US3755140A - Simultaneous production of aromatic hydrocarbons and isobutane - Google Patents

Simultaneous production of aromatic hydrocarbons and isobutane Download PDF

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US3755140A
US3755140A US00170801A US3755140DA US3755140A US 3755140 A US3755140 A US 3755140A US 00170801 A US00170801 A US 00170801A US 3755140D A US3755140D A US 3755140DA US 3755140 A US3755140 A US 3755140A
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E Pollitzer
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Honeywell UOP LLC
Universal Oil Products Co
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    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G59/00—Treatment of naphtha by two or more reforming processes only or by at least one reforming process and at least one process which does not substantially change the boiling range of the naphtha
    • C10G59/02—Treatment of naphtha by two or more reforming processes only or by at least one reforming process and at least one process which does not substantially change the boiling range of the naphtha plural serial stages only
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10L—FUELS NOT OTHERWISE PROVIDED FOR; NATURAL GAS; SYNTHETIC NATURAL GAS OBTAINED BY PROCESSES NOT COVERED BY SUBCLASSES C10G, C10K; LIQUEFIED PETROLEUM GAS; ADDING MATERIALS TO FUELS OR FIRES TO REDUCE SMOKE OR UNDESIRABLE DEPOSITS OR TO FACILITATE SOOT REMOVAL; FIRELIGHTERS
    • C10L1/00—Liquid carbonaceous fuels
    • C10L1/04—Liquid carbonaceous fuels essentially based on blends of hydrocarbons
    • C10L1/06—Liquid carbonaceous fuels essentially based on blends of hydrocarbons for spark ignition
    • Y—GENERAL TAGGING OF NEW TECHNOLOGICAL DEVELOPMENTS; GENERAL TAGGING OF CROSS-SECTIONAL TECHNOLOGIES SPANNING OVER SEVERAL SECTIONS OF THE IPC; TECHNICAL SUBJECTS COVERED BY FORMER USPC CROSS-REFERENCE ART COLLECTIONS [XRACs] AND DIGESTS
    • Y02—TECHNOLOGIES OR APPLICATIONS FOR MITIGATION OR ADAPTATION AGAINST CLIMATE CHANGE
    • Y02P—CLIMATE CHANGE MITIGATION TECHNOLOGIES IN THE PRODUCTION OR PROCESSING OF GOODS
    • Y02P20/00—Technologies relating to chemical industry
    • Y02P20/50—Improvements relating to the production of bulk chemicals
    • Y02P20/582—Recycling of unreacted starting or intermediate materials

Definitions

  • 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.
  • 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.
  • 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.
  • 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.
  • 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.
  • Naphthene dehydrogenation is extremely rapid, and constitutes the principal octaneimproving reaction.
  • 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.
  • 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.
  • the light gaseous material is substantially completely saturated and comprises methane, ethane, propane and butane.
  • 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.
  • a lowseverity catalytic reforming unit is dove-tailed with at least a separation system, a particular hydrocracking unit and an alkylation unit.
  • 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.
  • 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.
  • 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.
  • 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,
  • 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
  • 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.
  • 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.
  • the present invention involves a catalytic reforming zone, a separation zone and a particular saturate cracking zone.
  • 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.
  • the overall process includes a dehydrogenation reaction zone to produce the olefinic hydrocarbons utilized in the alkylation reaction zone.
  • the naphtha boiling range charge stock to the catalytic reforming zone may be derived from a multitude of sources.
  • 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.,
  • crystalline aluminosilicates such as the faujasites, or mordenite, or combinations of alumina with the various crystalline aluminosilicates.
  • 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.
  • 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.
  • the reactions effected therein are conducted at a relatively low operating severity.
  • 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.
  • the 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.
  • 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.
  • 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.
  • 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.
  • 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.
  • 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.
  • 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.
  • 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.
  • 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.
  • 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.
  • the charge may contain the pentane/hexane paraffins, a
  • 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.
  • the cracking of the paraffinic raffmate produces relatively large quantities of butanes, which butane concentrate is rich in isobutanes.
  • the hydrocracking reaction zone is referred to herein as l-cracking.
  • 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.
  • 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.
  • 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.
  • 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.
  • 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.
  • 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.
  • 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.
  • alumina and silica 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.
  • 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.
  • 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.
  • the noble metal component generally comprises from about 0.01 percent to about 2.0 percent by weight of the final composite.
  • 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.
  • 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.
  • 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.
  • 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.
  • the-finalcatalytic composite does not containany freeFriedel-Crafts metal halide.
  • 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.
  • 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.
  • Friedel-Crafts metal halides may be utilized, but not necessarily with equivalent results.
  • 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.
  • 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.
  • diluents such as inert gases.
  • the particularly preferred technique involves the sublimation of a metal halide directly to react with the alumina
  • the reaction product can result from a 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
  • halide-containing compounds include CCI SCl,, SOCl PCl POCl etc.
  • the catalytic composite Prior to its use, the catalytic composite may be subjected to a substantially water-free reduction technique.
  • 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.
  • the maximum catalyst bed temperature is maintained in the range of about 300 F. to about 480 F.
  • conventional quench streams either normally liquid, or normally gaseous and introduced at one or more intermediate loci of the catalyst bed, is contemplated.
  • the product effluent from the hydrocracking reaction zone is predominantly butanes, the greater proportion of which constitute the various isobutanes.
  • the hydrocracking reaction zone is herein referred to as l-cracking, the I alluding to isomer production.
  • 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.
  • 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.
  • a fluid catalytic cracking unit or a thermal cracking unit.
  • at least a portion of the isobutane concentrate is subjected to dehydrogenation in a dehydrogenation reaction zone to produce the alkylatable olefinic hydrocarbons.
  • 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.
  • 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.
  • 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.
  • 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.
  • 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.
  • 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).
  • 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.
  • 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.
  • dehydrogenation 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.
  • 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.
  • 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.
  • the high selectivity and relatively low conversion, in the dehydrogenation zone are advantageous.
  • 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.
  • 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.
  • 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.
  • 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:
  • 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.
  • 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.
  • 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.)
  • 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.
  • 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.
  • 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.
  • FIG. 2 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.
  • 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.
  • 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.
  • Isomerization zone 17 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.
  • 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:
  • 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:
  • 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.
  • hydrocracking catalytic composite comprises from about 0.1 percent to about 2.0 percent by weight of a platinum or palladium component.
  • hydrocracking catalytic composite comprises from about 1.0 percent to about 10.0 percent by weight of a nickel component.

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Cited By (8)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US3873439A (en) * 1973-02-26 1975-03-25 Universal Oil Prod Co Process for the simultaneous production of an aromatic concentrate and isobutane
US5135639A (en) * 1990-05-24 1992-08-04 Uop Production of reformulated gasoline
US5294328A (en) * 1990-05-24 1994-03-15 Uop Production of reformulated gasoline
US5536692A (en) * 1994-12-02 1996-07-16 Phillips Petroleum Company Isomerization catalyst and use thereof in isomerization of saturated hydrocarbons
US5543374A (en) * 1994-11-15 1996-08-06 Phillips Petroleum Company Isomerization catalyst and use thereof in alkane/cycloalkane isomerization
US5707918A (en) * 1995-08-29 1998-01-13 Phillips Petroleum Company Hydrocarbon isomerization catalyst blend
US6746495B2 (en) * 2000-10-24 2004-06-08 Exxonmobil Research And Engineering Company Method for controlling deposit formation in gasoline direct injection engine by use of a fuel having particular compositional characteristics
WO2021237034A1 (en) * 2020-05-22 2021-11-25 Exxonmobil Research And Engineering .Company High naphthenic content naphtha fuel compositions

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US2908628A (en) * 1956-06-28 1959-10-13 Sun Oil Co Hydrocarbon conversion
US3114696A (en) * 1958-10-03 1963-12-17 Socony Mobil Oil Co Inc Upgrading of naphthas
US3649520A (en) * 1970-03-13 1972-03-14 Mobil Oil Corp Production of lead free gasoline
US3658690A (en) * 1970-03-13 1972-04-25 Mobil Oil Corp Gasoline upgrading
US3660520A (en) * 1970-02-18 1972-05-02 Exxon Research Engineering Co Integrated oxydehydrogenation and alkylation process
US3692666A (en) * 1970-09-21 1972-09-19 Universal Oil Prod Co Low pressure,low severity hydrocracking process

Patent Citations (6)

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Publication number Priority date Publication date Assignee Title
US2908628A (en) * 1956-06-28 1959-10-13 Sun Oil Co Hydrocarbon conversion
US3114696A (en) * 1958-10-03 1963-12-17 Socony Mobil Oil Co Inc Upgrading of naphthas
US3660520A (en) * 1970-02-18 1972-05-02 Exxon Research Engineering Co Integrated oxydehydrogenation and alkylation process
US3649520A (en) * 1970-03-13 1972-03-14 Mobil Oil Corp Production of lead free gasoline
US3658690A (en) * 1970-03-13 1972-04-25 Mobil Oil Corp Gasoline upgrading
US3692666A (en) * 1970-09-21 1972-09-19 Universal Oil Prod Co Low pressure,low severity hydrocracking process

Cited By (10)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US3873439A (en) * 1973-02-26 1975-03-25 Universal Oil Prod Co Process for the simultaneous production of an aromatic concentrate and isobutane
US5135639A (en) * 1990-05-24 1992-08-04 Uop Production of reformulated gasoline
US5294328A (en) * 1990-05-24 1994-03-15 Uop Production of reformulated gasoline
US5543374A (en) * 1994-11-15 1996-08-06 Phillips Petroleum Company Isomerization catalyst and use thereof in alkane/cycloalkane isomerization
US5639933A (en) * 1994-11-15 1997-06-17 Phillips Petroleum Company Isomerization catalyst and use thereof in alkane/cycloalkane isomerization
US5536692A (en) * 1994-12-02 1996-07-16 Phillips Petroleum Company Isomerization catalyst and use thereof in isomerization of saturated hydrocarbons
US5707918A (en) * 1995-08-29 1998-01-13 Phillips Petroleum Company Hydrocarbon isomerization catalyst blend
US6746495B2 (en) * 2000-10-24 2004-06-08 Exxonmobil Research And Engineering Company Method for controlling deposit formation in gasoline direct injection engine by use of a fuel having particular compositional characteristics
WO2021237034A1 (en) * 2020-05-22 2021-11-25 Exxonmobil Research And Engineering .Company High naphthenic content naphtha fuel compositions
US11390820B2 (en) 2020-05-22 2022-07-19 ExxonMobil Technology and Engineering Company High naphthenic content naphtha fuel compositions

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IT961983B (it) 1973-12-10
JPS4848503A (de) 1973-07-10
FR2148636A1 (de) 1973-03-23
FR2148636B1 (de) 1975-03-07
CA969885A (en) 1975-06-24
DE2239282A1 (de) 1973-03-01
DE2239282C3 (de) 1975-08-21
DE2239282B2 (de) 1975-01-16

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