US3165461A - Neohexane production - Google Patents

Description

Jan. 12, 1965 D. E. WADE ETAL.

NEOHEXNE PRODUCTION 2 Sheets-Sheet l Filed Aug. 7, 1959 S nu Jan. 12, 1965 D. E. WADE ETAL 3,165,461

NEOHEXANE PRODUCTION 2 Sheets-Sheet 2 Filed Aug. 7, 1 959 ATTRNEY United States Patent O prix f? 1 1,1535 61 Nlliih XAilrt PRGEUCHN Dennis E. Wade and ii'iaynard L Anderson, Texas City, Tex., assigncrs, by rnesne assignments, to Standard @il Company, iiicago, iii, a corporation of indiana Filed Aug. 7, i959, Ser. No. 832,2@1 4 Ciaims. (Si. 26S- 64) rThis invention relates to high anti-knock gasoline production and it pertains more particularly to improvements in hexane feed stoel; preparation and hexane isomerization for the production chiey of neohexane, i.e., 2,2- dimethylbutane.

Heretofore in the production of neohexane by aluminum chloride isomerization (US. 2,446,608)- a liquid catalyst complex was employed, temperatures upwards of 200 F. were used commercially to obtain feasible reaction rates, corrosion problems were vexatious and neohexane was produced in relatively small conversions per pass which made product fractionation and recycle quite expensive. The object of this invention is to minimize or avoid such problems and diiiiculties. A further object is to provide an improved method and means for obtaining large quantities of n-hexane-methylpentane charge containing lessv than about five volume per cent hcptanes and substantially free from aromatics, sulfur and other impurities that might be detrimental to aluminum chloride-on-adsorbent isomerization catalyst. Another object is to provide an improved liquid phase hexane isomerization system employing solid aiuminum chloride-on-adsorbent catalyst. Other objectives will be apparent as the detailed description of the invention proceeds.

Our invention makes it possible to produce commercially a non-leaded super premium gasoline having an octane number substantially higher than iGO. Available naphthas are split by a fractional distillation to obtain light fractions (C7 hydrocarbons and lighter) and heavy fractions (C7 hydrocarbons and heavier}. The heavy fractions are hydroformed with platinum-on-alumina catalyst at a severity to obtain a reformate having a clear octane number upwards of 90 and such that the heavy reformate produced will have an octane number substantially higher than 10G. rhe reformate is split by fractional distillation to obtain a Cq+ fraction which may have an octane number of about 10B-104 which is a major cornponent of the super premium gasoline. The Cf, and lighter fraction of the reformate is combined with the light virgin naphtha fraction and the mixture is depentanized and then dearomatized. The dearomatization is preferably by selective extraction when aromatica aredesire'd as an additional product, but unless the extraction is suhciently exhaustive to reduce the aromatic content to less than 0.2 volume percent the depentanized stock should be hydrogenated or otherwise treated to insure removal of not only aromatics but also sulfur compounds and any other component which might be deleterious to the isomerization catalyst. The dearomatized charge is then distilled to eliminate heptanes, which is facilitated by the absence of azeotrope-forming aromatics. A hexane stream is thus produced which is particularly suited for isomerization with our aluminum chlorideon-adsorbent catalyst.

Our preferred catalyst is an acid washed and calcined alumina such as the activated alumina of commerce or bauxite on the surface of which there is deposited at least ICC a monomolecular layer of aluminum chloride. When Porocel (an acid Washed and calcined bauxite) is employed as the adsorbent the amount of aluminum chloride deposited on it should be in the range of about 15-30 weight percent. Other supports such as silica, silica alumina, Ca3(PO4)2, Cal-W04, MnHPOa, etc., maybe used instead of adsorbent alumina but the latter is particularly advantageous both from the standpoint of cost and electiveness. Instead of subliming the aluminum chloride on the support (note Perry lsomerization of Light Hydrocarbons, Transactions of American institute of Chemical Engineers, volume 42 V(194), pages 639 et seq.), we meter aluminum chloride powder and powdered Porocel separately into a circulating, noncomplex forming liquid to form a slurry so that the aluminum chloride is more uniformly adsorbed and the AlCl3 is dispersed more nearly over the entire surface area. The liquid is preferably an aliquot portion of the isomerized product. Aluminum chloride-on-adsorbent catalyst is granular in form and may have a particle size in the range of about 4-40 mesh although about 8-40 mesh particles are preferred. The Al/lg is thus deposited on adsorbent at about to 200 F.

Preferably the purified hexane charge contacts the solid catalyst in a stage-countercurrent system wherein catalyst about to be discarded or regenerated is contacted with in- Corning charge and recycled hydrocarbons and the final treatment of the charge stream is with freshly prepared catalyst. ln each stage of our preferred system the weight ratio of hydrocarbon to catalyst in the contacting vessel is in the range of about 0.5 :1 to 2:1, preferably about 0.7:1 to 1:1. ln the stirred portion of each vessel there may be about 4 to S pounds catalyst per gallon of liquid. The liquid which is introduced to the settler in cach stage may contain about .3 to 4 pounds of catalyst per gallon. The catalyst stream leaving the base of the settlers may contain 8-12 pounds of catalyst per gallon while liquid withdrawn from the upper part of the settlers may have only about 0.1 pound catalyst per gallon. About 1GO-400 and preferably 20G-300 parts of catalyst is recycled in the stage for each part catalyst that is discharged and replaced by incoming catalyst. f

Many features of the previously known neohexane process (US. 2,443,608) may be employed, but generally speaking our improved process may be operated with less aluminum chloride activator andrlower hydrogen pressure. Promoter is stripped from the' product and ab-` sorbed in incoming charge so that less than 1% make-up HC1 is usually necessary. The preferred stripping gas is hydrogen from the hydroformer and the incoming charge may dissolve enough hydrogen to make introduction of extraneous hydrogen unnecessary. The charge should contain about 2-20% of naphthenes such as methylcyclopentane and/or cyclohexane.` The isomerization is ef-V fected in liquid phase at a temperature of 60-180 F. pref-V erably Btl-140 F. and ordinarily in the range of 100'- F. for a period of time which may range from approximately 0.5 to 5 hours.

The invention will be more clearly understood from the following description of a preferred example thereof read in conjunction with the accompanying drawings which form part of the speciication and in which:

FIGURE l is a schematic ow diagram of our irnproved neohexane production system showing the charge 3 preparation and super premium gasoline production; and Y FIGURE 2 is a schematic ilow sheet showing our hexane isomerization system in greater detail.

In a refinery employing two hydroformers (ultraformers see Petroleum Renner, 33 (9), 1623 (1954), and Petroleum Reiiner, 33 (4), 153-56 (1954)) each having a charge rate, including recycle of about 18,000 to 20,000 barrels per day and having about 12,000 barrels per day of straight run gasoline the total hexane distribu- '.Ihe `system shown in FIGURE l makes it possible toY utilize hexanes from these sources and the conversion of almostall of these hexanes to neohexane. About 25,000 barrels per day of naphtha charge from line is fractionated in column 11 to obtain about 18,500 barrels per day. of light naphtha which is Withdrawn through line 12` and about 6,500 of C7 and heavier hydrocarbons which are withdrawn by line 13 to hydroformer 14.

While l.the hydrofor'ming may be `eifected by anyknown Itechnique which is Aeffective for converting the charge substantially entirely to reformate of which the highest boiling 75% will have an octane number of 102-104, the preferred hydroforming technique is Ultraforming as described in the cited references. The reformate produced bythe hydroforming step is introduced by line 15 to splitting tower 16 wherein heavy reformate (C7 and heavier) is removed by line 1'7 as a bottoms fraction while light reformate (C7 and lighter) is taken overhead from the fractional distillation tower by line 18 and introduced to depentanizer tower 19 along with light virgin naphtha from line 12. Pentanes and lighter hydrocarbons are taken overhead from depentanizer 19 by line 20 and may be employed largely for blending into regulargasoline. Y The depentanized light Vnaphtha mixture is introduced by line 21 to dearomatiza-tion system 22 which may .be A an extraction Vsystem when benzene and/orV toluene is 'to Abe separately produced.V The aromatics removal may be effected by extraction with a glycol solvent yorother commercially available technique (note Udex` process, Petroleum Processing, March 1953, pp. 384-7) or by any other extraction system which is'adequatefor removing substantially all aromatics from the depentaniZed stream. A particularlydesirable aromatics removal system is described in U.S. 2,768,986. If'the extraction leaves more than .2 volume Varomatics in the depentanized stream additional aromatics may be re. moved. by adsorption,` adduct formation, hydrogenation or any, othertechniqueiknown to the art. InA this example upwards of about 2,000 barrrels per day of aromatic product is withdrawn as a separate product through line 23 although. it should be understoodthatat least va part ofthe aromatics -may be introduced by line 24 into the .sinner- ,premiumY gasoline.' The dearomatization of the hexane stream facilitates ,removal ofV C7 hydrocarbons from hexanes by avoiding v azeotrope formation which would unavoidably be `encounteredfin the presence of aromatics. The aromatic'free-hexanefheptane stream is introduced by line 25 to distillation tower 26 wherein the C, and heavier hydrocarbons are separated and returned by line 27 for introduction into hydroformer 14. The hexane stream which leaves the top of tower 26 by line 28 substantially free from aromatics, sulfur compounds, and other impurities Which wouldbe deleteriousV to isomerization'catalyst will usually contain an amount of napthenes in the rangeof about 2-20%. This hexane charge is the feed stock for isoinerization system 29 which is shown in greater detail in FlGURE 2. Hydrogen produced as make gas in the hydroformer 14 is passed by line 30 for use in the isomerization system. Make-up catalyst is introduced by line 31 in carrier liquid introduced by line 32, the slurry being passed by line 53 to the isomerization system. Spent catalyst is removed by line 34 and it may either be discarded or regenerated for further use.

The isomerized product stream is introduced by line 35 to a fractionation system diagrammatically represented by tower 36 from which the neohexane product is taken overhead by line 37, cyclohexane is removed as bottoms through line 33 While an intermediate boiling range methylpentane-n-hexane stream is removed by line 39 for recycle to the isomerization system 29. An ,aliquot portion of the isomerate stream from line 40 or a portion of the cyclohexanerstream may be used for preparing the catalyst as hereinabove described. The net amount ofV cyclohexane produced Vmay be withdrawn through line 42 kand returned to the hydroformer particularly if this stream contains appreciable amounts of methyl or` dimethyl cyclohexanes. The composition of the isomerized Ypu'oduct charge to fractionator. 36 by line 35 and the ycompositions of the overhead removed by line 37 and the sidestrearn plus bottoms removed by lines 39 and 33 may be substantially as follows:

From .the foregoing it will be observed that the neohexane product stream has a phenomenally high neohexane concentration and that by recycle of the other hexanes almost all of thehexanes can be converted to neohexane. NeohexaneY has a research blending octane numberV kof approximately 100 in heavy hydroforrnate soV that the total gasoline obtained by blending as illustrated in FIGURE l in this Vparticular example may have an octane Ynumber of the order of 103.v

Referringto FIGURE 2 the hexane stream from line 23V together with recycled hexanes from line 50 are introduced at' the upper part ofA absorber 51 wherein it picks up both HC1 and hydrogen introduced by line S2, Vunabsorbed hydrogen being vented from the system by line 53. Make-up promotor HC1 may be introduced by line 54 in amountsto `bring the total HC1 content of the charge to Vabout .1 to 1% e.g. about .2 weight percent. The promotor-containing charge which is saturated with hydrogen then passes Vby line 55 to reactor 56 which is a vesselV about9 feet in diameter and 25 feet in height p provided with a motor driven stirrerV 57 and with a 45 of slurry containing about 514 pounds per hour of catalyst.`

The amount of charge plusrrecycle introduced by line 55 being about 14,400 barrels per day. The 100 horse power stirrer 57 maybe driven at about 1500 to 2000V r.p.m. for obtaining intimate VVcontactof catalyst with charge. YAbout 21,110 barrels per day of liquid is withdrawn Yfrom Vthe upper quiescent Zone by pump `62 and introduced by line 63 by settler 64 which may be 8 feet in diameter and 8 feet ltall. The stream entering the settler may contain about 31/2 pounds of catalyst per gallon. The settled material leaving the base of the settler through line 65 at the rate of about 6,700 barrels per day may contain about to 11 pounds of catalyst per gallon and about 271/2 barrels per day of this stream will be discharged through line 66 while the rest of this stream is returned by line 67 to the stirred portion of reactor S6. The 14,400 barrels per day of liquid from the upper part of the settler is passed by line 68 to the next stage of the countercurrent contacting system, this stream containing only about .1 pound of catalyst per gall-on.

In this particular example we employ five separate contacting stages, i.e. an initial step, a final step, and three intermediate stages but it will be apparent that n-intermediate stages may be employed Where n is 0 'to 5 or more. ln the intermediate `stage incoming charge is introduced by line 68 and 69 to the stirred portion of reactor 70 into which make-up catalyst is introduced by line 71. A relatively quiescent liquid is pumped from the upper part of vessel 70 by pump 72 and line 73 to settler 74 from which thickened slurry is withdrawn through line 7S, a minor portion thereof being withdrawn through line 61 to reactor 56 and the major portion being recycled by line 76 to reactor 70. Each of the n-intermediate stages will be substantially the same and therefore only one need be described.

The partially isomerized hexane `stream from the settler of the last intermediate stage is introduced by lines 78 and 79 to mixing zone in the final reactor 80 into which freshly prepared catalyst is introduced by line 81. The liquid from the upper quiescent zone in Vessel 80 is introduced by pump 82 and line $3 into settler 84 from which the more concentrated catalyst slurry is returned by line 85 and S6 except for lthe 271/2 barrels per day which is passed by line 71 to reactor 70. The nal isomerized product is introduced by line 88 to clarifier 39 which may be a sand filter but which is preferably afilter bed formed from adsorptive alumina. Filter-clarified efuent, which is preferably denuded of ydissolved aluminum chloride as well as catalyst solids, is passed by line 90 to the upper part of stripper tower 91 at .the base of which hydrogen is introduced by line 92, the hydrogen preferably being make gas from the hydroformer. The HCl-enriched stripping gas which leaves the top of stripper 91 may be introduced through line 52"` at the base of absorber 51 so that the incoming hexane f charge may absorb substantially all of the HC1 contentftherein and may also dissolve a substantial amount of hydrogen.

Stripped isonrerate leaves: `the base of stripper 91 through line 94, is caustic washed in scrubber system 95, water washed in scrubber system 96 and it is preferably dried before being introduced into fractionating tower 97 from which the neohexane product `stream is removed overhead by line 9S (which corresponds with line 37 in FIGURE 31). The bottoms from distillation tower 97 are withdrawn through line 99 to fractionating tower 100 although at least a part of such bottoms may be recycled through lines 101 and 50, respectively. The overhead from tower 100 consists chiefly of methylpentanes and normal hexanes and is recycled by line 50 to absorber 51. A cyclohexane stream 102 is withdrawn from the base of tower 100.

An aliquot part of the isomerate may be introduced via line 104 to catalyst preparation vessel 105 into which acid-washed and calcined adsorptive alumina such as Porocel is introduced by feeder 106 and aluminum chloride is introduced by feeder 107. These feeders may be of the type described in U.S. 2,792,152 and the two materials are fed into vessel 10S in such proportions as to give about 1 part by weight of aluminum chloride to 3 parts by weight of alumina. The resulting slurry is withdrawn from vessel 105 by pump 108 and a portion rev aluminum cycled tothe base of vessel 105 through line 109 in order that the slurry may remain in uniform intimately mixed suspension at about to 130 F. while the aluminum chloride is being deposited on the adsorptive alumina. The resulting slurry is withdrawn through line 81 for providing make-up catalyst in vessel Si) at the rate of about 271/2 barrels per day.

In this example the temperature is preferably about 1Z0-130 F. in reaction vessel 56 and about 90-100" F. in reaction vessel 81B with a gradient therebetween although the temperature may be uniformly about 110 F. throughout. The reaction pressure is maintained at about 50 p.s.i.g.; the pressure should be suflicient to maintain liquid phase conversion conditions but unduly high hydrogen pressures retard desired isomerization. Generally speaking the weight space velocity may be about 0.2 part by weight of oil per hour per part by weight of catalyst employed and the catalyst to oil ratio may be of the order of 1:1 and the hold-up time may be in the range of 0.5 to 5 hours.

Catalyst withdrawn from the system through line 66 may be treated to recover the alumina. After drying the alumina may be reemployed in the preparation of catalyst. However, in many instances economics may favor simple disposal of spent catalyst before or after recovering and/ or hydrolyzing material deposited thereon.

While a preferred example of our invention has been described in suitable detail it should be understood that modifications and alternative operating conditions will be apparent from the foregoing description to those skilled in the art.

We Claim:

1. The method of converting naphtha to an aromatic product and a super premium motor fuel having a clear octane number substantially higher than 100 which :method comprises splitting naphtha into a light fraction containing some C7 and lighter hydrocarbons and a heavy fraction containing C7 and heavier hydrocarbons, hydroforming the heavy fraction under a severity to obtain reformate of which the higher boiling portion has an octane number of about 102-104, splitting the reformate into a light fraction containing C7 and lighter hydrocarbons and a separate heavy reformate fraction, depentanizing the combined light fractions, removing substantially all aromatics from the depentanized light fractions to obtain a dearomatized stream, dehexanizing the dearomatized stream to obtain a hexane stream and a heptane stream, recycling at least a part of the heptane stream to the hydroforming step, isomerizing the hexane stream which is substantially free from aromatics and which contains 2-20% naphthenes by contacting with chloride-on-adsorbent catalyst containing fabout 15-30% aluminum chloride under `a pressure sumcient to maintain liquid phase conditions at a ten era- -ture in the range of about 60-180 F. for a time to convert a large proportion of the normal hexane and methylpentanes to dimethylbutanes, splitting the isomerization product by fractional distillation to obtain a stream which is chiefly neohexane and a stream consisting essentially of methylpentanes and normal hexane, recycling the latter stream to the isomerization step and combining the neohexane stream with the heavy reformate fraction to produce super premium gasoline having a clear octane number substantially higher than 100.

2. The method of claim 1 which includes the steps of separately metering aluminum chloride and absorbent alumina into a slurrying zone, introducing a liquid into said slurry zone which is inert and unreactive with said aluminum chloride and pumping said slurry in a closed circuit to deposit aluminum chloride on the alumina support at a temperature in the range of about 90 to F.

3. The method of converting normal hexane and methylpentanes to neohexane which method comprises contacting a charge stream consisting essentially of hexanes other than neohexane and containing about 2-20% of naphthenes with a solid aluminum chloride-onadsorbent catalyst slurry in a plurality of stages each operated at a temperature in the range of about 60- 180 F., introducing a small amount of make-up catalyst slurry and a large amount of a charge stream-to a mixing zone of each stage and intimately mixing the catalyst slurry therein, passing slurry from the mixing Zone to a quiescent zone and passing liquid `from the quiescent zone to a settling zone in each stage, recycling a minor portion of the liquid and a major portion of the catalyst from the settling zone to the mixing zone in each stage, withdrawing a major portion of the liquid and a minor amount of catalyst slurry from the settler to the adjacent prior stage, withdrawing a minor amount of settled catalyst to the adjacent subsequent stage, and recycling in each stage more than 100 times the amount of catalyst which is continuously added thereto and withdrawn therefrom.

4. An improved countercurrent slurry catalyst contacting system comprising a series of stages wherein each stage comprises (l) a contacting vessel having a lower mixing zone and an upper quiescent zone, a stirrer in said mixing zone, a baffle in said vessel arranged to isolate said upper quiescent zone, said baiiie having an opening for receiving catalyst slurry from said mixing zone into said quiescent zone and means for returning settled catalyst slurry from said quiescent zone to said mixing zone,

(2) a settler exterior of said vessel,

(3) a conduit from said quiescent zone to said exterior settler,

(4) branched conduit fromthe base of said exterior settler to recycle by a rst branch thereof a major portion of settled catalyst slurry from said settler to said lower mixing zone of said contacting vessel and by a second branch thereof to remove from said settler a minor portion of said settled catalyst,

(5) an outlet from the top of said exterior settler and,

(6) an inlet to said mixing zone of said contacting vessel for introducing catalyst slurry and charge stream;

said system further comprising a conduit for introducing a charge stream to the inlet of the mixing zone of the contacting vessel in the rst stage of said series; a conduit for introducing fresh catalyst to the inlet of the mixing zone of the contacting vessel in the last stage of said series; a purge conduit for removing said minor portion of catalyst from said second branch of said branched conduit of the rst stage; a conduit for transfer of said minor portion of catalyst from said second branch in each stage subsequent to said first stage to the inlet of the mixing zone of the contacting vessel in each preceding stage of said series; a conduit for transfer of charge stream from the outlet at the top of the exterior settler in each stage prior to the last stage into the inlet of the mixing zone of the contacting vessel in each subsequent stage; and a conduit for removal of a product stream from the outlet at the top of the exterior settler in the last stage.

References Cited in the tile of this patent UNITED STATES PATENTS 2,373,674 Crawford et al. Apr. 17, 1945 2,413,691 Crawford et al Jan. 7, 194'7 2,425,074 Waugh Aug. 5, 1947 2,433,079 Whiteley et al Dec. 23, 1947 2,459,636 Fenney Jan. 18, 1949 2,686,110 Carver Aug. 10, 1954 2,767,847 Russell et al. Oct. 23, 1956 2,905,736 Belden Sept. 22, 1959 2,937.915 Bleich et al. May 17, 1960

Claims (1)

1. THE METHOD OF CONVERTING NAPHTHA TO AN AROMATIC PRODUCT AND A SUPER PREMIUM MOTOR FUEL HAVING A CLEAR OCTANCE NUMBER SUBSTANTIALLY HIGHER THAN 100 WHICH METHOD COMPRISES SPLITTING NAPHTHA INTO A LIGHT FRACTION CONTAINING SOME C7 AND LLIGHTER HYDROCARBONS AND A HEAVY FRACTION CONTAINING C7 AND HEAVIER HYDROCARBONS, HYDROFORMING THE HEAVY FRACTION UNDER A SEVERITY TO OBTAIN REFORMATE OF WHICH THE HIGHER BOILING PORTION HAS AN OCTANE NUMBER OF ABOUT 102-104, SPLITTING THE REFORMATE INTO A LIGHT FRACTION CONTAINING C7 AND LIGHTER HYDROCARBONS AND A SEPARATE HEAVY REFORMATE FRACTION, DEPENTANIZING THE COMBINED LIGHT FRACTIONS, REMOVING SUBSTANTIALLY ALL AROMATICS FROM THE DEPENTANIZED LIGHT FRACTIONS TO OBTAIN A DEAROMATIZED STREAM, DEHEXANIZING THE DEAROMATIZED STREAM TO OBTAIN A HEXANE STREAM AND A HEPTANE STREAM, RECYCLING AT LEAST A PART OF THE HEPTANE STREAM TO THE HYDROFORMING STEP, ISOMERIZING THE HEXANE STREAM WHICH IS SUBSTANTIALLY FREE FROM AROMATICS AND WHICH CONTAINS 2-20% NAPHTHENES BY CONTACTING WITH ALUMINUM CHLORIDE-ON-ADSORBENT CATALYST CONTAINING ABOUT 15-30% ALUMINUM CHLORIDE UNDER A PRESSURE SUFFICIENT TO MAINTAIN LIQUID PHASE CONDITIONS AT A TEMPERATURE IN THE RANGE OF ABOUT 60-180*F. FOR A TIME TO CONVERT A LARGE PROPORTION OF THE NORMAL LEXANE AND METHYLPENTANES TO DIMETHYLBUTANES, SSPLITTING THE ISOMERIZATION PRODUCT BY FRACTIONAL DISTILLATION TO OBTAIN A STREAM WHICH IS CHIEFLY NEOHEXANE AND A STREAM CONSISTING ESSENTIALLY OF METHYLPENTANES AND NORMAL HEXANE, RECYCLING THE LATTER STREAM TO THE ISOMERIZATION STEP AND COMBINING THE NEOHEXANE STREAM WITH THE HEAVY REFORMATE FRACTION TO PRODUCE SUPER PREMIUM GASOLINE HAVING A CLEAR OCTANE NUMBER SUBSTATNIALLY HIGHER THAN 100.

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