US3143490A - Hydrocarbon conversion process to produce isoparaffins from olefins - Google Patents

Description

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Herm erw United States Patent O 3,143,490 HYDROCARBON CONVERSION PROCESS T PRO- DUCE ISOPARAFFINS FROM OLEFINS Harry M. Brennan, Hammond, and Clifton G. Frey and Herman S. Seelig, Valparaiso, Ind., assignors to Standard Oil Company, Chicago, Ill., a corporation of Indiana Filed July 27, 1961, Ser. No. 127,272 Claims. (Cl. 208-49) This invention relates to a process for the production of parafiinic branched chain hydrocarbons. More particularly, it relates to a hydrocarbon conversion process for converting light normal parains to light isoparafns boiling withing the gasoline range.

Light isoparaffns are desirable components of gasoline blends because of their desirable octane ratings. Such isoparaiiins have markedly higher leaded motor octane ratings than either the corresponding normal paraffins or olefins of the same carbon number. Consequently, there is a definite advantage for converting light normal paraffns to isoparafiins. In order to meet the requirements for the production of high octane gasolines, it is desirable for petroleum refiners to isomerize the light normal paraffins in refinery streams such as light virgin naphtha which contains primarily C5 to C7 paraffns.

Processes are presently available for isomerizing light parains. However, these known processes have serious disabilities. The high temperature isomerization processes operate in the range of unfavorable thermodynamic equilibrium so that relatively low yields of isoparafhns are produced. The low temperature processes use corrosive catalyst systems which increase the capital investment required for a plant and also present difhculties during operation.

The present invention provides a novel process for increasing the chain-branching in normal parafiins which process results in high yields of isoparafiins without the diiiiculties encountered in presently available isomerization processes. Brieiiy, according to the present invention, normal parafiinic hydrocarbons are dehydrogenated to olefins and the resulting olefinic hydrocarbons are contacted with a composite catalyst having an isomerization activity and a hydrogenation activity in the presence of a hydrogen affording gas under isomerization-hydrogenation conditions, the composite catalyst comprising of a solid acidic component and a metallic hydrogenation component wherein the activities of the catalyst are controlled to produce under the conditions of the process a parafiinic product containing more branched parafiins than the parafiin isomer equilibrium amount at the operating temperature.

Various petroleum refinery streams containing substantial quantities of straight chain paraflins can be utilized as feedstocks for the present process. A particularly desirable feed is light virgin naphtba which contains primarily straight chain pentanes, hexanes and heptanes; although other suitable hydrocarbon streams containing substantial amounts of light paraffins also may be utilized.

In the operation of the process, the feed is first dehydrogenated to an olefinic hydrocarbon by a suitable dehydrogenation reaction. The dehydrogenation may be accomplished by contacting the parafhnic feed under dehydrogenation conditions with a dehydrogenation catalyst such as nickel, copper, palladium black, chromia, molybdena, and vanadia, either alone or composited with alumina, carbon, silica gel, etc. Typically, a chromia on alumina catalyst is employed and the feed is contacted with the catalyst at elevated temperature ranging from about 1000 F. to about 1200 F., preferably in the range of about l050 F. to 1075 F. Advantageously, a low operating pressure is employed, preferably in the range from about 3,143,490 Patented Aug. 4, 1964 rice 2 to about 4 p.s.i.a. The space velocity employed will range from about 200 to about 2000 volumes of gas per hours per volume of catalyst.

The dehydrogenation step can also be conducted by first catalytically treating the parafiinic feed with a hydrogen-accepting light olefin such as ethylene or propylene under hydrogen transfer conditions which will yield an olefinic naphtha. Typically, a light virgin naphtha stream plus the olefin, at a feed to olefin ratio ranging from about 1:1 to 1:100, is introduced into a reactor at a temperature from about 400 F. to about 1200 F., preferably 800 F. to 1000 F. The feed and olefin are contacted with a known hydrogen transfer catalyst such as silica-alumina, iiuorided-alumina, silica-magnesia, acid treated clay or the like, preferably silica-alumina, at a space velocity (LHSV) ranging between about 0.1 to 5 volumes of feed (as liquid) per hour per Volume of catalyst, preferably 0.25 to 2 LHSV. The operating pressure will range from atmospheric to about 1000 p.s.i.g., and preferably, the pressure is less than about 200 p.s.i.g. A hydrogen to hydrocarbon mole ratio from about 0 to 5, and preferably 0.5 to 2, moles of hydrogen per mole of hydrocarbon feed is employed.

Subsequently, the resulting oleiinic hydrocarbon from the dehydrogenation reaction is contacted with a second catalyst in the presence of a hydrogen-affording gas such as substantially pure hydrogen, catalytic reformer recycle gas or another gas stream containing sufficient available hydrogen for olefin saturation. The catalyst employed in this step of the process is a composite catalyst having a hydrogenation activity and an isomerization activity and is comprised essentially of a solid acidic component and a metallic hydrogenation component. The relative activities of the catalyst are controlled by heretofore discovered means to produce a converted material containing more branched paraffins than the paraffin isomer equilibrium amount at the operating temperature.

The olefinic material from the dehydrogenation step is introduced into the isomerization-hydrogenation reaction zone Where it is contacted with the balanced catalyst in the presence of at least sufficient hydrogen for olefin saturation. Advantageously, a hydrogen-affording gas such as substantially pure hydrogen, catalytic reformer recycle gas or another gas stream containing sufficient available hydrogen for olefin saturation is introduced into the reaction zone with the charge. The minimum amount of hydrogen required will be the stoichiometric amount required for olefin saturation, and this will vary according to the nature of the olefim'c charge. Preferably an excess` above, eg., about 1200 s.c.f. per barrel of C5 mono-olefin,

about 960 s.c.f. per barrel of C7 mono-olefin, etc. Excess hydrogen is removed in the reactor efiiuent and may be recycled to the reaction zone.

The isornerization-hydrogenation reaction zone is operated under conditions promoting the chain branching and saturation of the olefnic charge. An elevated pressure is employed, ranging to about 3000 p.s.i.g. or more. Advantageously, the pressure is in the range of slightly superatrnospheric to 1500 p.s.i.g. and preferably is about 250 to 750 p.s.i.g. for a fixed bed operation. An elevated temperature in the range of about 200 F. to 1000 F. is employed, the operating temperature being determined by the nature of the catalyst employed. The process may be conducted in either the vapor phase, the liquid phase, or a mixed vapor-liquid phase. Catalyst activity, the nature of the material charged to the reaction zone, pressure and other operating conditions will affect the selection of the operating temperature. For example, when contacting a C4 to C7 olenic material with a suliided nickel on silica on alumina catalyst, the temperature typically is in the range of about 400 F. to 750 F. and preferably is about 450 F. to 650 F. Liquid hourly space velocities (LHSV) of from about 0.1 to 50 volumes of hydrocarbon (as liquid) per hour per volume of Vcatalyst are employed, optimally between about 0.1 and 10, with a preferred rate being about 0.5 to 3 LHSV for a ixed'bed operation.

The catalyst employed in the isomerization-hydrogenation zone possess isomerization activities and hydrogenation activities, with these activities being controlled to produce under the defined conditions of the process, e.g., temperature, pressure, space velocity, the presence of herein defined modifying materials, etc., a more branched parainic product than the paraffin isomer` equilibrium amount at the operating temperature. The catalyst is a composite catalyst comprising a solid acidic component and a metallic component and may be produced by a variety of methods such as, co-gelling the components, supporting one component on the other, mixing the components and pelleting to size, impregnating with another component either prior to use or during processing, etc., or a combination of the above methods. Known hydrogenation catalysts are employed in carrying out the process of the invention. Catalysts such as the metals of Group VIII of the periodic table, particularly nickel, platinum, palladium and cobalt can be incorpoporated in the present catalyst. Also, other hydrogenation catalysts, such as the compounds, i.e., the oxides and/or sullides, of the metals of Groups VB and VIB of the periodic table, particularly tungsten, chromium and vanadium may be employed. In addition to the above-named catalysts, combination of metals and/or metal compounds having hydrogenation activity may be employed in the present process. In the instance where the hydrogenation catalyst utilized imparts a very high hydrogenation activity, compared to the isomerization activity, to the composite catalyst the hydrogenation activity can be reduced to a suitable level by means hereinafter described. Examples of the above are nickel and platinum on silica-alumina, wherein small amounts of sulfur, arsenic, antimony, bismuth, phosphorous, selenium, tellurium, copper, lead, mercury, silver, etc. are employed to control the hydrogenation activities. In other cases, the metallic hydrogenation constituent may have an inherently low activity relative to the isomerization activity, and such catalyst can be employed without the use of activity control-affording materials. For instance, a metallic hydrogenation component such as vanadia may be employed with a solid acidic component such as silica-alumina.

Such metals as mentioned above are employed as the metal constituent of the metallic component of the catalyst in varying amounts to give the desired catalyst hydrogenation activity relative to the isomerization ability of the solid acidic component. Typically, from about 0.1 weight percent to about 30 weight percent of nickel and about 0.1 to 2% of platinum are incorporated in a catalyst employing a silica-alumina support. It is understood that the catalyst activities of the various metallic components vary and that the proper operating temperature must be selected for each to attain the desired degree of activity.

Various solid acidic components, known to posses the ability to promote the isomerization of hydrocarbons and to increase the degree of chain branching therein, are incorporated in the catalyst. Preferably, these materials are porous, having a high surface area of about 100-600 square meters per gram, and are employed as a support with the metallic component disposed thereon. In

general, it is necessary to provide an acidity, which under the conditions of the process, is not suicient to isomerize para'ins at a substantial rate, but is suliicient to promote skeletal isomerization of the oleins. For example, commercially available acidic cracking catalysts such as silica-alumina catalyst has the ability to promote olefin isomerization under the proper temperature conditions.`

Other acidic components having suicent isomerization activities are uorided alumina, synthetic or natural silicates and other solid materials which posses the necessary acidity when incorporated in the catalyst to provide a relatively fast rate of olen isomerization with respect to hydrogenation rate. genation and isomerization are properly balanced, the converted material has higher than equilibrium /n paraffin ratio, which may be as high as about 30 under optimum process conditions.

The preferred silica-alumina component of the catalyst can be a naturally occurring mineral such as montmorillonite clay or a synthetic material, or a combination of these. Preferably, an artificial aluminosilicate, such as one of the commercially available cracking catalyst is utilized as a support. Such catalysts are generally made` by coating a silica sol with alumina and the alumina portion of the support can vary from about 5 to aboutV For example, the so-called high 40 weight percent. alumina cracking catalyst containing about 20 to 30 weight percent A1203 is very effective.

Normally only small amounts of the aforementioned activity-control-aording substance are required to properly balance the isomerization-hydrogenation catalyst. The total amount required will be dependent upon the total amount of the particular hydrogenation metal incorporated in the catalyst and upon its chemical form. Further, the activity-control-aifording substances may be employed either singly or in combination, and they may be incorporated in the catalyst during its manufacture or subsequently during use, e.g., organic metallic compound additions made to the feed stock.

Typically, when a sulfded nickel on silica-alumina catalyst is employed, the catalyst can be pre-sulided by contacting a nickel on silica-alumina catalyst base with a sulfur-affording gas at elevated temperature, and the activity balance may then be maintained byintroducing at least about 0.01 weight percent available sulfur, and preferably at least about 0.1 weight percent, in the feed to the reaction zone to prevent reduction of the suliided catalyst.

When, for example, a normally solid element of Group VA of the periodic table, such as arsenic or antimony, is employed `as the -activity-control-aording substance, about 0.01 to 5, and preferably about 0.1 to l moles of these elements per mole of hydrogenation metal, such as nickel or platinum, is employed. Likewise, small amounts of lead, mercury, silver or copper may be employed as the activity-control-afording substance. Normally, about 0.03 to 5, and preferably about 0.1 to 1 moles of these latter elements per mole of the hydrogenation component are incorporated in the catalyst.

The above-mentioned isomerization-hydrogenation catalyst can be readily prepared -by impregnating a solid acidic support, such as a silica-alumina cracking catalyst, with an organic or inorganic solution of a hydrogenation component, such as a nickel lacetate -o-r chloroplatinic acid, drying `and calcining at about l000 F. Following this, the catalyst can be impregnated with an organic or inorganic compound of the activity-control-alfording substance, such as triphenyl arsine, hydrogen sulfide, or lead nitrate. In most instances it is desirable to treat the twice-impregnated catalyst with hydrogen at an elevated temperature, i.e., about 850 F., to reduce the catalyst.

The present invention will be better understood by a reading of the following specific example of the operation of the process according to the invention and by reference to the accompanying drawing which diagrammatically illustrates the ow scheme employed -in such a process.

Referring to the drawing, a light virgin naphtha feed boiling in the range of about F. to 300 F. and comprising substantial amounts of C5-C7 normal parans is passed by way of line 11 to a first dehydrogenation re- When the activities for hydro-V action zone 13, which dehydrogenation zone is of the hydrogen transfer type hereinabove mentioned. Propylene, ethylene or a propylene-ethylene mixture is introduced 4into the dehydrogenation zone by Way of line 12. The parafn to olefin ratio within the dehydrogenation zone is maintained at about 1:5. The hydrogen transfer dehydrogenation zone contains a fixed bed of silicaalumina catalyst. The material charged to this reaction zone is contacted With the catalyst therein at a temperature of about 950 F. and a pressure of about 100 p.s.i.g. The feed -is charged to the reaction zone at a rate of about 1 volume of feed (as liquid) per hour per volume of catalyst. The eiuent from the zone 13 is passed by way of line 16 to a separation zone 17, in this case a fractionator, wherein it is separated into at least ya light gaseous fraction rich in propane and/or ethane and a heavier oleiinic naphtha fraction. Desirably, the C2-C3 rich light fraction is passed by Way of line 18 to a second dehydrogenation Zone 19 wherein it is contacted with a chromia on alumina dehydrogenation catalyst at a temperature of about 1075 F., a pressure of about 2 p.s.i.a. and a space velocity of 1000 volumes of gas per hour per volume of catalyst. Effluent from zone 19, rich in ethylene and/ or propylene, is passed via lines 20 and 12 back to the dehydrogenation zone 13 for further use therein.

The heavier oleiinic naphtha from the separator 17 is then passed by way of line 21 to the isomerization-hydrogenation reactor 22 which contains a bed of arsenided nickel on silica-alumina catalyst. This catalyst is a 5 percent nickel-2.5 percent arsenicsilica-alumina composite. The yoleiinic naphtha and la hydrogen-rich gas supplied by 'way of line 23 are contacted in the zone 22 under isomerization-hydrogenation conditions which include a temperature of about 600 F., a pressure of 1000 p.s.i.g., a liquid hourly space velocity of volumes of hydrocarbon per hour per volume of catalyst and a hydrogen to hydrocarbon ratio of about 11,000 standard cubic feet of hydrogen per barrel of feed. The resulting eiuent is withdrawn from zone 22 via line 24 and passed to a high pressure separator 25 wherein a light gaseous fraction rich in hydrogen is separated therefrom. This hydrogen-rich gaseous fraction is then removed by Way of line 26 and recycled to the zone 22. The remaining fraction, rich in isoparains and containing unconverted normal paraiins is passed by way of line 27 to a separation zone 28 which typically employs `a bed of a crytsalline zeolite molecular sieve material, e.g., 4 or 5 Angstrom molecular sieve, to separate isoparains from normal parains. This latter separation step is operated in a conventional manner (e.g., see Petroleum Rener, volume 39, 6, pp. 125-132 (1960)), and normal paraflins substantially boiling under 200 F. are recycled by Way of line 29 and line 11 to the dehydrogenation zone 13. Isoparains boiling substantially Within the C5-C, range are Withdrawn from zone 2S by Way of line 31.

Alternately, the dehydrogenation zone 13 may employ a dehydrogenation catalyst such as chromia on alumina. With this type operation, of course, the light oleiins are not introduced into the dehydrogenation zone. The dehydrogenation zone 13 typically is operated under conditions which include a temperature of 1075 F., a pressure of 2 p.s.i.a. and Ia space Velocity of about 10,000 volumes of gas per hour per volume of catalyst. The effluent withdrawn through line 16 is passed directly by way of lines 32 and 21 to the isomerization-hydrogenation zone 22. As described above the hydrogen-rich fraction is separated from the eiuent from zone 22 and recycled, and the heavier liquid fraction is separated into a normal p-araln fraction recycled to the dehydrogenation zone 13 and an isoparainic product fraction.

The periodic table hereinabove referred to is the periodic table of elements contained in the volume College Chemistry, 2nd ed., by Paul R. Frey, Prentice- Hall, Inc., 1958. It is to be understood the activitycontrol-affording substances referred to above may be incorporated into the isomerization-hydrogenation catalyst in Various chemical combinations, such as compounds, alloy or in elemental form, with the other catalyst components to provide a catalyst capable of yielding a greater amount of isoparatiins, under the conditions of the process than the isoparain equilibrium amount at a comparable temperature, although the various forms are not necessarily equivalents.

An alternate but not necessarily equivalent method for conducting the combination hydrogen transfer and isomerization-hydrogenation process is to employ a single reactor having a rst hydrogen transfer stage and a second isomerization-hydrogenation stage. The iirst stage contains a hydrogen transfer catalyst and is operated at conditions as described above, and the second stage contains an isomerization-hydrogenation catalyst and is operated as hereinabove described. The feed is introduced through an inlet into the first stage and passes downstream into the second stage. The total effluent may be withdrawn from the second stage and the isoparailin-rich product separated therefrom.

From the foregoing description of the invention, variations and modifications in the operation of the process will be apparent to the skilled artisan, and, as such, fall Within the spirit and scope of the invention. For example, net hydrogen produced from the dehydrogenation may be utilized in the isomerization-hydrogenation zone 22.

What we claim is:

1. A hydrocarbon conversion process which comprises contacting a normal parainic hydrocarbon charge in a first reaction Zone With a first catalyst under dehydrogenation conditions to produce a normal olelinic hydrocarbon product, contacting said oleiinic product in the second reaction zone with a second catalyst having an isomerization activity and a hydrogenation activity in the presence of a hydrogen-affording gas under elevated pressure and at a temperature in the range of about 200 F. to 1000 F., said second catalyst comprising a solid acidic component and a metallic hydrogenation component wherein the activities of said second catalyst components are controlled to produce under the conditions of the process a paraiinic product containing more branched parafns than the paralin isomer equilibrium amount at the operating temperature.

2. The process of claim 1 wherein said normal parafnic charge is contacted in said first reaction zone with a light olefin under hydrogen transfer conditions to effect the dehydrogenation of said paratiinic charge to a corresponding oleiinic product and wherein said tirst catalyst is an acidic hydrogen transfer catalyst.

3. The process of claim 1 wherein said second catalyst comprises from about 0.5 to 5 percent nickel and from about 0.1 to 1 mole of arsensic per mole of said nickel on a silica-alumina support containing from about 5 to 40 Weight percent alumina.

4. The process of claim l wherein said second catalyst comprises from about 0.5 to 5 percent nickel and from about 0.1 to l mole of lead per mole of said nickel on a silica-alumina support containing from about 5 to 40 Weight percent alumina.

5. The process of claim l wherein said second catalyst comprises from about 0.5 to 5 percent nickel and from about 0.1 to 1 mole of mercury per mole of said nickel on a silica-alumina support containing from about 5 to 40 weight percent alumina.

6. The process of claim l wherein said normal parafnic hydrocarbon is a light virgin naphtha.

7. The process of claim 2 wherein said hydrogen transfer conditions include a temperature in the range of about 800 F. to 1000 F., a pressure in the range of atmospheric to 200 p.s.i.g., a space Velocity in the range of about 0.25 to 2 liquid volumes of feed per hour per volume of catalyst and a hydrogen to hydrocarbon mole ratio in the range of about 0.5 to 2 moles of hydrogen per mole of hydrocarbon feed; and said isomerization-hydrogenation conditions include a temperature in the range of about 450 F. to 650 F., a pressure in the range of about 250 to 750 p.s.i.g., a liquid hourly space velocity inthe range of about 0.1 to 10 volumes of hydrocarbon per hour per volume of catalyst and a hydrogen to hydrocarbon ratio of at least about 1200 standard cubic feet of hydrogen per barrel of said olen..

8. A hydrocarbon conversion process which comprises contacting a normal parainic hydrocarbon charge boiling in the gasoline range with a light olenic hydrocarbon under hydrogen transfer conditions in the presence of an acidic hydrogen transfer catalyst in a dehydrogenation zone; fractionating the resulting eliluerit from said dehydrogenation zone into at least a light fraction rich in paraiins and a heavy oleiinic fraction; contacting said light fraction with a dehydrogenation catalyst under dehydrogenation conditions; recycling the resultingV dehydrogenated light parains to said hydrogen transfer zone; contacting said heavy olenic fraction with a catalyst comprising a Group VIII hydrogenation metal, a solid acidic component and a small amount of an element selected from the group consisting of sulfur, arsenic, antimony, bismuth, phosphorous, selenium, tellurium, copper, lead, mercury and silver in the presence of a hydrogen-affording gas under elevated pressure and at a temperature in the range of about 400 F. to 750 F.; and recovering a product rich in isoparains having a carbon number distribution corresponding to said hydrocarbon charge.

9. A hydrocarbon conversion process which comprises contacting a gasoline boiling range hydrocarbon charge containing substantial normal parains with a C2-C3 olenic hydrocarbon in the presence of an acidic hydrogen transfer catalyst in a hydrogen transfer zone at a temperature in the range of about 800 F. to 1000 F., a pres- 8 sure 4in the range of atmospheric to y200 p.s.i.g., a space velocity in the range of about 0.25 to 2 liquid volumes of feed per hour per volume of catalyst and a hydrogen to hydrocarbon mole ratio in the range of about 0.5 to 2 moles of hydrogen per mole of hydrocarbon feed; fractionatng the resulting ellluent from said hydrogen transfer zone into at least a light'fraction rich in para'ns and an olenic fraction; contacting said light fraction with a dehydrogenation catalyst under dehydrogenation conditions; recycling the resulting dehydrogenated light parafns to said hydrogen transfer zone; contacting said heavy olenic fraction with a catalyst comprising from about 0.5 to about 5 Weight percent nickel and from about 0.1 to l mole of arsenic per mole of nickel on a silica-alumina cracking catalyst support in the presence of at least about 1200 standard cubic feet of hydrogen per barrel of said olenic fraction, at a temperature in the range of about 450 F. to 650 F., a pressure in the range of about 250 to 750 p.s.i.g., and a liquid hourly space velocity the range of about 0.1 to l0 volumes of hydrocarbon per hour per volume of catalyst; and recovering a product n'ch in isoparaflin having a carbon number distribution corresponding to said hydrocarbon charge.

10. The process of claim 9 wherein said hydrocarbon charge is a light virgin naphtha.

References Cited in the le of this patent UNITED STATES PATENTS 2,333,625 Angell Nov. 9, 1943 2,456,672 Block et al. Dec. 21, 1948 2,880,249 Raley et al Mar. 3l, 1959 v2,897,137 Schwarzenbek July 28, 1959 3,018,244 Stanford et al. Ian. 23, 1962

Claims (1)

1. 9. A HYDROCARBON CONVERSION PROCESS WHICH COMPRISES CONTACTING A GASOLINE BOILING RANGE HYDROCARBON CHARGE CONTAINING SUBSTANTIAL NORMAL PARAFFINS WITH A C2-C3 OLEFINIC HYDROCARBON IN THE PRESENCE OF AN ACIDIC HYDROGEN TRANSFER CATALYST IN A HYDROGEN TRANSFER ZONE AT A TEMPERATURE IN THE RANGE OF ABOUT 800*F. TO 1000F., A PRESSURE IN THE RANGE OF ATMOSPHERIC TO 200 P.S.I.G., A SPACE VELOCITY IN THE RANGE OF ABUT 0.25 TO 2 LIWUID VOLUMES OF FEED PER HOUR PER VOLUME OF CATALYST AND A HYSROGEN TO HYDROCARBON MOLE RATION IN THE RANGE OFABOUT 0.5 TO 2 MOLES OF HYDROGEN PERMOLE OF HYDROCARBON FEED; FRACTIONATING THE RESULTING EFFLUENT FROM SAID HYDROGEN TRANSFER ZONE INTO AT LEAST A LIGHT FRACTION RICH IN PARAFFINS AND AN OLEFINIC FRACTION; CONTACTING SAID LIGHT FRACTION WITH A DEHYDROGENATION CATALYST UNDER DEHYDROGENATION CONDITIONS; RECYCLING THE RESULTING DEHYDROGENATED LIGHT PARAFFINS TO SAID HYDROGEN TRANSFER ZONE; CONTACTING SAID HEAVY OLEFINIC FRACTION WITH A CATALYST COMPRISING FROM ABOUT 0.5 TO ABUT 5 WEIGHT PERCENT NICKEL AND FROM ABOUT 0.1 TO 1 MOLE OF ARSENIC PERMOLE OF NICKEL O A SILICA-ALUMINA CRACKING CATALYST SUPPORT IN THE PRESENCE OF AT LEAST ABOUT 1200 STANDARD CUBIC FEET OF HYDROGEN PER BARREL OF SAID OLEFINIC FRACTION, AT A TEMPERATURE I THE RANGE OF ABOUT 450*F. TO 650*F., A PRESSURE IN THE RANGE OF ABOUT 250 TO 750 P.S.I.G., AND A LIQUID HOURLY SPACE VELOCITY IN THE RANGE OF ABOUT 0.1 TO 10 VOLUMES OF HYDROCARBON PER HOUR PER VOLUME OF CATALYST; AND RECOVERING A PRODUCT RICH IN ISOPARAFFIN HAVING A CARBON NUMBER DISTRIBUTION CORRESPONDING TO SAID HYDROCARBON CHARGE.

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