US 3755140 A
A naphtha boiling range charge stock is converted into aromatic hydrocarbons and isobutane via a combination process involving catalytic reforming, separation and hydrocracking. The catalytic reforming is effected at low-severity conditions to maximize the reaction of naphthene dehydrogenation while simultaneously decelerating the dehydrocyclization and cracking of paraffinic hydrocarbons. Following separation of an aromatic concentrate, the remaining saturated hydrocarbons are subjected to a particular hydrocracking technique which results in exceedingly high yields of isobutane.
Description (OCR text may contain errors)
United States Patent [1 1 51 Aug. 28, 1973 Pollitzer 41 SIMULTANEOUS PRODUCTION OF AROMATlC HYDROCARBONS AND ISOBUTANE Y  lnventor: Ernest L. Pollitzer, Skokie, lll.
 Assignee: Universal Oil Products Company,
Des Plaines, lll.
 Filed: Aug. ll, 1971  Appl. No.: 170,801
52 vs. Cl. 208/62, 208/66 [5-1] Int. Cl C103 37/10  Field of Search 208/62, 66, 87
 References Cited UNITED STATES PATENTS 3,692,666 9/1972 Pollitzer 208/5112 2,908,628 l0/l959. Schneider et al...... 3,114,696 12/1963 Weisz 3,649,520 3/1972 Graven 208/93 Gala/yI/c Reforming Extraction 4/1972 Graven 208/62 3.660.520
Primary Exa m 'ner'Delbert E. Gantz Assistant Examiner-S. Berger Attorney-James R. Hoatson, Jr. et al'. I V
ABSTRACT A naphtha boiling range charge stock is converted into aromatic hydrocarbons and isobutane via a combination process involving catalytic reforming, separation and hydrocracking. The catalytic reforming is effected.
6 Claims, '2 Drawing Figures .5/1972 Hemminger 260/68343 1 SIMULTANEOUS PRODUCTION OF AROMATIC HYDROCARBONS AND ISOBUTANE APPLICABILITY OF INVENTION The present invention involves a multiple-stage process for the conversion of naphtha, or gasoline boiling range hydrocarbons, to produce an aromatic concentrate and exceedingly large quantities of isobutane. More specifically, the inventive concept herein described is directed toward an integrated refinery process for producing a high octane, unleaded gasoline pool.
Aromatic hydrocarbons, principally benzene, toluene, ethylbenzene and xylene isomers, are required in large quantities to satisfy an every-increasing demand for various petrochemicals which are synthesized therefrom. For example, benzene may be hydrogenated to 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. A principal utilization of aromatic hydrocarbons is as gasoline blending componenets in view of their exceedingly high research octane blending values. For example, benzene has a clear research octane blending value of 99, while toluene and all other aromatics have a value in excess of 100. lsobutane finds widespread use in organic synthesis, as a refrigerant and as an aerosol propellant, etc. Other uses include conversion to isobutenes for use in the production of butyl-rubber, copolymer resins with butadiene, acrylonitrile, etc. In accordance with one embodiment of the present invention, the multiplestage process, for an aromatic concentrate and isobutane, is integrated into a refinery scheme for the production of a high octane, unleaded gasoline pool. The aromatic concentrate is sent directly to the unleaded gasoline pool while the isobutane concentrate is subjected to alkylation, with an olefinic hydrocarbon, the normally liquid alkylate product being recovered 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 to the atmosphere stems from vehicular exhaust consisting 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 material. In developing these catalytic converters, it was learned that the efficiency of conversion and particularly the stability of the selected catalytic composite were severely impaired when the exhaust fumes resulted from the combustion of lead-containing fuels. Compared to operations of the catalytic converter during the combustion of clear, unleaded gasolines, both the conversion of noxious components and catalyst stability decreased as much as 50.0 percent when the motor fuel contained lead additives. Therefore, it has been recognized throughout the petroleum industry, as well as in the major gasoline-consuming countries, that suitable gasoline must be produced for consumption in current internal combustion engines without requiring the addition of lead to increase the octane rating. Also being recognized is the fact that unburned hydrocarbons and carbon monoxide are not the only extremely dangerous pollutants being discharged via vehicular exhaust. Japan has recently experienced an increase in the incidence of lead poisoning, and has enacted legislation to reduce significantly the quantity of lead in motor fuel gasolines.
One natural consequence of the removal of lead from motor fuel gasoline, in addition to many others, resides in the fact that petroleum refining operations will necessitate modification in order to produce voluminous quantities of high octane, unleaded motor fuels in an economically attractive fashion. One well-known and well-documented refining process, capable of significantly improving the octane rating of gasoline boiling range fractions, is the catalytic reforming process. In such a process, the primary octane-improving reactions are naphthene dehydrogenation, naphthene dehydroisomerization, paraffin dehydrocyclization and paraffin hydrocracking. Naphthene dehydrogenation is extremely rapid, and constitutes the principal octaneimproving reaction. With respect to a S-membered alkyl naphthene, it is necessary first to effect isomerization to produce a six-membered ring naphthene, followed by dehydrogenation to an aromatic hydrocarbon. Paraffin aromatization is achieved through dehydrocyclization of straight-chained paraffins; this reaction is rate limited in catalytic reforming operations. Unreacted, relatively low octane paraffins, therefore, are present in the catalytically reformed product effluent and effectively reduce the overall octane rating thereof. When operating at a relatively high severity, the paraffinic hydrocarbons within the reforming zone are subjected to cracking. While this partially increases the octane rating of the gasoline boiling range product, substantial quantities of normally gaseous material are produced. In view of the fact that hydrogen is present within the reaction zone, the light gaseous material is substantially completely saturated and comprises methane, ethane, propane and butane.
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, subsequent processing of the saturates for octane rating improvement can be eliminated by increasing the operating severity within the catalytic reforming reaction zone. A high severity operation produces a two-fold effect while increasing the octane rating; first, additional high-octane aromatic components are produced and, secondly, the low octane rating components are at least partially eliminated by conversion either to aromatic components or light normally gaseous hydrocarbons. The results, therefore, include lower liquid yields 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. These problems are further compounded when the desired end result is the production of a high octane, unleaded gasoline pool. In accordance with an overall refinery operation, into which the present invention is integrated, a lowseverity catalytic reforming unit is dove-tailed with at least a separation system, a particular hydrocracking unit and an alkylation unit. As hereinafter indicated, the end result is the production of a high octane, unleaded gasoline pool, in higher yields than would be attainable by direct, high severity reforming alone.
The 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 100 F., and an end boiling point less than about 450 F., and is inclusive of intermediate boiling range fractions often referred to in the art as light naphtha and heavy naphtha. Light naphtha generally refers to a hydrocarbon mixture having an end boiling point in the range of about 280 F. to 340F. These can be recovered directly from a crude distillation unit. A heavy naphtha is considered a hydrocarbon mixture having an initial boiling point of about 180 F. and an end boiling point of about 400 F. to about 450 F. and primarily includes those hydrocarbons having seven or more carbon atoms per molecule. It is not intended, however, to limit the present invention to a charge stock having a particular boiling range. Suffice to say, that suitable charge stocks will generally have an initial boiling point above about 100 F. and an end boiling point below about 450 F. The precise boiling range of any given naphtha charge stock will be dependent upon the economic and processing considerations prevalent in the particular locale where the charge stock is available.
A key feature of the present invention is a particular hydrocracking reaction zone wherein the saturated hydrocarbons, remaining after the separation of the aromatic concentrate from the catalytic reforming effluent, are converted into exceedingly high yields of isobutane.
* OBJECTS AND EMBODIMENTS A principal object of the present invention is the simultaneous production of aromatic hydrocarbons and an isobutane concentrate. A corollary objective resides in the production of a high-octane, unleaded motor fuel gasoline pool.
Another object of my invention is to provide an integrated refinery operation for producing high liquid yields of a high octane, unleaded gasoline pool.
Therefore, in a broad embodiment, the present invention affords a process for the simultaneous production of an aromatic concentrate and an isobutane concentrate, from a naphtha boiling range charge stock,
which comprises the steps of: (a) reacting said charge stock in a low-severity catalytic reforming reaction zone, at reforming conditions selected to convert naphthenic hydrocarbons to aromatic hydrocarbons; (b) separating the resulting reformed product effluent to recover said aromatic concentrate and to provide a saturated normally liquid stream; (c) reacting at least a portion of said normally liquid stream 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, or a nickel component, and the reaction product of alumina and a sublimed Friedel-Crafts metal halide; and, (d) recovering said isobutane concentrate from the resulting hydrocracked product effluent.
A more limited embodiment of my invention relates to a process for producing a high octane, unleaded gasoline pool which comprises the steps of: (a) reacting a naphtha boiling range charge stock, in a low-severity catalytic reforming reaction zone, at reforming conditions selected to convert naphthenic hydrocarbons to aromatic hydrocarbons; (b) separating the resulting reformed product effluent to recover aromatic hydrocarbons and to provide, (1) a heptane-plus saturate stream and (2) a propane/butane stream; (0) reacting at least a portion of said heptane-plus stream 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, or a nickel component, and the reaction product of alumina and a sublimed Friedel-Crafts metal halide; (d) separating the resulting hydrocracked product effluent to provide an isobutane concentrate and alkylating said isobutane with an olefinic hydrocarbon, in an alkylation reaction zone at alkylating conditions selected to produce a normally liquid alkylated hydrocarbon stream; and, (e) recovering said aromatic concentrate and said alkylated hydrocarbon stream as said high octane, unleaded gasoline pool.
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 to provide a pentane/hexane stream which is reacted with hydrogen in a hydroisomerization reaction zone at conditions selected to produce pentane and hexane isomers, said isomers being recovered as part of said high octane, unleaded gasoline pool. In another such embodiment, at least a portion of the propane/butane stream, recovered from the reformed product effluent, is reacted in a dehydrogenation reaction zone to produce a propene/butene concentrate which is reacted in said alkylation reaction zone as said olefinic hydrocarbon.
SUMMARY OF INVENTION As hereinbefore set forth, the present invention involves a catalytic reforming zone, a separation zone and a particular saturate cracking zone. Additionally, in other embodiments, an integrated refinery scheme incorporating the process of the present invention, utilizes a solvent extraction zone, an isomerization reaction zone and an alkylation reaction zone. In a specific embodiment the overall 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 is obtained, a brief description 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 to United States Patents will be made in order that more detail will be readily available where desired. Such references are not intended to be exhaustive or limiting, but merely exemplary and illustrative.
* CATALYTIC REFORMING ZONE The naphtha boiling range charge stock to the catalytic reforming zone may be derived from a multitude of sources. For example, one such source constitutes those naphtha distillates which are derived from a full boiling-range petroleum crude oil; another source is the naphtha fraction obtained from the catalytic cracking of gas oils, while another source constitutes the gasoline boiling range etiluent from a hydrocracking reaction zone processing heavier-than-gasoline charge stock. Since the greater proportion of such naphtha fractions are contaminated through the inclusion of sullytic reforming have indicated that catalyst activity and stability are significantly enhanced through the addition of variousmodifiers,specifically tin, rhenium, nickel and/or germanium.
Suitable porous carrier materials include refractory inorganic oxides such as alumina, silica, zirconia, etc.,
and crystalline aluminosilicates such as the faujasites, or mordenite, or combinations of alumina with the various crystalline aluminosilicates. Generally favored metallic components include ruthenium, rhodium, palladium, osmium, iridium, platinum, rhenium, germa nium, nickel and/or tin. These metallic components are employed in concentrations ranging from about 0.01 percent to about5.0 percent by weight, and preferably from about 0.01 percent to about 2.0 percent by weight. Reforming catalysts may also contain combined halogen selected from the group of fluorine, chlorine, bromine, iodine and mixtures thereof.
Illustrations of catalytic reforming process schemes are found in US. Pat. No. 2,905,620 (Cl. 208-65), 3,000,812 (Cl. 208-138) and 3,296,118 (Cl. 208-l00). Effective reforming operating conditions include catalyst temperatures within the range of about 800 F. to about 1,100 F., preferably having an upper limit of about l,050 F. The liquid hourly space velocity, defined as volumes of hydrocarbon charge per hour per volume of catalyst disposed within the reforming reaction zone, is preferably in the range of about 1.0 to about 5.0, although space velocities from about 0.5 to about 15.0 may be employed. The quantity of hydrogen-rich gas in admixture with the hydrocarbon feed stock to the reforming reaction zones is generally from about 1.0 to 20.0 moles of hydrogen per mole of hydrocarbon. The reforming reaction zone effluent is generally introduced into a high-pressure separation zone at a temperature of about 60 F. to about 140 F., to separate lighter components from heavier, normally liquid components. Since normal reforming operations produce large quantities of hydrogen, a certain amount of the recycle gaseous 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 hydrogen-consuming hydrocracking reaction zone as make-up hydrogen, as well as in the hydroisomerization reaction zone. Pressures in the range of about 100 to about 1,500 psig. are suitable foreffecting catalytic reforming reactions.
With respect to the catalytic reforming reaction zone utilized in the present combination process, the reactions effected therein are conducted at a relatively low operating severity. To those familiar with the catalytic reforming art, the term relatively high severity indicates high temperature or low space velocity, or both high temperature and low space velocity. The most noticeable direct result of a 'high severity operation is found in the octane rating of the normally liquid reformed product effluent. Whilethe reforming zone utilized in-the present process does not necessarily upgrade the octane rating of the charge stock to the level ultimately attained with respect to the unleaded gaso line pool, the charge stock is substantially improved in octane rating. I
In the presentspecification and appended claims, the
term low+severity reforming alludes'to a reforming process in which substantial quantities of naphthenic hydrocarbons are 'dehydrogenated tohigh octane arcmatic hydrocarbons, while thedehydrocycli'zation and cracking of paraffinic hydrocarbonsis inhibited. Lowseverity reforming operations may be defined by stating that from about 8'0.0 to about 100.0 moles of aromatics areproduced for every 100.0moles of naphthenes in the charge stock, while lessthan about 40.0 moles of aromatics are produced for every 100.0'moles of alkanes. In determining the degree of conversion of naphthen'es to aromatics (dehydrogenation) and al kanes to aromatics (dehydrocyclization), it is assumed that a relatively small amount of naphthenes' are cracked'or otherwise converted to hydrocarbons other than aromatics, and that a major portion of the alkanes which disappear areconve'rted to'aromatic hydrocarbons with some naphthenes and higher molecular weight alkanes being" converted to low molecular weight normally gaseous components.
* AROMATIC SEPARATION ZONE tained therein. Although any separation scheme maybe 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 raffmate, is a well known technique which is thoroughly described in the literature. Suitable tech niques involve the operations illustrated in US. Pat. Nos. 2,730,558 (Cl. 260674') and 3,361,664 (Cl. 208 313).
As previously stated, the product effluent from the reforming reaction zone is generally-introduced into a high-pressure separator at a temperature sufficient to provide a normally liquid hydrocarbon phase and a hydrogen-rich recycle gaseous phase. Other separations, contemplated within the scope of the present invention, which may be considered either within the catalytic reformingsystem or the aromatic separation system, include the recovery of a propane/butane concentrate, an ethane-minus gaseous phase and a pentanelhexane concentrate. lt is also within the scope of the present invention to introduce the total pentane-plus portion of thereformed product effluent into the solvent extraction zone, subsequently separating a pent'ane/hexane fraction from theparaffinic'raffinate.
In any event, the solvent extraction process utilizesa solvent having a-greater'selectivity and solvency for the aromatic components of 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. The particular solvent is one which boils at a temperature above the boiling point of the hydrocarbon mixture at the 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 the glycols, including ethylene glycol, propylene glycol, butylene glycol, tetraethylene glycol, glycerol, diethylene glycol, dipropylene glycol, dimethylether of ethylene glycol. triethylene glycol, tri-propylene 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 solvents are those characterized as the sulfolane-type. Thus, as indicated in U.S. Pat. No. 3,470,087 (Cl. 208321), the 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 includes the sulfolenes such as 2-sulfolene and 3-sulfolene.
The aromatic selectivity of the preferred solvents can be further enhanced by the addition of water. This increases the selectivity of the solvent phase for aromatic hydrocarbons over non-aromatic hydrocarbons without reducing substantially the solubility of the solvent phase for aromatic hydrocarbons. In general, the solvent composition contains 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, depending primarily on the particular solvent and the process conditions under which the extraction, extractive distillation and solvent recovery zones are operated.
In general, solvent extraction is conducted at elevated temperatures and pressures selected to maintain the charge stock and solvent in the liquid phase. Suitable temperatures are within the range of about 80 F. to about 400 F., and preferably from about 150 F. to about 300 F. Operating pressures include superatmospheric pressures up to about 400 psig., and preferably from about 15.0 psig. to about 150 psig.
Typical extractive distillation zone pressures are from about atmospheric to about 100 psig., although the pressure at the top of the distillation zone will generally be maintained in the range of about 1 psig. to about 20 psig. The reboiler temperature is dependent upon the composition of the feed stock and the selected solvent, although temperatures of from about 275 F. to about 360 F. appear to yield satisfactory results. The solvent recovery system is operated at low pressures 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 pressures of from about 100 to about 400 millimeters of mercury absolute. These low pressures must be used since the reboiler temperature should be maintained below about 370 F. in order to avoid thermal decomposition of the organic solvent.
* HYDROCRACKING REACTION ZONE In a preferred embodiment, the charge to the hydrocracking reaction zone will be the heptane-plus paraffinic concentrate remaining in the catalytically reformed product efi'luent following removal of the aromatics in the solvent extraction zone. Although the charge may contain the pentane/hexane paraffins, a
preferred technique, as hereinafter set forth, involves separate recovery of a pentane/hexane concentrate for utilization as the charge to an isomerization reaction zone wherein the same is converted into pentane and hexane isomers of significantly increased octane rating. The hydrocracking reaction zone of the present process is unlike present-day hydrocracking processes both in function and result. Initially, the charge to the hydrocracking reaction zone constitutes paraffinic hydrocarbons boiling within the naphtha boiling range, and the product effluent contains very little, if any, methane and ethane. In those instances where the product effluent contains propane, the same can be utilized for subsequent alkylation or isopropyl alcohol production; another valuable use is as LPG. Through the utilization of a particular catalytic composite and operating conditions, the cracking of the paraffinic raffmate produces relatively large quantities of butanes, which butane concentrate is rich in isobutanes. In view of the unique character of the product effluent, being exceedingly rich in isobutane, the hydrocracking reaction zone is referred to herein as l-cracking. Thus, with respect to increasing the yield of normally liquid hydrocarbons in the unleaded gasoline pool, the butane concentrate can be subjected to alkylation with suitable olefinic hydrocarbons. The hydrocracking reaction conditions, under which the process is conducted, will vary according to the physical and chemical characteristics of the charge stock. In the past, hydrocracking reactions have 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 scfJBbl. and maximum catalyst bed temperatures in the range of about 700 F. to about 950 F.
As discussed in the prior art, the heavier charge stocks require a relatively high severity of operation including high pressures, high catalyst bed temperatures, a relatively low liquid hour space velocity and high hydrogen concentrations. A lowerseverity of operation is employed with comparatively lighter feed stocks such as kerosenes and light gas oils. Through the practice of the present invention, regardless of the characteristics of the charge stock, the hydrocracking process is effected at a relatively lower severity than those commonly in use. In accordance with the present invention, the hydrocracking reaction zone has disposed therein a catalytic composite comprising a Group VIII noble metal component, or a nickel component, and the reaction product of alumina and a sublimed Friedel-Crafts metal halide. The conversion conditions include a liquid hourly space velocity of 0.5 to about 10.0, a hydrogen-circulation rate of about 3,000 to about 20,000 scf./Bbl., a pressure from about 200 to about 2,000 psig., preferably from about 500 psig. to 1,000 psig., and, of greater significance, a maximum catalyst bed temperature of from 300F. to about 480 F. In many instances, particularly with respect to the naphtha boiling range charge stocks, the operating pressure will consistently be in the range of about 200 to about 500 psig., the hydrogen concentration from about 3,000 to about 10,000 scf./Bbl. and the liquid hourly space velocity from 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 utilizes a catalytic composite containing a Group VIII noble metal component, or a nickel component, and the reaction product of alumina and a sublimed Friedel-Crafts metal halide. Thus, where the metal halide is, for example, aluminum chloride, the catalyst is characterized in that it contains the following group:
preferred prior art carrier material appears to be acomposite 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 inthe area of catalysis have further shown that various crystalline aluminosilicates can be utilized to advantage in somehydrocracking situations. Suchzeolitic material includes mordenite, faujasite, Type A or TypeU mo--- lecular sieves, etc.
In view of the fact that a sublimed Friedel-Crafts.
metal halide is not strong enough to react with silica, to form the type of group hereinbefore described, the preferred carrier is alumina. While the action and effeet of the sublimed metal halideon refractory material other than aluminaand silica, for example, zirconia, is not known with accuracy, it is not believed that reaction takes place to a degree sufficient to produce-the desired catalyst and result.
The catalytic composite contains a Group VIII noble metal component, or a nickel component. Thus, suitable metals are those of the group including platinum, palladium, rhodium, ruthenium, osmium, iridium and nickel. Iron and cobalt components do not appear to possess the propensity for effecting the desired degree of hydrocracking, and are, therefore, excluded from the group of suitable metallic components. A particularly preferred catalytic composite contains a platinum, palladium or nickel component. These metal omponents, for example, platinum, may exist within the final composite as a compound such as an oxide, sulfide halide, etc., or in an elemental state. Generally the amount of the noble metal component is small compared to the quantities of the other components combined therewith. Calculated on an elemental basis, the noble metal component generally comprises from about 0.01 percent to about 2.0 percent by weight of the final composite. With respect to the nickel component, again calculated on the basisof the elemental metal, it will be present within the catalytic composite in an amount of from about 1.0 percent to about 10.0 percent by weight.
The metallic components may be incorporated within the catalytic composite in any suitable manner including co-precipitation or co-gellation with the carrier material, ion-exchange or impregnation. The latter constitutes a preferred method of preparation, utilizing water-soluble compounds of the metallic components. Thus, a platinum component may be added to the 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,
An essential ingredient of the catalytic composite is a Friedel Crafts metal halide which, when sublimed, combines with the alumina by way of reaction therewith. The method of incorporating-the Friedel-Crafts metal halide involves asublimation, or vaporization technique,zwith the vaporized metal halide contacting alumina containing the Group VIII 'noble metal component, or th enickel c'omponent. That is, the catalytic active. metallic component is already composited withthe alumina before the latter is contacted with the sub limed 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. When the sublimation technique is utilized, themetal halide will be vaporized onto the carrier, and then heated'to a-temperature of about 300 C., and for a time sufficient to remove any unreacted metal halide. Thus, the-finalcatalytic composite does not containany freeFriedel-Crafts metal halide. Following vaporization of the Friedel-Crafts metal halide, and heating of the thus-formed composite, the refractory oxide will be increased in weight by from about 2.0 percentto about 25.0 percent basedupon the original weight of the carrier material. While-the exact increase in weight does not appear to-becritical, 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. On thebasis of the quantity of the metal halide combined therewith, the treated carrier material will contain from about 1.96 percent to about 20.0 percent by weight of the metal halide, and preferably from about 4.76 percent to about 16.67 percent by weight, as the metal halide. Further details of this sublimation technique may be found in U.S. Pat. No. 2,924,628 (Cl. 260-666). Since the desired groups, as hereinbefore set forth, are sensitive to moisture, the sublimation technique is effected after the Group VIII noble metal component, or nickel component, has been combined with the alumina.
Various Friedel-Crafts metal halides may be utilized, but not necessarily with equivalent results. Examples of such metal halides include aluminum bromide, aluminum chloride, antimony pentachloride, beryllium chloride, germanium tetrachloride, ferric bromide, ferricv chloride, gallium trichloride, stannie bromide, stannic chloride, titanium tetrabromide, titanium tetrachloride, zinc bromide, zine-chloride and zirconia chloride. The Friedel-Crafts aluminum halides are preferred, with aluminum chloride and/or aluminum fluoride being particularly preferred. This is so, not only due to the ease of preparation, but also because the thusprepared catalysts have an unexpectedly high activity for the selective production of isobutane.
Temperatures at which the FriedeLCrafts metal halide is vaporized onto the alumina will vary'in accordancewith the particular metal halide utilized. In most instances, the vaporization is carried out either at the boiling, or sublimation'point of the particular Friedel- Crafts metal halide, or at a temperature not greatly exceeding these points; for example, not greater than 100 C. higher than the boiling point, or sublimation point. In effecting one catalyst preparation, the amorphous carrier material has aluminum chloride sublimed thereupon. Aluminum chloride sublimes at 178 C., and thus a suitable vaporization temperature will range from about 180 C. to about 275 C. The sublimation technique may be carried out under pressure, and also in the presence of diluents such as inert gases. Although the particularly preferred technique involves the sublimation of a metal halide directly to react with the alumina, the reaction product can result froma halide-containing compound which initially reacts with the alumina to form an aluminum halide which, in turn, reacts with additional alumina, thereby forming groups of -Al-O-AlCl Such halide-containing compounds include CCI SCl,, SOCl PCl POCl etc. Prior to its use, the catalytic composite may be subjected to a substantially water-free reduction technique. This is designed to insure a more uniform and finely-divided dispersion of the metallic components throughout the carrier material. Substantially pure and dry hydrogen is employed as the reducing agent at a temperature of about 800 F.to about l,200 F., 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 hydrogen and paraffinic raffinate traverse the catalyst bed. In accordance with the present process, the maximum catalyst bed temperature, virtually the same as that measured at the outlet of the reaction zone, is maintained in the range of about 300 F. to about 480 F. In order to assure that the catalyst bed temperature does not exceed the maximum allowable, the use of conventional quench streams, either normally liquid, or normally gaseous and introduced at one or more intermediate loci of the catalyst bed, is contemplated.
As hereinbefore set forth, the product effluent from the hydrocracking reaction zone is predominantly butanes, the greater proportion of which constitute the various isobutanes. For this reason, the hydrocracking reaction zone is herein referred to as l-cracking, the I alluding to isomer production.
* ALKYLATION ZONE Since the preferred use of the present inventive concept is the integration thereof into an overall refinery scheme for the production of a high octane, unleaded motor fuel gasoline pool, the isobutane-rich effluent from the l-cracking zone is utilized as fresh feed to an alkylation reaction zone. The alkylation is effected by intimately commingling the isobutane feed, olefinic hydrocarbons and a particular catalyst as hereinafter de' scribed. It is understood that the particular 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 brought into the described process from any suitable source such as a fluid catalytic cracking unit, or a thermal cracking unit. However, as stated in another specific embodiment of the present invention, at least a portion of the isobutane concentrate is subjected to dehydrogenation in a dehydrogenation reaction zone to produce the alkylatable olefinic hydrocarbons. Similarly, the propane/butane concentrate obtained via the separation of the catalytically reformed product effluent 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 fluoridecatalyzed system, or one which utilizes sulfuric acid. Hydrogen fluoride alkylation is particularly preferred, and may be conducted substantially as set forth in US Pat. No. 3,249,650 (Cl. 260683.48). Briefly, the allrylation 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 10.0 percent water. Excessive dilution with water is generally to be 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 tendency of the olefinic portion of the charge stock to undergo polymerization prior to alkylation, the molar proportion of isoparafiins to oleflnic hydrocarbons 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 0 to about 200 F., and preferably from about 30 F. to about 125 F. The pressure maintained within the alkylation system is ordinarily at a level sufi'icient 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 con- I veniently expressed in terms of spacetime, being 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 spacetime 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, along with isopentane and isohexane from the I-cracking reaction zone. Unreacted isobutane and olefinic hydrocarbons, if any, 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 As previously indicated, a significant quantity of pentanes and hexanes are produced in the catalytic reforming reaction zone. Additionally, in those instances where the fresh feed charge stock to the process is a full boiling range naphtha distillate, the same may contain a pentane/hexane concentrate. In view of the fact that nonnal pentane has a clear research octane rating of 62 and normal hexane a clear research octane rating of 25, these components are not desirable in a gasoline pool which is intended to be free from lead additives.
In still another embodiment of the present invention, therefore, the pentane/hexane stream is introduced into an isomerization reaction zone for the purpose of producing an effluent product rich in pentane and hexane isomers. For example, iso-pentane has a reseach clear octane rating of 93, while 2,2-dimethylbutane has a rating of 92 and 2,3-dimethylbutane a rating of 104; the average clear research octane rating of the monomethylpentanes is 74. 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 through the production of pentane/hexane isomers without incurring a detrimental volumetric yield loss.
As indicated in U.S. Pat.No. 3,131,235 (Cl. 260-683.3), 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 naturallyoccurring refractory inorganic oxides. Of these, a synthetically-prepared gamma-alumina is preferred. The Group VIII noble metal is generally present in an amount of about 0.01 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 paltinum 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 2.5 percent to about 5.0 percent by weight, is preferred.
The isomerization reaction is 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 200 F. to about 800 F., although temperatures within the more limited range of about 300 F. to about 525 F. will generally be utilized. The pressure, under which the reaction zone is maintained, will range from about 50 to about 1,500 psig. The reaction products are separated from the hydrogen, which is recycled, and subjected to fractionation and separation to produce the desired reaction product. Recovered starting material is also recycled so that the overall process yield is high. The liquid hourly space velocity will be maintained in the range of about 0.25 to about 10.0, and preferably within the range of about 0.5 to about 5.0. Another suitable isomerization process is found in U.S. Pat. No. 2,924,628 (Cl. 260-666).
* DEHYDROGENATION ZONE As hereinbefore set forth, at least a portion of the isobutane-rich effluent from the I-cracking reaction zone may be subjected to dehydrogenation to produce the olefins required to alkylate the same in the alkylation reaction zone. In another embodiment, at least a portion of the propane/butane concentrate recovered from the catalytic reforming reaction zone effluent may also be subjected to dehydrogenation. 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 deemded desirable, it may be effected essentially as set forth in U.S. Pat. No. 3,293,219 (Cl. 260-683.3). Briefly, dehydrogenation reactions are generally effected at conditions including a temperature in the range of from 400 C. to about 700 C., a pressure from about atmospheric to about 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 1:1 to about 10:1 based upon the paraffin charge.
The dehydrogenation catalyst is a composite of an inorganic oxide carrier material, an alkali metal component, a Group VIII metal component and a catalytic attenuator from the group consisting of arsenic, antimony and bismuth. A particularly preferred catalyst comprises lithianted alumina containing about 0.05 percent to about 5.0 percent by weight of a Group VIII noble metal, especially platinum. The catalytic attenuator is employed in amounts based upon the concentration of Group VIII 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, mixtures thereof, etc. Still another preferred catalyst contains, in addition to the noble metal component, a component from the group of tin, germanium and rhenium.
The dehydrogenation conditions and catalysts are selected to result in relatively low conversion per pass, accompanied, however, 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 normal paraffins over olefinie hydrocarbons, the high selectivity and relatively low conversion, in the dehydrogenation zone, are advantageous.
* DESCRIPTION OF DRAWING The inventive concept, encompassed by the present process, and a particularly preferred embodiment, are illustrated in the accompanying two drawings. The illustrations are presented by way of block-type flow diagrams, in which each block represents one particular step, or stage of the process. Miscellaneous appurtenances, not believed necessary for a clear understanding in the present combination process, have been eliminated from the drawing. The use of details such as pumps, compressors, instrumentation and controls, heat-recovery circuits, miscellaneous valving, start-up lines and similar hardware, etc., is well within the purview of one skilled in the art. Similarly, 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 recylce lines, and vent gas streams, etc., have also been eliminated. FIG. 1 illustrates the basic inventive concept wherein the normally liquid portion of a catalytic reforming reaction zone effluent is subjected to extraction to recover aromatics and to provide a paraft'inic raffinate which is subjected to lcracking. FIG. 2 presents the integration of the inventive concept in a preferred embodiment encompassing 7 both an isomerization reaction zone and an alkylation reaction zone.
With reference now to the drawings, they will be described in conjunction with a commercially-sealed unit designed to process a C -plus straight-run naphtha fraction which had previously been hydrorefined for contaminant removal and olefin saturation. The intended object is the maximum production of a high octane, unleaded gasoline pool from 100,000 Bbl./day of the fresh charge stock. Pertinent properties of the naphtha fraction include a gravity of 564 API, an initial boiling point of 194 F., a 50.0 percent volumetric distillation temperature of 255 F. and an end boiling point of 362 F. Hydrocarbon-type analyses indicate that the fresh feed consists of 44.4 vol. percent paraffins, 48.8 vol. percent naphthenes and 6.8 vol. percent aromatics.
The charge stock enters the process by way of line 1, and is introduced thereby into catalytic reforming zone 2. Reforming zone 2 constitutes a low-severity catalytic reforming system intended to produce maximum quantities of a 90.0 research octane normally liquid product effluent. The operating conditions are selected to maximize the dehydrocyclization of naphthenes to aromatics while simultaneously minimizing the hydrocracking of paraffins. Therefore, the naphtha feed is reformed at conditions including a pressure of about 150 psig., a liquid hourly space velocity of 3.0, a hydrogen molal concentration of 6.0 and an average catalyst bed temperature of about 900 F. 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 the basis of the elements.
The product effluent is separated to provide a hydrogen-rich recycle gas stream, a methane-ethane vent gas stream in line 3 and a propane-butane concentrate in line 4. In the embodiment illustrated, the C -plus liquid portion of the product effluent is introduced through line 5 into extraction zone 6, from which a pentanelhexane concentrate is withdrawn through line 7. It is contemplated that the separation system employed for the recovery of the pentane/hexane fraction may be an integral part of the reforming system, and that the material in line 5 constitutes the heptane-plus liquid portion of the reformed product effluent. The pentaneplus reformate has a clear research octane rating of 90.3; component yields and product distribution are presented in the following Table I:
TABLE I: Reforming Zone Effluent Distribution Component WL 11 Vol. I; Hydrogen 2.64
Propane 1.10 1.59 lsobutane 0.47 0.63 N-butane 0.37 0.48
lsopentane 1.82 2.19 N-pentane 0.65 0.78 Hexane-plus 91.49 85.75
Catalytic reforming is a hydrogen-producing process, and the 2.64 percent by weight of hydrogen, or 1,310 scf./Bbl., may be utilized to advantage in the I-cracking zone wherein hydrogen-consuming reactions are effected.
Of the hexane-plus portion, 6.7 percent by volume (5,750 Bbl./day) constitutes hexanes which are withdrawn through line 7 along with the 2,970 Bbl./day of pentanes. The remaining 80,000 Bbl./day is introduced into a lower portion of a solvent extraction column countercurrently to a lean solvent stream which is introduced into an upper portion of said column, the mole ratio of solvent to hydrocarbon being about 32:10. The selected solvent is sulfolane, and the extraction column functions at a top pressure of about 15.0 psig. and a reboiler temperature of about 320 F. A saturate-rich raff'mate stream is withdrawn as an overhead product, while the rich solvent bottoms stream is introduced into an extractive distillation zone. Additional rafi'inate is withdrawn as an overhead stream, combined with the saturate-rich raft'mate from the extraction column, and passed through line 9 into lcracking zone 11. Rich solvent is introduced into a solvent recovery system functioning at sufficiently low pressures and high temperatures to drive aromatic hydrocarbons overhead while producing a lean solvent bottoms stream for recycle to the extraction column. The aromatic concentrate, in an amount of 48,450 BbL/day (based upon 100,000 barrels of fresh feed) is withdrawn from reaction zone 6 via line 8. In regard to the raffinate stream in line 9, beneficial results are obtained, with respect to the subsequent l-cracking reactions, when the stream is substantially free from the selected solvent. One suitable technique, for removing solvent from the rafiinate stream in line 9, is that disclosed in US. Pat. No. 3,470,087 (Cl. 208-321).
The component analysis of the heptaneplus raffinate stream, charged to the l-cracking zone 11, is presented in the following Table II; the distribution being based upon the 31,550 Bbl./day of the raffinate stream:
TABLE II: Rafi'mate Stream Analysis Com nent Vol. 31: C-7 araffins 33.6 C-8 Parafi'ms 27.7 C-9 Paraffin: 15.3 C-l0 Paraffins 9.8 011 Paraffins 2.7
O? Naphthenes 3.4 C-8 Naphthenes 5.] C-9 Naphthenes 1.9 C-l0 Naphthenes 0.5
The raffinate stream is introduced into l-cracking zone 11 by way of line 9, and is admixed with hydrogen in an amount to yield a hydrogen/hydrocarbon mol ratio of about 6.0:l.0. Other operating conditions in clude a pressure of about 750 psig., a liquid hourly space velocity of 1.0, based upon combined feed which includes about 10,950 BbL/day of unconverted heptane-plus material, and a catalyst bed temperature of 350 F. to 400 F. (a temperature gradient of 50 F.). The catalyst is a composite of alumina, 5.0 percent by weight of nickel and 7.5 percent by weight of aluminum chloride sublimed thereon to react with the alumina as aforesaid.
The hydrocracked product efiluent is separated to provide a normally vaporous phase in line 10, containing propane and lighter gaseous hydrocarbons, a butane concen-trate in line 12 and a pentane-plus normally liquid stream in line 13. The latter containing 10,950 BbL/day of heptane-plus hydrocarbons, may be further separated to recover the pentanes and hexanes, in which case the heptane-plus material is recycled to TABLE III: I-cracking Yields and Product Distribution Component Vol. Bbl/day Hydrogen (1007) (ScfJBbL) Methane 82.5 (Scf./Bbl.)
Ethane 16.5 (Scf./Bbl.) Propane 2 .1 7,290 lsobutane 72.2 22,800 N-butane 9.5 3,000 lsopentane 10.5 3,240 N-pentane 1.8 568 lsohexanes 7.0 2,210 N-hexane 1.4 442 The distinct advantage of the particular hydrocracking reactions, to which the substantially aromatic-free raftinate is subjected, are readily ascertained from foregoing Table Ill. Hydrogen consumption is 1,007 scf./Bbl., and only 82.5 scf./Bbl. of methane and 16.5 scf./Bbl. of ethane are produced. A total of 25,800 Bbl./day of butanes are produced, of which about 88.3 percent by volume constitutes isobutane. The breakdown of the isohexane fraction indicates 220 Bbl./day of 2,2-dimethylbutane, 252 BbL/day of 2,3- dimethylbutane, 1,140 Bbl./day of 2-methylpentane and 662 BbL/day of 3-methylpentane.
The following Table IV summarizes the results obtained from the present inventive concept as illustrated in FIG. 1. Only propane and heavier components are indicated since they are significantly more valuable than methane and/or ethane. The yields are inclusive of 5 the propane, butanes, pentanes and hexanes recovered from the catalytic reforming reaction zone.
TABLE IV: Overall Volumetric Yields and Product Distribution Component Vol. BbL/day Propane 8.93 8,880 lsobutane 23.55 23,430 N-butane 3.50 3,480 lsopentane 5.46 5,430 N-pentane 1.35 1,348 2,2-dimethyl butane 0.65 643 2,3-dirnethyl butane 0.74 734 Z-methyl pentane 3.21 3.195 3-methyl pentane 1.94 1.932 N-hexane 1 .97 1 ,962 Aromatic Gasoline 48.70 ,4
TABLE V: Clear Gasoline Pool, FIG. 1
Component Bbl/day Vol.% Research Octane lsopentane 5,430 8.53 93 N-pentane 1,348 2.12 62 2,2-DiMeBu 643 1.01 92 2,3-DiMeBu 734 1.15 104 Z-MePe 3,195 5 .02 74 3-MePe 1,932 3.03 74 N-hexane 1,962 3 .07 25 Aromatic Gasoline TOTALS The 8,880 Bbl./day of propane can be subjected to dehydrogenation to produce propylene which can then be converted by hydrolysis into isopropyl' alcohol having a clear research octane rating of about 1 10 to about 120. This, of course, further increases both the yield and octane rating of the unleaded gasoline pool. On the other hand, the propylene could be employed in an alkylation reaction zone for the production of C alkylate having a clear research octane rating of about 92.0.
A preferred embodiment, utilizing the inventive concept above described, is illustrated in the accompanying FIG. 2, in which an alkylation zone and an isomerization zone have been added to the overall scheme. In this situation, the propane will again be withdrawn from the process as a by-product stream, normal hexane and normal pentane will be subjected to isomerization and the isobutane concentrate will be alkylated with outside butylenes from a catalytic cracking unit. As previously set forth, the hydroret'med naphtha feed is introduced, via line 1, into catalytic reforming zone 2. A propane-minus phase is withdrawn through line 3 and further separated to recover the propane and to concentrate the hydrogen for recycle to reforming zone 2. Butanes are recovered in line 4, and introduced thereby into alkylation zone 14. Pentanes and heavier hydrocarbons are introduced via line 5 into solvent extraction zone 6 wherein the aromatic concentrate is recovered via line 8, and from which a pentane/hexane concentrate is removed via line 7, for introduction into isomerization zone 17.
The heptane-plus rafi'inate is again introduced via line 9 into I-cracking zone 11, with propane and lighter hydrocarbons being withdrawn through line 10. The isobutane concentrate is introduced into alkylation reaction zone 14 via line 12, and pentanes and heavier components are recycled to the extraction zone for removal of the pentanes and hexanes. The operating conditions in both the l-cracking and catalytic reforming reaction zones are identical to those previously set forth with respect to the description of FIG. 1.
Prior to introducing the pentane/hexane concentrate into hydroisomerization zone 17, a preferred technique involves removal of the pentanelhexane isomers which are sent directly to unleaded gasoline pool 16. Therefore, the feed to isomerization zone 17 will constitute 1,348 Bbl/day of normal pentane and 1,962 Bbl/day of normal hexane. Isomerization zone 17 utilizes a fixedbed of catalytic composite of alumina, about 4.0 percent by weight of aluminum chloride and 0.375 percent by weight of platinum, calculated as the elemental metal. The reactions are efiected at a pressure of about 300 psig., a temperature of about 330 F. and a hydrogen to hydrocarbon molal ratio of about 1.0: l .0; the reactants traverse the catalyst bed at a liquid hourly space velocity of about 1.0. Conversion to isomers is about 99.0 percent efl'icient and with the volumetric increase due to molecular size, and conversion of some hexane to isopentane, 1,370 Bbl/day of isopentane, 669 Bbl./day of 2,2-dirnethyl butane, 214' Bbl./day of 2,3- dimethylbutane, 636 Bbl./day of Z-methylpentane and 365 BbL/day of 3-methylpentane are introduced via line 18 into unleaded gasoline pool 16.
Alkylation zone 14 is a hydrofluoric acid system which requires 20,400 BbL/day of outside butylenes to produce 36,100 BblJday of C -alkylate having an octane rating of about 97.0. The reaction time, utilizing a pumped acid settler reaction, is about nine minutes, and the acid to hydrocarbon ratio is about 1.5:1 .0. The alkylation reactions are effected at a temperature of about 100 F. and a pressure of about 20 atmospheres. Following separation of unreacted isobutanes, which are recycled, the alkylate gasoline passes through line 15 into unleaded gasoline pool 16.
One preferred technique constitutes introducing the 3,480 BbL/day of butane into the isomerization zone for conversion into isobutane which is also alkylated in alkylation zone 14. At a conversion efficiency of 99.0 percent, and with the volumetric increase due to molecular size, an additional 3,575 Bbl./day of isobutane becomes available. In this situation, 27,005 Bbl/day of isobutanes require 23,483 Bbl/day of outside butylenes to produce 41,564 Bbl/day of C -alkylate.
The unleaded clear gasoline pool, including the 3,480 BbL/day of butanes which are unreacted in the alkylation reaction zone, has the characteristics shown in the following Table VI:
TABLE VI: Clear Gasoline Pool, FIG. 2
Component BbL/day Vol. Research Octane N-butane 3,480 3.38 94 lsopentane 6,800 6.58 93 2,2-DiMeBu 1,312 1.27 92 2,3-DiMeBu 948 0.92 104 Z-MePe 3,831 3.71 74 3-MePe 2,297 2.22 74 C -alkylate 36,100 34.98 97 Aromatic Gasoline 48,450 46.94 l 15 TOTALS 103,218 100.00 103.7
Based upon the 100,000 Bbl./day of fresh feed and the additional 20,400 BbL/day of outside butylenes, the volumetric yield of the gasoline pool, having a clear research octane of 103.7, is 85.7 percent.
The foregoing demonstrates the method by which the present invention is effected and the benefits afforded through the utilization thereof.
1 claim as my invention:
1. A process for the simultaneous production of an aromatic concentrate and an isobutane concentrate, from a naphtha boiling range charge stock which comprises the steps of:
a. reacting said charge stock in a low-severity catalytic reforming reaction zone, at reforming conditions selected to convert naphthenic hydrocarbons to aromatic hydrocarbons;
b. separating the resulting reformed product effluent to recover said aromatic concentrate and to provide a saturated normally liquid stream;
c. reacting at least a portion of said normally liquid stream 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, or a nickel component, and the reaction product of alumina and a sublimed Friedel-Crafts metal halide; and,
d. recovering said isobutane concentrate from the resulting hydrocracked product effluent.
2. The process of claim 1 further characterized in that said reformed product effluent is separated in a solvent extraction zone.
3. The process of claim 1 further characterized in that said hydrocracking conditions include a maximum catalyst bed temperature from 300 F. to 480 F., a liquid hourly space velocity in the range of 1.0 to about 10.0, a hydrogen concentration of 3,000 to about 20,000 scf./Bbl. and a pressure from about 500 to about 2,000 psig.
4. The process of claim 1 further characterized in that said hydrocracking catalytic composite comprises from about 0.1 percent to about 2.0 percent by weight of a platinum or palladium component.
5. The process of claim 1 further characterized in that said hydrocracking catalytic composite comprises from about 1.0 percent to about 10.0 percent by weight of a nickel component.
6. The process of claim 1 further characterized in that said Friedel-Crafts metal halide is an aluminum halide.