authority: US
Inventors: E Pollitzer
Original assignee: Universal Oil Products Co
Current assignee: Honeywell UOP LLC; Universal Oil Products Co
Priority date: 1972-11-21
Filing date: 1972-11-21
Publication date: 1974-01-22
Anticipated expiration: 1991-01-22

Abstract

A naphtha boiling range charge stock is converted into a motor fuel which does not necessitate the incorporation of metalcontaining additives otherwise required for suitable anti-knock characteristics. The process involves a combination of hydrocracking and catalytic reforming, and is effected in a manner which significantly decreases the quantity of methane and ethane produced. The novel form of hydrocracking results in a product predominantly comprising naphthenic hydrocarbons and highly branched paraffins, the latter being predominantly isobutane. Following separation to recover the isobutane, catalytic reforming is utilized to dehydrogenate the naphthenic compounds to produce an aromatic concentrate.

Description

ilnited States Patent [191 Poiiitzer Jan. 22, 1974 1 PRODUCTION OF HlGH-OfiTANE,

UNLEADED MOTOR FUEL Ernest 1L. Poiiitzer, Skokie, ll].

[73] Assignee: Universal Oil Products Company, Q aiitt nsszlllt [22] Filed: Nov. 21, 1972 [211 App]. No.2 308,352

[75} Inventor:

[52] US. Cl 208/60, 208/80, 208/93 [51] Int. Cl Clllg 39/00 [58] Field at Search 208/60, 80, 93

[56] References Cited UNITED STATES PATENTS 2,758,064 8/1956 Haensel 208/60 2,987,466 6/1961 Sanger et a 208/60 3,497,448 2/1970 Hamner et al. 208/60 Fracn'ana/ar I-Crack/ng Primary Examiner-Herbert Levine Attorney, Agent, or Firm.lames R. Hoatson et al.

[5 7] ABSTRACT A naphtha boiling range charge stock is converted into a motor fuel which does not necessitate the incorporation of metal-containing additives otherwise required for suitable anti-knock characteristics. The

process involves a combination of hydrocracking and catalytic reforming, and is effected in a manner which significantly decreases the quantity of methane and ethane produced. The novel form of hydrocracking results in a product predominantly comprising naphthenic hydrocarbons and highly branched paraffins,

the latter being predominantly isobutane. Following separation to recover the isobutane, catalytic reforming is utilized to dehydrogenate the naphthenic compounds to produce an aromatic concentrate.

10 Claims, 1 Drawing Figure Gala/yin: fie form in g 1 PAIENTEDJAN2-2m Q ESx w Q Q PRODUCTION OF IIIGII-OCTANE, UNLEADED MOTOR FUEL APPLICABILITY OF INVENTION The present invention involves a multiple-stage combination process for the conversion of naphtha, or gasoline boiling range hydrocarbons, to produce a highoctane motor fuel. The process is effected in a manner which produces an aromatic concentrate and isomeric paraffins, the latter being predominantly isobutane. More specifically, the inventive concept herein described encompasses an integrated refinery process for producing a high-octane, unleaded motor fuel gasoline pool.

Aromatic hydrocarbons, principally benzene, toluene, ethylbenzene and the xylene isomers, are required in large quantities to satisfy an ever-increasing demand for a wide variety of petrochemicals. For example, large quantities of benzene are hydrogenated to produce cyclohexane for use in the manufacture of nylon; toluene is often used as a solvent and as the starting material for various medicines, dyes, perfumes, etc. Possibly, the principal utilization of aromatic hydrocarbons resides in gasoline blending in view of their exceedingly high research octane blending values. For example,

benzene has a clear research octane value of about 99,

while toluene and the other alkyl-substituted aromatics have values exceeding lOO. lsobutane finds widespread use in organic synthesis, as a refrigerant and as an aerosol propellant, etc. Other uses include conversion to isobutenes for the subsequent production of butyl rubber, copolymer resins with butadiene, acrylonitrile, etc. In accordance with one embodiment of the present invention, the multiple-stage process is integrated into an overall refinery scheme for the production of a high-octane, unleaded gasoline pool. The aromatic concentrate is directly employed in the gasoline pool, while the isobutane is subjected to alkylation, with a suitable olefini'c hydrocarbon, the normally liquid alkylate product beingrecovered as a part of the gasoline pool.

Relatively recent investigations into the causes and cures of environmental pollution have shown that more than half of the violence perpetrated upon the atmosphere stems from vehicular exhaust, and consists primarily of unburned hydrocarbons and carbon monoxide. These investigations have brought about the development of catalytic converters which, when installed within the automotive exhaust system, are capable of converting more than 90.0 percent of the noxious components into innocuous materials. While developing these catalytic converters, it was learned that the efficiency of conversion and the stability of the selected catalytic composite were severely impaired when the exhaust fumes resulted from the combustion of leadcontaining motor fuel. When compared to the operation of the converter during the combustion of clear, unleaded gasoline, both the conversion of noxious components and the stability of the catalytic composites decreased as much as 50.0 percent. Therefore, it is rapidly being recognized throughout the petroleum industry, as well as in major gasoline-consuming countries, that suitable motor fuel must ultimately be produced for consumption in current internal combustion engines without requiring the addition of lead to increase the octane rating and thereby enhance the anti-knock properties. Also being recognized is the fact that unburned hydrocarbons and carbon monoxide are not the only extremely dangerous pollutants being discharged by way of vehicular exhaust. Japan has recently experienced an increase in the incidence of lead poisoning, and has enacted legislation which reduces significantly the quantity of lead permitted in motor fuel intended for consumption in that country.

A natural-flowing consequence of the removal of lead from motor fuel gasoline, in addition to others, resides in the fact that petroleum refining operations and techniques will necessarily undergo modification in order to produce voluminous quantities of a highoctane, unleaded motor fuel in an economically attractive fashion. One well-known and well-documented refining process, capable of significantly improving the octane rating of naphtha boiling range fractions, is the catalytic reforming process. In such a process, the principal octane-improving reactions are naphthene dehydrogenation, naphthene dehydroisomerization, paraffin, isomerization, dehydrocyclization and hydrocracking. Naphthene dehydrogenation is an extremely rapid reaction, and constitutes the principal octaneimproving reaction. With respect to a five-membered ring alkylnaphthene, it is necessary first to effect isomerization to produce a six-membered ring naphthene, followed by dehydrogenation to an aromatic hydrocarbon. Paraffin dehydrocyclization is achieved through the conversion of straight-chain paraffins having at least six carbon atoms per molecule. The degree of paraffin aromatization is limited by virtue of the fact that the aromatic concentration increases as the reactants traverse the reforming reaction zone. Unreacted, relatively low-octane paraffins are, therefore, present in the reformed product effluent and effectively reduce the overall octane rating thereof. In the past, these components could be tolerated as a result of their high susceptibility to the addition of a lead compound.

The present invention is in part founded upon recognition of the fact that the higher paraffins i.e. C,,-, C,,-, and C -paraffins, etc. are more easily converted to aromatic hydrocarbons, via dehydrocyclization, at relatively low operating severity levels, whereas the lower paraffins necessarily require a relatively higher severity level which is accompanied by the adverse effect attendant cracking. Additionally, while the novel hydrocracking system, forming an integral part of the present combination process, retains virtually 100.0 percent of the cyclic hydrocarbons in the feed stock, alkyl-substituted aromatic compounds can lose their alkyl groups. These two considerations contribute to an increase in the production of light gaseous hydrocarbons.

When operating at a relatively high severity, the paraffinic hydrocarbons within the reforming reaction zone are subjected to hydrocracking. Although this partially increases the octane rating of the normally liquid gasoline boiling range products, substantial quantities of gaseous material are produced. This light gaseous material is substantially completely saturated and comprises methane, ethane, propane and butane. As hereinafter set forth, the propane/butane concentrate may be recovered and utilized in an alkylation reaction system to produce a normally liquid, high-octane alkylation product. At a relatively low reforming severity, paraffin cracking is decreased with the result that an increased quantity of low octane rating saturates is produced. In order to upgrade the overall quality of the gasoline pool, either the addition of lead becomes necessary, or the low octane rating saturates must be subjected to further processing to produce higher octane components. As previously stated, further processing of the saturates can be eliminated by increasing the operating severity within the catalytic reforming reaction zone. This type of operation produces a two-fold effect, notwithstanding an increase in the octane rating of the final product, first, additional high-octane aromatic components are produced and, secondly, the lowoctane rating components are at least partially eliminated by conversion either to aromatic hydrocarbons, or to light normally gaseous material. Therefore, the results include a lower volumetric liquid yield of gasoline due both to shrinkage in molecular size when paraffins and naphthenes are converted to aromatics, and to the production of the aforesaid light gaseous components. In accordance with one overall refinery operation, into which the present invention is integrated, the hydrocracking system is dove-tailed with at least a separation system, a catalytic reforming unit and an alkylation unit. The end result is the production of a high-octane, unleaded gasoline pool, unaccompanied by substantial liquid yield loss.

Hydrocarbonaceous charge stocks, contemplated for conversion in accordance with the present invention, constitute naphtha boiling range hydrocarbon fractions and/or distillates. Gasoline boiling range hydrocarbons" generally connotes those hydrocarbons having an initial boiling point of at least about IF., and an end boiling point less than about 410F., and is inclusive of intermediate boiling range fractions often referred to in the art as light naphtha and heavy naphtha." It is not intended to limit the present invention to a charge stock having a particular boiling range. Suffice to say, a suitable charge stock will generally have an initial boiling point above about 100F. and an end boiling point below about 4l0F. In those instances where the reformed product effluent is subjected to further separation to recover a substantially pure aromatic concentrate, the paraffinic hydrocarbons may be recycled to the hydrocracking reaction zone. In this manner, the charge stock is reduced in boiling range, while maintaining the ring structure and eventual aromatic yields. The precise boiling range of any given naphtha charge stock will be primarily dependent upon the economic and processing considerations prevalent in the particular locale where the charge stock is available.

OBJECTS AND EMBODIMENTS The principal object of the present invention constitutes the simultaneous production of aromatic hydrocarbons and an isobutane concentrate.

Another objective resides in the production of a highoctane, unleaded motor'fuel gasoline pool. A corollary objective of my invention is to provide an integrated refinery operation for producing high liquid yields of a high-octane, unleaded gasoline pool.

Therefore, in one embodiment, the present invention involves a process for the simultaneous production of an aromatic concentrate and isobutane, from a naphtha boiling range charge stock, which process comprises the steps of: (a) separating said charge stock in 'a first separation zone, to provide a first fraction containing hydrocarbons having ten carbon atoms per molecule, and a second fraction containing hydrocarbons having less than about ten carbon atoms per molecule; (b) reacting said second fraction with hydrogen, in a hydrocracking reaction zone, at hydrocracking conditions and in contact with a hydrocracking catalytic composite of a Group VIII noble metal component and the reaction product of alumina and a gem-polyhalo compound; (c) separating the resulting hydrocracked product effluent, in a second separation zone, to provide a heptane-plus concentrate and to recover said isobutanes; (d) reacting said heptane-plus concentrate and said first fraction in a catalytic reforming zone, at reforming conditions selected to convert paraffinic and naphthenic hydrocarbons to aromatic hydrocarbons; and, (e) recovering said aromatic concentrate from the resulting reformed product effluent.

Other embodiments of my invention involve the use of various catalytic composites, operating conditions and processing techniques. In one such other embodiment, the reformed product effluent is separated, in a third separation zone to provide a saturated paraffinic stream and to recover the aromatic concentrate. In another such embodiment, the paraffinic stream recovered from the third separation zone is reacted with hydrogen in said hydrocracking reaction zone.

In a more limited embodiment, a pentane/hexane concentrate is separated from the hydrocracked product effluent and subsequently reacted with hydrogen in a hydroisomerization reaction zone to produce pentane and hexane isomers. Similarly, a C /C concentrate may be recovered from the reformed product effluent and subjected to hydroisomerization. In still another embodiment, at least a portion of isobutane concentrate is reacted with an olefinic hydrocarbon, in an alkylation reaction zone, to produce a normally liquid alkylate product.

SUMMARY OF INVENTION As hereinbefore set forth, the present invention involves a particular saturate cracking zone and a catalytic reforming zone. Additionally, in another embodiment, an integrated refinery scheme, incorporating the present inventive concept, utilizes a solvent extraction zone, an isomerization zone and an alkylation reaction zone. In a specific embodiment, the overall refinery process includes a dehydrogenation reaction zone to produce the olefinic hydrocarbons utilized in the alkylation reaction zone. In order that a clear understanding of the integrated refinery process be obtained, a brief description of each of the various reaction and separation zones, utilized in one or more embodiments, is believed to be warranted. In describing each individual zone, one or more references will be made to United States Patents in order that more details will be available where desired. Such references, are not to be construed as either exhaustive, or limiting, but merely exemplary and illustrative.

I have observed that, notwithstanding the preservation of the ring compounds while effecting the reactions in the hydrocracking reaction zone, alkyl groups will be removed from those ring compounds containing the same. These alkyl groups contribute to the net liquid yield loss in the form of light gaseous material; however, a more significant consideration is that the aromatic hydrocarbons in the ultimate product will contain fewer alkyl groups and possess, therefore, a lower clear octane blending value. For example, benzene has a clear blending value of about 99, ethylbenzene a value of about 124, while the various xylenes have values ranging from about 120 to about 150. These clear octane blending values, as well as those of other hydrocarbon species, are obtained from APl Research Project Number 45, Tabulated Knock-test Data to June 30, 1954," Table 1V, page 119. it is noted, therefore, that the removal of the alkyl groups from the ring compounds results in a lower clear octane blending value.

In accordance with the present invention, the charge stock is initially separated, for example in a fractionation system, to provide a first fraction containing hydrocarbons having carbon atoms per molecule, and a second fraction containing hydrocarbons having less than about ten carbon atoms per molecule. 1n a preferred mode of operation, the separation is effected such that substantially all of the C -ring compounds in the fresh feed charge stock are withdrawn with the heavier fraction and introduced into the catalytic reforming reaction zone. in those instances where the catalytically reformed product effluent is subjected to a solvent extraction technique to recover the atomatic concentrate, the paraffinic raftinate may be recycled to the hydrocracking reaction zone for conversion therein into valuable normally liquid isomeric hydrocarbons.

HYDROCRACKING REACTION ZONE The naphtha boiling range charge stock processed in accordance with the present invention, may be derived from a multitude of sources. For example, one source constitutes those naphtha distillates derived from full boiling range petroleum crude oils; another source is the naphtha fraction obtained from the catalytic cracking of gas oils and other, heavier hydrocarbon mixtures, while another source constitutes the gasoline boiling range effluent from a hydrocracking reaction zone which processes heavier-than-gasoline feed stocks. In view of the fact that the greater proportion of such naphtha fractions are contaminated through the inclusion of sulfurous and nitrogenous compounds, it is contemplated that the process may have integrated therein a hydrorefining zone, complete details of which are well known and thoroughly described within the prior art. It is understood that such pretreatment of the naphtha charge stock does not constitute an essential feature of the present combination process.

As hereinbefore set forth, one key feature of the present inventive concept resides in the use of a particular hydrocracking reaction zone. This hydrocracking reaction zone is unlike present-day hydrocracking processes both in function and end result. Initially, the charge to the hydrocracking reaction zone constitutes the lower boiling portion of the naphtha charge stock, and the product effluent contains very little, if any, methane and ethane. The propane in the product effluent can be recovered and subsequently utilized as LPG, or as feed to the allcylation reaction zone, or for the synthesis of isopropyl alcohol. Through the utilization of a particular catalytic composite and operating conditions, the cracking of paraffinic hydro-carbons in the charge stock produces relatively large quantities of butane, which is exceedingly rich in isobutane. Similarly, virgin pentanes and hexanes are converted to various isomers thereof. With respect to the overall octane rating of the gasoline pool, this aspect is of great significance. Whereas n-pentane has a blending octane rating of only 62, isomers thereof average about 95.9; similarly, while n-hexane has a clear octane blending value of about 20, the isomers thereof have an average rating of about 90.8.

in view of the unique character of the product effluent, being exceedingly rich in isoparaffins, the present hydrocracking reaction zone is herein referred to as 1- cracking. The selective nature of the hydrocracking reactions taking place include the retention of cyclic rings and a reduction in molecular weight of those rings via isomerization and the splitting of isobutane from the parent molecule. Thus, cyclic compounds boiling in the higher temperature range of the feed stock are converted to lower-boiling naphthenes which are, in turn, converted into gasoline boiling range aromatics in the subsequent catalytic reforming reaction zone. Benefits are thus afforded since the high octane aromatic hydrocarbons will be more uniformly distributed throughout the final gasoline boiling range product. With respect to increasing the yield of normally liquid hydrocarbons in the unleaded gasoline pool, the butane concentrate may be subjected to alkylation with suitable olefinic hydrocarbons.

The hydrocracking reaction conditions, under which the process is conducted, will vary according to the chemical and physical characteristics of the particular charge stock. l-llydrocracking reactions have heretofore generally been effected at pressures in the range of about 1,500 to about 5,000 psig., a liquid hourly space velocity of about 0.25 to about 5.0, hydrogen circulation rates of about 5,000 to about 50,000 scf./Bbl. and maximum catalyst bed temperatures in the range of about 7001F. to about 950F. As discussed in the prior art, heavier charge stocks require a relatively high severity of operation including high pressures, high catalyst bed temperatures and relatively low liquid hourly space velocities. A lower severity of operation is employed with comparatively lighter feed stocks such as the kerosenes and light gas oils. in accordance with the present invention, regardless of the characteristics of the naphtha boiling range charge stock, the hydrocracking process is effected at a relatively lower severity than is now commonly in use. The hydrocracking reaction zone has disposed therein a catalytic composite comprising a Group Vlll noble metal component and the reaction product of alumina and a gempolyhalo compound. The conversion conditions include a liquid hourly space velocity of about 0.5 to about 10.0, a hydrogen circulation rate of about 3,000 to about 20,000 scfJBbL, a pressure from about 200 to about 2,000 psig., and preferably less than about 1,000 psig. and, of greater significance, a maximum catalyst bed temperature from about 300F. to about 600F. In many instances, the operating pressure will consistently be in the range of about 200 to about 500 psig., the hydrogen concentration is from about 4,000 to about 10,000 scfJBbl. and the liquid hourly space velocity about 2.0 to about 10.0, without inducing serious effects either in regard to the effective life of the catalytic composite, or with respect to the desired product slate.

As hereinbefore set forth, the hydrocracking reaction zone has disposed therein a catalytic composite containing a Group Vlll noble metal component and the reaction product of alumina and a gem-polyhalo compound which may be a sublimed Friedel-Crafts metal halide. Thus, where the gem-polyhalo compound is, for example, aluminum chloride, the catalyst appears to be characterized by the following group:

Considering first the porous carrier material serving as the support for the catalytically active metallic components, it is preferred that it be absorptive and possess a high surface area from about 25 to about 500 square meters per gram. Heretofore, suitable carrier materials have been selected from the group of amorphous refractory inorganic oxides including alumina, titania, zirconia, silica, mixtures thereof, etc. When of the amorphous type, the preferred carrier material appears to be a composite of alumina and silica, with the latter being present in an amount of about 10.0 percent to about 90.0 percent by weight. Recent developments in the area of catalysis have further shown that various crystalline aluminosilicates can be employed to advantage in some hydrocracking situations. Such zeolitic material includes mordenite, faujasite, Type A or Type U molecular sieves, etc. In view of the fact that a gempolyhalo compound does not appear sufficiently strong to react with silica, to form the type of group previously described, the preferred carrier material constitutes alumina. While the action and effect on refractory material other than alumina and silica, for example zirconia, is not known with accuracy, it is believed that reaction does not take place to a degree sufficieut to produce the desired catalyst and result.

The hydrocracking catalytic composite contains a Group VIII noble metal component. Suitable metals are those from the group including platinum, palladium, rhodium, ruthenium, osmium and iridium, with platinum and palladium being particularly preferred. These metal components, for example platinum, may exist within the final composite as a compound such as an oxide, sulfide, halide, etc., or in an elemental state, the latter being preferred. Generally, the amount of the noble metal component is small compared to the quantities of the other components combined therewith. On an elemental basis, the noble metal component comprises from about 0.01 percent to about 2.0 percent by weight of the final catalytic composite. Bi-metallic catalysts, containing germanium, rhenium, or tin, in addition to the noble metal, are also suitable for use in the hydrocracking reaction zone.

The catalytically active metallic components may be incorporated within the catalytic composite in any suitable manner including co-precipitation or cogellation with the carrier material, ion-exchange or impregnation. The latter constitutes a preferred method of preparation, and utilizes water-soluble compounds of the metallic components. Thus, a platinum component may be added to a carrier material by commingling the latter with an aqueous solution of chloroplatinic acid. Other water-soluble compounds may be employed, and include ammonium chloroplatinate, platinum chloride, chloropalladic acid, palladic chloride, etc. Following impregnation, the carrier material is dried and subjected to a calcination, or oxidation technique which is generally followed by reduction in hydrogen at some elevated temperature.

The method of incorporating a Freidel-Crafts metal halide, as the gem-polyhalo compound, involves a sublimation, or vaporization technique, with the vaporized metal halide contacting the alumina containing the Group VIII noble metal component or the nickel component. That is, the catalytically active metallic component is composited with the alumina prior to contact with the sublimed metal halide. Briefly, therefore, the preferred technique involves the incorporation of the Friedel-Crafts metal halide after the catalytically active metal components have been impregnated onto the carrier material, and following drying, calcination and reduction in hydrogen. The metal halide is vaporized onto the carrier and heated to a temperature of about 300C. for a time sufficient to remove any unreacted metal halide. Thus, the final catalytic composite does not contain any free, uncombined Friedel-Crafts metal halide. Following vaporization of the metal halide and heating of the thus-formed composite, the refractory inorganic oxide will be increased in weight by from about 2.0 percent to about 25 .0 percent based upon the original weight of the carrier material. While the exact increased weight does not appear to be critical, high activity catalysts are obtained when the thus-treated refractory material has a weight increase of about 5.0 percent to about 20.0 percent. Further details of this sublimation technique may be found in US. Pat. No. 2,924,628 (Class 260666). Since the desired group, as hereinbefore set forth, is sensitive to moisture, the sublimation technique is effected after the Group Vlll noble metal component has been combined with the alumina, and the composite reduced in hydrogen.

Various Friedel-Crafts metal halides may be utilized in the hydrocracking catalytic composite, but not necessarily with equivalent results. Examples of such metal halides include aluminum bromide, aluminum chloride, antimony pentachloride, beryllium chloride, germanium petrachloride, ferric chloride, ferric bromide, gallium trichloride, stannic bromide, stannic chloride, titanium tetrabromide, titanium tetrachloride, zinc bromide, zinc chloride, and zirconium chloride. The Friedel-Crafts aluminum halides are preferred, with aluminum chloride and/or aluminum fluoride being particularly preferred. This is due to the ease of preparation and the fact that the thus-prepared catalysts have an unexpectedly high activity for the selective production of isoparaffins, and particularly for isobutane.

Temperatures at which the Friedel-Crafts metal halide is vaporized onto the alumina will vary in accordance with the particular metal halide utilized. Vaporization is carried out either at the boiling, or sublimation point of the particular Friedel-Crafts metal halide, or at some temperature not substantially exceeding these points; for example, not more than about C. higher than the boiling point or sublimation point. In effecting one catalyst preparation technique, the amorphous carrier material has aluminum chloride sublimed thereupon. Aluminum chloride sublimes at a temperature of about 178C. and a suitable vaporization temperature will, therefore, range from about 180C. to about 275C. The sublimation technique may be carried out under pressure and in the presence of a diluent such as an inert gas. Although the particularly preferred technique involves the sublimation of a metal halide directly to react with the alumina, the reaction product may result from other gem-polyhalo compounds which react with alumina to form the type of group previously described.

Thus, the catalyst may comprise the reaction product of an alumina-Group Vlll noble metal composite and one or more gem-dihalo or gem-polyhalo compounds. The interaction produces catalytic sites of higher acidity than can be produced, for example, by treatment with hydrochloric acid. Such compounds include car bon tetrachloride, CI-lCl sulfur dichloride, sulfur oxychloride, PCl POCl R--Cl-I-Cl etc.

Prior to its use, the catalytic composite is subjected to a substantially waterfree reduction technique. This is designed to insure a more uniform and thorough distribution of the metallic components throughout the carrier material. Substantially pure and dry hydrogen is employed as the reducing agent at a temperature of about 800F. to about l,200F., and for a time sufficient to reduce the metallic components.

In view of the fact that the reactions being effected are exothermic in nature, an increasing temperature gradient is experienced as the reactants traverse the hydrocracking catalyst bed. In accordance with the present process, the maximum catalyst bed temperature is maintained in the range of about 300F. to about 600F. In order to assure that the maximum catalyst bed temperature does not exceed the allowable limit, the use of conventional quench streams, either normally liquid, or normally gaseous, introduced at one or more intermediate loci of the catalyst bed is contemplated.

As hereinbefore set forth, the product effluent from the hydrocraclcing reaction zone, considering only the normally gaseous portion thereof, is predominantly butane, the greater proportion of whichconstitutes isobutane. Similarly, the pentane-hexane concentrate is rich in isomers of higher octane rating; for this reason, the hydrocraclcing reaction zone is herein referred to as I- cracking," the I alluding to isomer production. In addition to the production of exceedingly large quantities of isobutane, accompanied by little yield loss to methane and ethane, an unusual and unexpected result is the virtually complete retention of cyclic hydrocarbons originally present in the fresh feed charge stock. Furthermore, those heavier cyclic hydrocarbons in the fresh feed have been reduced in molecular weight such that the subsequent reformed product effluent exhibits a more uniform distribution of the high octane aromatic components. This achieves importance from the standpoint of the possible lowering of the end boiling point of motor fuel at some future date.

The foregoing is evidenced by results which were obtained when a Mid-continent, straight-run naphtha fraction was subjected to I-cracking. This charge stock has a gravity of 55.0 APl, an initial boiling point of about 210F. and an end boiling point of about 369F. Virgin cyclic hydrocarbons in the charge stock constituted 52.6 percent by weight. The l-cracking yields, being based on weight percent of the naphtha charge stock were as follows: butanes-minus, 21.2 percent by weight; pentanes, 11.7 percent; hexanes, 11.9 percent; and, heptane-plus hydrocarbons, 56.5 percent, which values are inclusive of a hydrogen consumption in the amount of l .3 percent by weight of the naphtha charge stock. From the standpoint of selectivity, butanes were produced in an amount of 92.0 percent by weight of the total butane-minus portion, the isobutane content of the total butanes being 92.0 percent by weight; of the total pentanes produced, 89.0 percent by weight were isomeric in nature. On a molar basis, the cyclic retention amounted to 99.0 percent.

As previously set forth, the selective l-cracking operation also has an effect on the boiling range of the aromatics produced in the subsequent catalytic reforming step. When a comparison is made between the product produced by direct catalytic reforming of the Midcontinent charge stock, and the product resulting from I-cracking the naphtha followed by catalytically reforming the heptane-plus portion of the hydrocracked product, it is noted that the last 50.0 percent by volume indicates a lower boiling range to the extent that there is a 40F. difference in the end boiling point; the end boiling point is, in fact, lower than that of the original feed stock. Catalytic reforming of itself results in a product having a gravity of 43.2 API, an initial boiling point of 144F. and an end boiling point of 404F., having a clear octane rating of 96.3 and containing 61.8 percent by volume of aromatic hydrocarbons. Where the naphtha charge stock is initially subjected to I- cracking, followed by the catalytic reforming (at the same reforming conditions) of the heptane-plus portion, the product has a gravity, API of 36.5, an initial boiling point of 162F. and an end boiling point of 35 8F., and possesses a clear octane rating of 105.4 and contains 84.5 percent by volume of aromatic hydrocarbons. In this particular comparison, the lower-boiling front of the product resulting from the combination cannot be compared directly since it does not include the pentane/hexane portion of the hydrocracked product effluent. The overall yields, considering a butaneplus product in an amount of 101 .2 percent by volume, based upon the naphtha charge stock, having a clear research octane rating of 98.9, is produced with an accompanying yield loss to methane, ethane and propane of 4.0 percent by weight.

One of the principal objects of the present inventive concept is to afford a method for achieving a distinct improvement in the foregoing results. Although the degree of cyclic hydrocarbon retention is approximately the same, and the end product consists of a similar quantity of aromatic hydrocarbons, the octane blending values are increased as a result of a greater degree of retention of alkyl groups on the alkyl-substituted ring compounds. Furthermore, there exists a lower yield loss resulting principally from a decreased production of methane and ethane.

CATALYTIC REFORMING ZONE The product effluent from the I-cracking reaction zone is separated, in suitable fractionation facilities, into various component streams. A butane concentrate, consisting predominantly of isobutane, is recovered and subjected to either alkylation, or dehydrogenation as hereinafter set forth. A pentane/hexane concentrate, rich in isomers thereof, is separately recovered and may be introduced directly into the unleaded gasoline pool. In a preferred embodiment, normal pentane and normal hexane are separately recovered and subjected to isomerization to produce additional pentane/hexane isomers. The heptane-plus portion of the I-craclted product effluent constitutes a portion of the charge of the catalytic reforming reaction zone.

Catalytic composites, suitable for utilization in the reforming reaction zone, generally comprise a refractory inorganic oxide carrier material containing a metallic component selected from the noble metals of Group VIll. Recent developments in the area of catalytic reforming have indicated that activity and stability are significantly enhanced through the addition of various catalytic modifiers, especially tin, rhenium, nickel and/or germanium, thereby forming a bi-metallic catalyst. Suitable porous carrier materials include refractory inorganic oxides such as alumina, silica, zirconia,

etc. Generally favored metallic components include ruthenium, rhodium, palladium, osmium, rhenium, platinum, iridium, germanium, nickel and tin, and mixtures thereof. These metallic components are employed in concentrations ranging from about 0.01 percent to about 5.0 percent by weight, and preferably from about 0.01 percent to about 2.0 percent by weight. Reforming catalyst may also contain combined halogen selected from the group of chlorine, fluorine, bromine, iodine and mixtures thereof, with chlorine and fluorine being particularly preferred.

Since the reforming reaction zone processes substantially only a heptane-plus product, the conditions required result in a lower degree of operating severity. Operating severity level increases with increased temperature and lower liquid hourly space veolicty. Those skilled in the reforming art will immediately recognize that a lower severity operation is more advantageous from the standpoint of catalyst life and decreased production of methane and ethane. Furthermore, at a higher space velocity, the refiner enjoys the added advantage of greater throughput per unit of time.

Illustrative catalytic reforming processes are found in US. Pat. Nos. 2,905,620 (Class 208-65), 3,000,812 (Class 208-138) and 3,296,118 (Class 208-100). Effective prior art reforming operating conditions include a catalyst temperature within the range of about 800F. to about 1,100F., a liquid hourly space velocity about 1.0 to about 5.0 and a pressure of about 500 to about 1,000 psig. The quantity of hydrogen-rich gas, in admixture with the hydrocarbon charge stock, is generally in the range of about 1.0 to about 20.0 moles of hydrogen per mole of hydrocarbon. In accordance with the present combination process, the catalytic reforming reaction zone will normally function at a temperature in the range of about 800F. to about 950F., a liquid hourly space velocity of about 2.0 to about 10.0, or more, and a pressure in the range of about 100 to about 750 psig. The hydrogen concentration will generally be in the same range as that of the prior art. The product effluent from the catalytic reforming reaction zone is generally introduced into a high-pressure separation zone at a temperature in the range of about 60F. to about 140F., to separate lighter components from heavier, normally liquid components. Since normal reforming operations are hydrogen-producing, a certain amount of a hydrogen-rich stream is generally removed from the reforming system by way of pressure control. It is within the scope of the present invention that such excess hydrogen be employed in the hydrogenconsuming hydrocracking reaction zone, as make-up hydrogen, as well as in the hydroisomerization reaction zone.

AROMATIC SEPARATION ZONE As hereinabove set forth, the catalytic reforming reaction zone is maintained at relatively low severity operating conditions in order to produce a product effluent rich in aromatic hydrocarbons, and for the purpose of dehydrocyclization of the paraffinic material in the charge stock. Although any aromatic separation scheme, such as fractionation, may be utilized, a greater degree of efficiency is achieved through the use of a solvent extraction system. Solvent extraction, to produce an aromatic concentrate and a paraffinic rafflnate, is a well known technique thoroughly described in the literature. Suitable techniques involve the operations illustrated in US. Pat. Nos. 2,730,558 (Class 260-674) and 3,361,664 (Class 208-313). Solvent extraction processes utilize a solvent which possesses a greater selectivity and solvency for aromatic components in the reformed product effluent than for the paraffinic components. Selective solvents may be selected from a wide variety of normally liquid organic compounds of generally polar character; that is, compounds containing a polar radical. In any given situation, the particular solvent is one which boils at a temperature above the boiling point of the hydrocarbon mixture at an ambient extraction pressure. Illustrative specific organic compounds, useful as selective solvents in extraction processes for the recovery of aromatic hydrocarbons, include the alcohols, such as glycols, including ethylene glycol, propylene glycol, butylene glycol, tetra-ethylene glycol, glycerol, diethylene glycol, dipropylene glycol, dimethylether of ethylene glycol, triethylene glycol, tripropylene glycol, etc.; other organic solvents well known in the art, for extraction of hydrocarbon components from mixtures thereof with other hydrocarbons, may be employed. A particularly preferred class of such other solvents are those characterized as the sulfolane-type. Thus, as indicated in US. Pat. No. 3,470,087 (Class 208-321), a particularly preferred solvent is one having a five-membered ring, one atom of which is sulfur, the other four being carbon and having two oxygen atoms bonded to the sulfur atom. In addition to sulfolane, the preferred class include the sulfolenes, such as 2-sulfolene and 3- sulfolene.

Selectivity of the foregoing described solvents may be enhanced further through the addition of water. This increases the selectivity of the solvent phase for aromatic hydrocarbons over non-aromatic hydrocarbons. As a general practice, the solvent composition will contain from about 0.5 percent to about 20.0 percent by weight of water, and preferably from about 2.0 percent to about 15.0 percent by weight, principally depending on the particular solvent and the process conditions under which the extraction, extractive distillation and solvent recovery zones are functioning.

In general, solvent extraction is effected at elevated temperatures and pressures which are selected to maintain the charge stock and solvent in a liquid phase. Suitable temperatures are within the range of about F. to about 400F., and preferably from about 150F. to about 300F. Operating pressures include superatmospheric pressures up to about 400 psig. and preferably from about 15.0 psig. to about 150 psig. Extractive distillation zone pressure at the top of the distillation zone will generally be maintained in the range of about 1.0 psig. to about 20 psig. The reboiler temperature is dependent upon the composition of the feed stock and the selected solvents, although temperatures from about 275F. to about 360F. appear to yield more satisfactory results. The solvent recovery system is operated at relatively low pressure and sufficiently high temperatures to drive the aromatic hydrocarbons overhead, thus producing a lean solvent bottoms stream. Preferably, the top of the solvent recovery zone is maintained at a pressure from about to about 400 mm. Hg., absolute. These low pressures must be utilized since the reboiler temperatures should be maintained below about 370F. in order to avoid thermal decomposition of the organic solvent.

ALKYLATION REACTION ZONE Since the preferred use of the present inventive concept constitutes the integration thereof into an overall refinery scheme for the production of voluminous quantities of a high-octane, unleaded motor fuel gasoline pool, the isobutane-rich, butane concentrate from the I-cracking zone is utilized as fresh feed charge stock to an alkylation reaction zone. Alkylation is effected by intimately commingling the isobutane feed, an olefinic hydrocarbon and a particular catalyst as hereinafter described. It is understood that the source of the olefinic hydrocarbon, for utilization in the alkylation reaction zone, is not essential to the process encompassed by the present invention. Thus, outside olefinic material may be supplied from any suitable source such as a fluid catalytic cracking unit, or a thermal cracking unit. However, as indicated in another specific embodiment, at least a portion of the isobutane concentrate may be subjected to dehydrogenation in a dehydrogenation reaction zone to produce the alkylatable olefinic hydrocarbons. In still another embodiment, the propane produced within the process may also be dehydrogenated and introduced into the alkylation reaction zone.

The alkylation reaction zone may be any acidic catalyst reaction system such as a hydrogen fluoride, or sulfuric acid catalyzed system, or one which utilizes a boron halide in a fixed-bed reaction system. Hydrogen fluoride alkylation is particularly preferred, and may be conducted substantially as set forth in U.S. Pat. No. 3,249,650 (Class 260-68348). Briefly, the alkylation conditions, when effected in the presence of hydrogen fluoride catalyst, are such that the catalyst to hydrocarbon volume ratio within the alkylation reaction zone is in the range of about 0.5 to about 2.5. Ordinarily, anhydrous hydrogen fluoride will be charged to the alkylation system as fresh catalyst; however, it is possible to utilize hydrogen fluoride containing as much as about 10.0 percent water, although excessive dilution with water is generally avoided since it tends to reduce the alkylating activity of the catalyst and further introduces a variety of corrosion problems into the process. In order to reduce the prevailing disposition of the olefinic portion of the charge stock to undergo polymerization prior to alkylation, the molar proportion of isoparaffin to olefinic hydrocarbon within the alkylation reaction zone is desirably maintained at a value greater than 1.0, and preferably from about 3.0 to about 15.0. Alkylation reaction conditions, when catalyzed by hydrogen fluoride, include a temperature from to about l50F. and preferably from about 30F. to about 100F. The pressure maintained within the alkylation system is ordinarily at a level sufficient to maintain the hydrocarbons and catalyst in substantially liquid phase; that is, from about atmospheric to about 40 atmospheres. The contact time within the alkylation reaction zone is conveniently expressed in terms of spacetime, and is generally defined as the volume of catalyst within the contact zone divided by the volume rate per minute of hydrocarbon reactants charged to the zone. Usually, the space-time factor will be less than 30 minutes and preferably less than about 15 minutes.

The alkylation reaction zone effluent is separated to provide an acid phase and a hydrocarbon phase, the latter being separated to recover the normally liquid alkylate product and unreacted isobutane. The alkylate product, in combination with the aromatic concentrate from the solvent extraction zone forms part of the unleaded gasoline pool. Unreacted isobutane may be recycled to the alkylation reaction zone, or a portion thereof may be diverted to the dehydrogenation reaction zone for the purpose of producing additional olefinic hydrocarbons for utilization in the alkylation reaction zone.

ISOMERIZATION REACTION ZONE The n-pentane and n-hexane separated from the product effluent of the I-cracking reaction zone possess clear research blending values of about 62 and 25 respectively. These components are not, therefore, desirable in a gasoline pool which is intended to be free from lead additives. Therefore, in still another embodiment of the present invention, a normal pentane/normal hexane stream is separately recovered and introduced into an isomerization reaction zone for the purpose of producing an effluent product rich in pentane and hexane isomers. Since the selectivity of conversion in the isomerization reaction zone is virtually 100.0 percent, the unleaded gasoline pool can be significantly increased in its clear research octane rating without incurring a detrimental volumetric yield loss.

As indicated in U.S. Pat. No. 3,131,235 (Class 260683.3), the isomerization process is effected in a fixed-bed system utilizing a catalytic composite of a refractory inorganic oxide carrier material, a Group VIII noble metal component and combined halogen, preferably selected from fluorine and chlorine. The refractory inorganic oxide carrier material may be selected from the group including alumina, silica, titania, zirconia, mixtures of two or more, and various naturally occurring refractory material. Of these, a synthetically prepared gamma-alumina is preferred. The Group VIII noble metal is generally present in an amount of about 0.0l percent to about 2.0 percent by weight, and may be one or more metals selected from the group of ruthenium, rhodium, osmium, iridium, and particularly platinum and/or palladium. The amount of combined halogen will be varied from about 0.01 percent to about 8.0 percent by weight. Both fluorine and chlorine may be used to supply the combined halogen, although the use only of fluorine in an amount of about 2.5 percent to about 5.0 percent by weight is preferred.

The isomerization reactions are preferably effected in a hydrogen atmosphere utilizing sufficient hydrogen so that the hydrogen to hydrocarbon mole ratio to the reaction zone will be within the range of about 0.25 to about 10.0. Operating conditions will additionally include temperatures ranging from about 200F. to about 600F., although temperatures within the more limited range of about 230F. to about 320F. will generally be employed. The pressure, under which the reaction zone is maintained, will range from about 400 to about 1,000 psig., and the liquid hourly space velocity from 1.0 to about 3.0. Hydrogen is separated from the reaction products and recycled, while the normally liquid effluent is subjected to fractionation and separation to produce the desired isomerized product. Recovered starting material is also recycled to the reaction zone to increase the overall process yield. Another suitable isomerization process is that described in U.S. Pat. No. 2,924,628 (Class 260666).

DEHYDROGENATION REACTION ZONE As previously stated, at least a portion of the recovered butane concentrate may be subjected to dehydrogenation to produce the olefin required for alkylation within the alkylation reaction zone. In another embodiment, the propane stream is also subjected to dehydrogenation to provide additional olefins. The advisability of the utilization of either, or both techniques will be primarily dependent upon the availability of outside olefins; for example, from a catalytic or thermal cracking unit.

When dehydrogenation is deemed desirable, it may be effected essentially as set forth in U.S. Pat. No. 3,293,219 (Class 260683). Briefly, dehydrogenation re-actions are generally effected at conditions including a temperature in the range of from 400C. to about 700C., a pressure from about atmospheric to about 100 psig., a liquid hourly space velocity within the range of about 1.0 to about 40.0 and in the presence of hydrogen in an amount to result in a mole ratio of from 10:10 to 10011.0, based upon the paraffmic charge.

The dehydrogenation catalyst is a composite of an inorganic oxide carrier material, an alkali metal component, a Group Vlll metal component and a catalytic modifier from the group consisting or arsenic, antimony, bismuth, rhenium, germanium and tin. A particularly preferred catalyst comprises lithiated alumina containing from 0.05 percent to about 5.0 percent by weight of a Group VIII noble metal, particularly platinum. The catalytic modifier is employed in an amount based upon the concentration of Group Vlll noble metal components. For example, arsenic is present in an atomic ratio of arsenic to platinum in the range of about 0.20 to about 0.45. Although lithium is the preferred alkalinous metal component, the catalyst may contain calcium, magnesium, strontium, cesium, rubidium, potassium, sodium and mixtures thereof, etc.

Dehydrogenation conditions and catalysts result in a relatively low equilibrium conversion per pass, accompanied by relatively high selectivity to the desired olefinic hydrocarbons. Thus, while the conversion per pass might range from about 10.0 percent to about 35.0 percent, the selectivity of conversion will range from about 93.0 percent to about 97.0 percent, or higher. in view of the fact that the alkylation reactions are effected with a molar excess of paraffms over olefinic hydrocarbons, the high selectivity and relatively low conversion in the dehydrogenation zone are advantageous.

DESCRlPTlON OF DRAWING The inventive concept, encompassed by the present process, is illustrated in the accompanying drawing. The illustration is presented by way of a simplified block-type flow diagram representing the l-cracking and catalytic reforming reaction zones. Also illustrated are two fractionating columns utilized in the separation of the fresh feed charge stock and the product effluent from the l-cracking reaction zone. Miscellaneous appurtenances, not believed necessary for a completely clear understanding of the present process, have been eliminated. The utilization of details such as pumps, compressors, instrumentation and controls, heatrecovery circuits, miscellaneous valving, startup lines and similar hardware, etc., is well within the purview of those skilled in the petroleum refining art. Similarly, 6

with respect to the flow of materials throughout the system, only those major streams required to illustrate the interconnection and interaction of the various zones are presented. Thus, various recycle lines and vent-gas streams have also been eliminated.

With reference now to the drawing, it will be described by way of processing the Mid-continent naphtha charge stock previously described. The charge stock is introduced into the process by way of line 1 and is separated, in fractionator 2, to provide a light fraction containing C -paraffinic hydrocarbons in line 3, which fraction is substantially free from the isomeric xylenes. The catalytic composite, disposed in lcracking reaction zone 4, constitutes alumina, 0.75 percent by weight of palladium and 7.5 percent by weight of aluminum chloride, sublimed thereon to react with the alumina to form the type of group hereinbefore described. Operating conditions include a pressure of about 450 psig., a maximum catalyst bed temperature of about 375F., a liquid hourly space velocity of about 2.0 and a hydrogen to hydrocarbon mole ratio of about 4.0:l.0.

The product effluent is withdrawn from l-cracking zone 4 by way of line 5, and introduced therethrough into fractionator 6. Butanes and lighter components are withdrawn through line 7 while a pentane/hexane concentrate is withdrawn by way of line 8. A heptane-plus, normally liquid material is removed by way of line 9, combined with the heavy fraction in line 10 (from fractionator 2), the mixture continuing through line 10 into catalytic reforming zone 11. The reforming reactions are effected at a pressure of about 250 psig., a temperature of about 950F., a liquid hourly space velocity of about 2.0 and a hydrogen/hydrocarbon mole ratio of about :10. The catalytic composite constitutes an alumina carrier material containing 0.55 percent by weight of platinum, 0.20 percent by weight of rhenium and 0.87 percent by weight of combined chloride, all of which are computed on an elemental basis. The reformed product effluent is withdrawn from catalytic reforming zone 11 by way of line 12.

The overall yields, considering the hydrocracking and catalytic reforming reaction zones, indicates a butane-plus product in an amount of about 102.5 vol. percent, based upon the naphtha charge stock. This product stream has a clear research octane rating of about 99.5, and is produced with an accompanying yield loss to methane, ethane and propane of only about 2.5 percent by weight. These results are substantially improved over those obtained via the combination process wherein the fresh feed charge stock iS not initially separated in accordance with the method of the present invention.

The foregoing demonstrates the method by which the present invention is effected and the benefits afforded through the utilization thereof.

I claim as my invention:

1. A process for the simultaneous production of an aromatic concentrate and isobutane, from a naphtha boiling range charge stock, which process comprises the steps of:

a. separating said charge stock, in a first separation zone, to provide a first fraction containing hydrocarbons having ten carbon atoms per molecule, and a second fraction containing hydrocarbons having less than about 10 carbon atoms per molecule;

b. reacting said second fraction with hydrogen, in a hydrocracking reaction zone, at hydrocracking conditions and in contact with a hydrocracking catalytic composite of a Group Vlll noble metal component and the reaction product of alumina and a gem-polyhalo compound;

0. separating the resulting hydrocracked product effluent, in a second separation zone, to provide a heptane-plus concentrate and to recover said isobutane;

d. reacting said heptane-plus concentrate and said first fraction in a catalytic reforming zone, at reforming conditions selected to convert naphthenic hydrocarbons to aromatic hydrocarbons; and,

e. recovering said aromatic concentrate from the resulting reformed product effluent.

2. The process of claim 1 further characterized in that said reformed product effluent is separated, in a third separation zone to provide a saturated paraffinic stream and to recover said aromatic concentrate.

3. The process of claim 2 further characterized in that said paraffinic stream is reacted with hydrogen in said hydrocracking reaction zone.

4. The process of claim 1 further characterized in that said hydrocracked product effluent is separated to provide a pentane/hexane concentrate and a propane/- butane concentrate.

5. The process of claim 4 further characterized in that said pentane/hexane concentrate is reacted with hydrogen in a hydroisomerization reaction zone, at isomerization conditions selected to produce pentane and hexane isomers.

6. The process of claim ll further characterized in that said hydrocracking catalytic composite comprises from about 0.01 percent to about 2.0 percent by weight of a platinum or palladium component.

7. The process of claim ll further characterized in that said gem-polyhalo compound is a Friedel-Crafts metal halide.

8. The process of claim 1 further characterized in that said hydrocracking conditions include a maximum catalyst bed temperature of from 300F. to about 600F. and a pressure in the range of about 200 to about 500 psig.

9. The process of claim 1 further characterized in that said isobutane is reacted with an olefinic hydrocarbon, in an alkylation reaction zone, at alkylation conditions selected to produce a normally liquid alkylated hydrocarbon stream.

10. The process of claim 2 further characterized in that said third separation zone is a solvent extraction zone.

Claims

2. The process of claim 1 further characterized in that said reformed product effluent is separated, in a third separation zone to provide a saturated paraffinic stream and to recover said aromatic concentrate.
3. The process of claim 2 further characterized in that said paraffinic stream is reacted with hydrogen in said hydrocracking reaction zone.
4. The process of claim 1 further characterized in that said hydrocracked product effluent is separated to provide a pentane/hexane concentrate and a propane/butane concentrate.
5. The process of claim 4 further characterized in that said pentane/hexane concentrate is reacted with hydrogen in a hydroisomerization reaction zone, at isomerization conditions selected to produce pentane and hexane isomers.
6. The process of claim 1 further characterized in that said hydrocracking catalytic composite comprises from about 0.01 percent to about 2.0 percent by weight of a platinum or palladium component.
7. The process of claim 1 further characterized in that said gem-polyhalo compound is A Friedel-Crafts metal halide.
8. The process of claim 1 further characterized in that said hydrocracking conditions include a maximum catalyst bed temperature of from 300*F. to about 600*F. and a pressure in the range of about 200 to about 500 psig.
9. The process of claim 1 further characterized in that said isobutane is reacted with an olefinic hydrocarbon, in an alkylation reaction zone, at alkylation conditions selected to produce a normally liquid alkylated hydrocarbon stream.
10. The process of claim 2 further characterized in that said third separation zone is a solvent extraction zone.