US2938936A - Isomerization of saturated hydrocarbons - Google Patents

Isomerization of saturated hydrocarbons Download PDF

Info

Publication number
US2938936A
US2938936A US658877A US65887757A US2938936A US 2938936 A US2938936 A US 2938936A US 658877 A US658877 A US 658877A US 65887757 A US65887757 A US 65887757A US 2938936 A US2938936 A US 2938936A
Authority
US
United States
Prior art keywords
per day
fraction
barrels per
hydrocarbon
isomerization
Prior art date
Legal status (The legal status is an assumption and is not a legal conclusion. Google has not performed a legal analysis and makes no representation as to the accuracy of the status listed.)
Expired - Lifetime
Application number
US658877A
Inventor
Donald H Belden
Current Assignee (The listed assignees may be inaccurate. Google has not performed a legal analysis and makes no representation or warranty as to the accuracy of the list.)
Universal Oil Products Co
Original Assignee
Universal Oil Products Co
Priority date (The priority date is an assumption and is not a legal conclusion. Google has not performed a legal analysis and makes no representation as to the accuracy of the date listed.)
Filing date
Publication date
Application filed by Universal Oil Products Co filed Critical Universal Oil Products Co
Priority to US658877A priority Critical patent/US2938936A/en
Application granted granted Critical
Publication of US2938936A publication Critical patent/US2938936A/en
Anticipated expiration legal-status Critical
Expired - Lifetime legal-status Critical Current

Links

Images

Classifications

    • 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
    • C10G35/00—Reforming naphtha
    • C10G35/04—Catalytic reforming
    • C10G35/06—Catalytic reforming characterised by the catalyst used
    • C10G35/085—Catalytic reforming characterised by the catalyst used containing platinum group metals or compounds thereof

Definitions

  • this invention relates ⁇ to a particular series of process steps, the use of which result in maximum catalyst life and optimum product quality.
  • This invention is particularly directed to an improved process for, the production of dimethylbutanes by the isomerization of relatively straight chain hexane hydrocarbons.
  • n-Butane and npentane have been isomerized to isobutane and isopentane ⁇ respectively by various prior art processes utilizing either liquid or vapor phase.
  • cracking occurs along with isomerization, and that this cracking increases with increasing molecular weight of the hydrocarbon reactant.
  • hexane fractions have never been successfully isomerized on a commercial scale.
  • a process for such isomerization is particularly attractive when it is realized thata hexane fraction can be converted by proper isomerization and fractionation into a pool of motor fuel having an F-l +3 cc. octane number of over 100. It is there fore an object of this invention to provide a process which will yield these desired high octane hexane isomers.
  • Hexane hydrocarbonsboil from about 40 to about 90 C. (m4-185 F.) and analyses of typical hexane fractions show that they contain varying quantities ofcyclopentane, 2,2-dimethylbutan e, 2,3- dimethylbutane, ⁇ Z-methylpentane, 3-methylpentane, nhexane, methylcyclopentane, dimethylpentanes, cyclohexane, and benzene.
  • the boiling points ⁇ and octaneunurnbers vof these six carbon atom hydrocarbons are given in the following table:
  • the high octane number components are not passed to the isomerization ⁇ reactor but are fractionated from the fresh feed.
  • cornponents include the 2,2-dimethylbutane, ⁇ 2,3-dimethylbutane, methylcyclopentane, benzene, and cyclohexane.
  • the reaction zone feed will comprise 2-methylpentane, 3-methylpentane, and n-hexane.
  • reaction zone feed will be converted therein to an equilibrium mixture of hexane isomers including the dimethylbutane cornpounds.
  • a total gasoline pool havingan F-l-l-3 cc. octane number of greater than 100 is produced.
  • the present invention relateshto a process for the isomerization of an isomerizable saturated hydrocarbon fraction, characterized hy'an average molecular weight greater than about 80, to produce a composite hydrocarbon fraction having an F-l-l-B cc. octane number of at least 100, said isomer-ization being carried out in the presence of an isomerizationcatalyst,
  • octane number of at least as one product from the process passingf said substantially naphthene free overhead hydrocarbon fraction from said firstY fractionation zone to a second fracjdi''rrailonsarefisoinerizecl to higher octane number 'hydroc'arbons, recycling tlieisonieriz'ation'zone eliuent to 'thefrstnientioned fractionation zone as aforesaid to'recoverthe *hig'h octane number hydrocarbons therefrom rs vproduct'andto separate low"octanejnumber hydroey'arbgns'as" aforesaid for reuse fandv internal recyclel in jsaid procesa "and v'compositiilg 'the f aforesaid i product streams.
  • Afspecicernbodiment ofthe invention relates to a ce'ss uiforth'e"isor'ne'rrization lof an isomerizable hexane "friction to 'proluce ⁇ a Composite hexane fraction having jf'niFi-'l-l-:Bcc octane number of'at "least 100, said isoime'riza'tio'n lbeing carried-out in the presence of hydroge'n and nisomeriz'ation catalyst comprising platinum,
  • halogen which comprises pass- "ingitoa 'first vfractionation zone said first mentioned Yhex- ⁇ ⁇ ane'"hydro'carbon fraction in combination with liquid phaseisomerization zone eluent producedl as hereinafter escribed, ylfra'ctionating -said hydrocarbons to produce lsiibstantiallyr naph'tliene ⁇ free yoverhead hydrocarbon frac- 'i-tin and tofproducea substantially pure cyclics-hydrocarbon bottoms fraction containing methylcyclopentane and cyclohexane, removing said vcyclics hydrocarbon bot- ⁇ toms fraction containing methylcyclopentane and cyclo- 'hexane as one product from the process, passing said ,substantially naphthene free overhead hexane fraction from'lsaid rst fractionation zone to a lsecond fraction-
  • the combination process ofgthis invention has diverse advantages, all interrelated.
  • the process prevents a buildup of heavier products in the combined feed to the reaction zone since these heavier products are eliminated by removal in the bottoms from the first fractionation zone-"along with methylcyclopentane and cyclohexane.
  • k"11n-a preferred embodiment the isomerization process of "thefpresent invention-is carried out'in the presence of fhydogen. "In'the Vpresence ofhydrogen, cracking of yfnaphthenic hydrocarbons occurs readily due to the acid )function ofthe preferred catalysts.
  • the Vprocess v prevents *iso-called reverse Zisomerization of 'high octane nurnber "dimethylbutanes to other"hexane Visomers by separation of these high Aoctane number cornponentsiprior to'isomeriz'ation. Atany particular isomeiization'reaction temperature,the formation of hexane 'isomers 'is limited by theequilibrium distribution-of these isomers.
  • Vthis invention relates to ⁇ a process V:forl the isomerization of "an isomerizable saturated hydrocarbon 'fraction characterized by'ran average moleculanweight-greater than about'SOQ ⁇ Hydrocarbons within'the scope 'of thel-aboveflimitation and'utilizable in ⁇ vthe process of this invention include methylcyclopentane, ⁇ cyclohexane, n-hex'ane, Y 2m ⁇ e ⁇ thylpentane, 3-methyl ticulafrlyy applicable to'V the isomerization off-*hexane -fracvarioussonrcesincluding fractionationlfrom straight run 'gas'olinaA straight rim1 naphtha, natural gasoline, catalyti- 76 eally'refr'f'ied hpltha, andfcatalytic'ally reformedl'gasol. Analyses of typical hexane fractions
  • Various isomerization catalysts are utilizable Within the generally broad scope of the process of the present invention. These catalysts include a support, an acidacting function, and a hydrogenation component.
  • the support may be selected from diverse refractory oxides including silicia, alumina, silica-alumina, silica-aluminamagnesia, silica-alumina-zirconia, silica-zirconia, etc. Depending up'on the method of preparation and upon the treatment of the support thereafter, these various supports will have surface areas ranging from about 25 to about 500 square meters per gram. In some of these supports the acid-acting function is inherently present, as when silica-alumina is used as the support.
  • the amount of effectiveness of this acid-acting function is then controlled by the quantity of silica which is combined with the ⁇ alumina, and by the treatment of the silica-alumina, particularly by calcination, prior to or after compositing the hydrogenation component therewith.
  • alumina is preferred, and particularly gammaalumina having a surface area of from about 150 to about 450 square meters p'er gram.
  • the acid-acting function can be added to the catalyst by the incorporationtherein of what is known in the art as'combined halogen.
  • the amount of combined halogen can be varied from about 0.01 to about 8% by weight based on the alumina.
  • both uorine and chlorine can be used satisfactorily.
  • alumina type catalyst to be utilized -at reaction temperatures of from 750 to about 850 F., about 0.3% by weight of uorine and about 0.43% by Weight of chlorine may be incorporated therein.
  • the combined halogen which will be utilized along with the alumina support is uorine, and this fluorine will be utilized in an amount of from about 2.5% to about 4.5% by weight.
  • the composite will then have the desired hydrogenation component combined therewith.
  • This hydrogenation component will normally be selected from groups VI(B) and VIII of the periodic table or mixtures thereof.
  • Such hydrogenation components include chromium, molybdenum, tungsten, iron, cobalt, nickel, and the socalled platinum group metals including platinum, palladium, ruthenium, rhodium, osmium, and iridium. Of the various hydrogenation components which may be utilized, those of the platinum group metals are preferred, and of these platinum group metals, platinum itself is particularly preferred.
  • the hydrogenation component ofthe catalysts of the present invention will normally be utilized inan amount of from about 0.01% to about ⁇ thequantity utilized will range from about 0.01% to about 2% by weight.
  • a particularly preferred catalyst comprising platinum, combined halogen, and alumina will contain 0.4% platinum, 4.0% iluorine, and alumina. Because of equilibrium considerations and because it is often desirable and/or advisable to carry out the isomerization reaction at the lowest possible temperature, for example, from about 300 to about 500 F., catalysts may also be prepared by impregnating ⁇ composites such as described hereinabove with a metal halide of the Friedel-Crafts type. For example, an excellent low temperature isomerization catalyst can be prepared by impregnating from about 5 to about 20% aluminum chloride onto a composite of platinum, alumina, and combined halogen.
  • the isomerization process of the present invention can be carried out ⁇ at varying conditions of temperature, pressure, space velocity, and combined feed ratio.
  • the temperature utilized will particular catalyst selected. Thus the temperature may range over a wide range of from about 300 to about 800 F.
  • the pressure will be selected so as to insure vapor phase operation and depending upon theparticular temperature utilized will range from about to about 1000 pounds per square inch or more.
  • Liquid hourly space velocity will range from about 0.1 to about 10 or more, the only limitation being that equilibrium mixtures of isomerized hydrocarbonsl or a close approach thereto shall be obtained in the reaction zone etluent.
  • the combined feed ratio which is defined as the total amount of fresh feed entering the reactor and recycle divided by the quantity of fresh feed will range from 1 up to about 5 or more.
  • Fractionation zone 4 is heated by reboiling agportion of the 'higherboiling hydrocarbons which are withdrawn through lines 14 and 15, and are passed through heat vexchangerjlr. The lheated hydrocarbons are thenreturned to a'lowfer portion of fractionation zone 4 through line 17. 'Ihe higher boiling substantially 4pure cyclics hydrocarbon bottoms fraction having an F-l-Jf-K cc. octane number of atleast -100 fis withdrawn from the bottom of fractionation zone 4 through Vlines Vll4 and 18, and is pumped by with the other high octane number fraction hereinafter described.V l
  • fractionationzone 21 The purpose of fractionationzone 21 is to fractionate the feedvthereto to produce a hydrocarbon Vfraction characterized by a major proportion of the hydrocarbons with at least two methyl substituents per molecule and and F-l-l-3 ce. octane number of at least 100 and to produce ⁇ reactionzone feed.
  • the 'overhead product from this fractionation zone 21 thus comprises dimethylalkanes naturally occurring in the feed and dirnethylalkanes from thereactionrzone effluent which have been formed as a result ofthe isomer'iza'tion-of riparans and vrn lonomethylalkanes as setv forth hereinafter.
  • the overheadkproduct which is separated Vfrom the process at this point comprises 2,2- and 2,3Ldimethylbutanes naturallyroccurring in the feed and Y2,2- and 2,3-dimethylbutanes from the reaction zone effluent which have been formed as a result of the isomerization'of -n-hexaneV and 2- and 3-methylpentanes as set forth hereinafter.
  • any water which is dissolvedin-the feed or produced during the processing separates overhead atthispoint due to distillation drying.
  • the ⁇ net hydrocarbon stream from fractionator 21 is pumped bypump 37 through line 38 'and Aline 39.V From "line 39 thisrhydrocarbon stream isrheat exchanged with the hot reaction Azone effluent in heat exchanger 40 ⁇ ,and then passes through line 41 to' heater 42. 'In heater .42,
  • Vthe stream is heated to the ⁇ desired reaction temperature, and then passes through lineA V43 ⁇ to reaction zone'44l It is in this reaction zone 44 thatthe-less highly lzyrarrched Chain and Straight chain' hy droiciarbolns,Y are isomerized. to an equilibrium mixture ofi-saturated hydrocarbonfiso.- mers.
  • the hydrocarbons are processed inthe presence of hydrogen whichfis introducedvia lines 52 and 5 4iintoi1ine39.
  • ThisV hydrogen also -hydrogenates any aromatic'hydrocarbonsin'reaction zone 44 and ⁇ the thus produced naphthenie ⁇ hydrocarbons arethen isomerized to an equilibriumrmixture of the sameduring isomerization of the remaining saturated hydrocarbons.
  • reaction zone 42 will depend upon the particular isomerization catalyst utilized therein. It was pointed out above that a preferred catalyst in the process of this invention 'is one comprising va platinum group metal, particularly platinum, combined halogen, and alumina. With such a catalyst, the pressure utilized in the reaction zone will range from about to about 1000 pounds per square inch, the temperature will range from about 300 to about 800 FL, and the hourly liquid space velocity will range from about 0.1 to about l0. The hydrogen to hydrocarbon ratio in the reaction zone will range from about 0.2.5 to about 10 mois of hydrogen per mol of hydrocarbons.
  • the reaction temperature will be lower than the higher part of the above set forth range, for example, from about 500 to about 750 F.
  • the catalyst comprises platinum, alumina, and lcombined halogen onto which has been impregnated a metal halide of the Friedel- Crafts type, such as aluminum chloride, the temperature will still be lower, for example, from about 300 ⁇ to aboutr50 F.
  • reaction zone eiuent passes from reactor 44 through ⁇ line'f45 in indirect heat exchange with ⁇ the reaction zone feed through heat exchanger 40 and then passesthrough line 46, is cooled in condenser 47, and passesthrough line ⁇ 48 to high pressure separator 49.
  • the Yhigh pressure separator 49 is utilized for separating hydrogen from the reaction rzone eflluent, which hydrogen is separated and passed through line Y50 to compressor 51 where its pressure is increased to the desired number of pounds per square inch, and then the hydrogen is discharged through lines 52 and 54 described hereinabove.
  • a hydrogen halide' is utilized along with the hereinabove described catalystV as a catalyst promoter, it will pass through line 50 to compressor 51 as hereinabove described. Hydrogen and/ or hydrogen-halide for startup and makeup are added through line 53.
  • the cooled reaction zone effluent after high pressure separation is discharged from separator '49 through line 55 to debutanizer 56.
  • This debutanizer is a conventional fractionation zone by means of which a light hydrocarbon gas or low boiling cracked products and dissolved H2 is removed from the process. This removal is accomplished through Vline :57 whichwilll includere;
  • EXAMPLE I One specific example of .the operation of the process with a platinum-aluminacombined halogen catalyst and as the process is carried outsimilar to that set forth hereinabove with reference to thedrawing'is described herewith nowin connectionwith the drawing.
  • the catalyst utilized comprises alumina containing 0.4% platinum and about 0.5 %v combined halogen.
  • the combined halogen comprises about 0.3% uorine and about ⁇ 0.2% chlorine.
  • This example ⁇ illustrates the isomerization of a Gulf Coast hexane fraction, boiling point 40-74 C.
  • composition of the Gulf Coast hexane fraction is as follows: cyclopentane, 3%, ⁇ 2,2-dimethylbutane, 3%; 2,3- dimethylbutane, 5%; Z-methylpentane, 25%; S-methylpentane, 20%; n-hex-ane, 30%; methylcyclopentane, 3%; dimethylpentanes, 1%; cyclohexane, 7%; andiA benzene, 3%.
  • this hexane fraction having an average molecular weight greater than 80 inthe quantity of 1000 barrels per day is passed as a liquid under pressure through line l1 and through line 3 to an upper portion of fractionator 4.
  • This fresh feed stream has combined therewith in line 1, ⁇ 3830 barrels per day of recycle reaction zone elfluent from line 2 as hereinafter described.
  • This recycle hexane stream consisting of the liquid reaction zone eiuent contains 452 barrels per day of 2,2-dimethylbutane, 298 barrels per day of 2,3-dimethylbutane, 1310 barrels per day of Z-methylpentane, 8,62 barrels per'day of B-methylpentane, and 908l barrels per lday of n-hexane.
  • the total combined ⁇ feed to fractionator 4 amounts to 4830 ⁇ barrels per day comprising 30 barrels per day ofcyclopentane, 482 barrels per day of 2,2-dimethylbutane, 348 barrels per day of 2,3-dimethylbutane, 1560 ⁇ barrels per day of 2-methylpentane, 1062 barrels per day of S-methylpentane, 1208 barrels per day of nhexane, 30 barrels per day of methylcyclopentane, barrels per day of dimethylpentane, 70 barrels per day of cyclohexane, and 30 barrels per day of benzene.
  • the combined feed to fractionator 4 is fractionated therein under a 4:1 molal ⁇ reflux to feed ratio and a substantially naphthene free stream -is separated overhead therefrom in an amount of 4690 barrels per day.
  • This 4690 barrels per day contains 30 barrels per day of cyclopentane, 482 barrels per day of 2,2-dimethylbutane, 348 barrels per day of 2,3- dimethylbutane, 1560 barrels per day of 2-methylpentane, 1062 barrels per day of 3methylpentane, and 1208 barrels per day of nhexane.
  • This tfractionator ⁇ 4 overhead is withdrawn therefrom through line 5, is condensed in cooler 6, and is passed through line 7 to receiver 8. From receiver 8 the liquid is withdrawn through line 9 by pump10 and supplies feed to fractionator 21, described hereinafter, via line 13. i
  • the feed stream to ⁇ fractionation ⁇ zone 21 in the quantity of 4690 barrels per day and of the composition describedv hereinabove is fractionated therein molal redux to feedjratio ⁇ and ⁇ the overhead passed therepump19 through lines 20 and 31 ⁇ per day of n-hexane.
  • Fractionatio zone 21 ⁇ separates a substantially monomethylpentane and n-hexane free dimethylbutane stream overhead therei from in an amount of 860 barrels per day. This ⁇ 860 barrels per day contains 30 barrels per day of cyclopen tane, 482 barrels per day of 2,2-dimethylbutane, ⁇ and 348 barrels per day of ⁇ 2,3-dimethylbutane- This stream has an F-ll3 cc. octane number of 112.0.
  • This fractionator overhead ⁇ is withdrawn from receiver 25 through line 26 by pump 27 andthe net product is pumped through lines 30 and 31 to storage.
  • the substantially dimethylbutane free bottoms stream from fractonator 21 in the quantity of 3830 barrels per day is the reactor feed.
  • the composition of this stream is as follows: 1560 barrels per day of 2methylpentane, ⁇ 1062 barrels per day of-3-methylpentane, 1208 barrels
  • This bottoms stream is pumped by pump 37 through lines 38 and 39, where it is joined with hydrogen from line 54, through heat exchange zone 40, and line 41 to heater 42.
  • the quantity of hydrogen which is continuously supplied is sufficient to maintain a hydrogen to hydrocarbon mol ratio of 2:1 in the reaction zone. This is accomplished by recycling hydrogen from the high pressure separator back to the reaction zone.
  • a small quantity of substantiallyk pure makeup hydrogen is added through line S3 to make up for the hydrogen con sumed in the reaction and for that which is dissolved in the separator liquid efuent.
  • the combined feed to thereaction zone is heated to a temperature of 800 F.
  • the reaction is carried out at a pressure of about 500 p.s.i.g. in vapor phase at a liquid hourly space velocity of about 2.0.
  • reaction zone 44 isomerization of the reaction zone
  • feed is accomplished in reaction zone 44 in the presence of the hereinabove described catalyst and results in the production of 3830 barrels per day of reaction zone effluent containing 452 barrels per day of 2,2- dimethylbutane, 298 barrels per day ⁇ of 2,3-dimethylbutane, 1310 barrels per day of 2-methylpentane, 862 barrels per day of 3-methylpentane, and 908 barrels per day of n-hexane.
  • reaction zone effluent after heat exchange with incoming reaction zone feed is passed via lines 45 and 46 to condenser 47, and the liquid product and hydrogen gas are then passed through line 48 to high pressure separator 49.
  • the hydrogen gas and minor amounts of low boiling hydrocarbons formed by cracking in the reaction zone are separated in high pressure separator 49 and passed through line 50 to compressor 51.
  • High pressure separator liquid which discharges through line 55 is passed to debutanizer 56 to further separate this small amount of cracked products produced during isomerization and to prepare the liquid recycle stream as hereinabove set ⁇ forth.
  • the liquid recycle stream passes from debutanizer 56 through lines 58 and 62 to pump r63 wherein it is recycled to fractionation zone 4 via line 2.
  • the catalyst utilized comprises alumina ,containing 0.4% ⁇ platinum and about 4.0% combined iluorine.
  • This example illustrates the isomerization of the Gulf Coast hexane fraction of the first specific example.
  • This hexane fraction in they quantity of 1000 barrels per day isfpas'sed as a Vliquid under pressure through line 1 and through line 3 to Yan upper portion of fractionator 4.
  • This' l000barrels per day is made up of 30 barrels per day of cyclopentane, 30 barrels per day of 2,2-dimethylbutane, 5,0 barrels per day of 2,3-dimethylbutane, 250 barrels per day of 2 rnethylpentane, 200 barrels per day of 3-meth ⁇ ylpentane, 300 barrels per day of n-hexane, 30 barrels per day of methylcyclopentane, barrels per day of dimethylpentanes, 70 barrels per day yof cyclohexane, and 30 barrels per day of benzene.
  • This fresh feed stream has combined therewith in line 1, 3290 barrels per day of recycle reaction zone efuent from Iline 2 as hereinafter described.
  • This recycle hexane stream consisting of liquid reaction zone eiuent contains 494 barrclsper day of 2,2-dimethylbutane, 257 barrels per day of 2,3- dimethylbutane, 1085 barrels per day of 2-methylpentane, 727 barrels per, day of 3-methylpentane, and 727'barrels per day of n-hexane.
  • the total combined feed to fractionatorl amounts to 4290 barrels per day consisting of 30 Abarrels per vday of cyclopentane, 524 barrels per day of 2,2-dimethylbutane, 307 barrels per day of 2,3- dimethylbutane, 1335 barrels per day of 2-methylpentane, 927 barrels per day of 3methylpentane, 1027 barrels per day of n-hexane, 30 barrels per day of methylcyclo-V pentane, 1,0 barrels per day of dimethylpentanes, 70 barrels per day of Vcyclohexane, and 30 barrels per day of benzene.
  • the combined feed to fractionator 4 is fractionated thereinl under va 4:1 molal reux to feed ratio and a substantially naphthene Vfree stream is separated overhead therefrom in an amount of 4150 barrels per day.
  • This -4150 barrels per day contains 30 barrels per day vof cyclopentane, 524 barrels per day of 2,2-dimethylbutane, 307 barrels per day of 2,3-dimethylbutane, 1335 barrels per day of Z-methylpentane, 927 barrels per day of S-methylpentane, and 1027 barrels per day of nhexane.
  • This fractionator 4 overhead is withdrawn therefrom through lineY 5, is condensed in cooler 6, and is passed throuh line 7 to receiver 8. From receiver 8 the liquid is withdrawn through line 9 by pump 10 and supplies feed to fractionator 21, described hereinafter, via lne'13 ⁇ .
  • 'Fractionator 4 bottoms are withdrawn through line 14 andthrough line 18 and have a composition as follows: 30 barrels per day of methylcyclopentane, 10 bar-A rels per day of dimethylpentanes, 70 barrels per day of cyclohexane, andv 30 barrels per day of benzene.
  • the F-1+3 ,octane number of this stream is 103.7.
  • Fractionation zone 21 separates a substantially monomethylpentane and n-hexane free dimethylbntane stream overhead therefrom in an amount of 86,1V barrels per day.
  • This 861 barrels per day contains30 ⁇ ibarrels penday of cyclopentane, 524 barrels per day of 2,'2-'dimethylbutane, and 307 barrels per day' of 2,3-dimethylbutane.
  • This Vstream has an F-1+3 cc. octane number of 109.8.
  • This fractionator overhead is withdrawn from receiver 25 through line 26 by pump 27 and the net product'is pumped through lines 30 Aand 31 to storage.
  • this stream with its F.l+3 cc..octane.number of 109.8 is blended with the aforementioned vbottomslstream for fractionator .4, said 'latter ystream havingV an F-l-PS ce. octane '1.1wr1br"' f v '103.7, ⁇ fthe combined pool thus having an 'F-l-t-B ce. octanenumberof greater than 100'. f'
  • Thefsubstantia-lly Y 1 Y from ⁇ fractonator '21l in the quantity of 3289 barrels-'per day is thelreactor feed.
  • the composition of this s'tream is as ⁇ followsfl335 barrels Vper day'ofV 2-methylpentane, V927 barrels-per day of 3methylpenftane,"and I027"ba'rrels per day of n-hexane.Y 'This bottoms stream is pumped by pump 37 through lines 3 8 and 39, where it; is? joined with hydrogen from line 54, through heat exchange zone 40, and line 41 to heater 42.
  • ⁇ T he quantity of hydrogen which is continuously suppliedis-suicient tomaintain-a hydrogen toA hydrocarbon mol ratio of 2:1 inthe reacl tion zone. This is' accomplished by recycling'hydrogen from the high pressure separator back to the yreaction zone. -A small amount of substantially 'pure'makeup hydrogen is added through line '53 tei-make upj-for-the hy'- drogen consumed in the reaction and for that whichlis dissolved in theseparator liquid ethuent. By 'means of heat exchanger 40 and heater 42, the combined feed to the reaction zone is heated to a temperature ofV 600 F.
  • TheY reaction is Vcarried out at av pressure' of about 400 p.s.i.g. in vaporphase at an hourly liquid space'velocity of 2.0.
  • av pressure' of about 400 p.s.i.g. in vaporphase at an hourly liquid space'velocity of 2.0.
  • isomerization of the reaction zone feed is accomplished in reaction zone 44 in Vthe presence of the hereinabove described catalyst and results in the production of 3289 barrels per day of reaction zone effluent cont, 'ng 494 barrels per vday of 2,2-dimethylbutane, 257 barrels per ⁇ day of '2,3-dimet ⁇ hyl butane, 1085 barrelsper day of Y2-rnethylpen'tane, V727 barrels per day of B-methylpentane, 'and 727. barrels per day of n-hexane.
  • the ⁇ reactionf-'zone -efuentafterV heat exchange with the incoming reaction zone feed is passed via lines'45 and 46 to condenser 47, and the liquid ⁇ product and;hydrog en gas are then'passedthrough line 48 to high pressure sepa,- rator 49;- k'The Vhydrogenl gas andjminor amounts of iovv boiling hydrocarbons formed by cracking in the reaction ozone areV separated in high pressure separatorV 49 and passed through vliue50 to compressor (5,1.
  • High pressure separator liquid v which discharges through line 55 is passed to .debutanizer Y56 to furtherseparatethe small amount ofunreacted products produced duringisomerization and to prepare the liquid recycle stream jas hereinabove set forth.
  • the liquidrecycle stream passes from debutanizer ,'56 through lines 58 and 62 to pump 6,3 wherein it is recycled to fractionation zone 4 via line-H2.
  • thisheirane fraction A having anV average vmolecular ,weight greater than ,i,n the quantity of 100,0,.barrels per day lis passed as a liquid under pressure through line 1 Aand through line 3 to an upper portion of fractionation ,zone 4Q
  • This 1,0 00 barrels per day of .Gulf Coast hexane'fnactioncontams 30 barrels per day of cyclopentane, 30 barrels per dimethylbutanefree bottoms streaml d by pump 27 and day of 2,2dimethylbutane, SObarrels per day of 2,3-dimethylbutane, 250 barrels per ⁇ day of Q-methylpentane, ⁇ 200 barrels per day of S-methylpentane, 300 barrels per ⁇ day of n-hexane, 30 barrelsper day of' methylcyclopent
  • This feed ⁇ stream has combined therewith'in
  • This recycle ⁇ hexane stream consistingy of the liquid reaction zone effluent contains 495 barrels per day of 2,2-dimethylbutane, 254 ⁇ barrels per day of 2,3-dimethylbutane, 818 barrels per day of 2-methylpentane, 475 barrels per day of 3methylpentane, and 358 barrels per day of n-hexane.
  • the total combined feed to fractionator 4 amounts to 3400 barrels Vper.
  • day comprising 30 barrels per day of cyclopentane, 525 barrels per day of 2,32-dimethylbutane, 304 barrels per day of 2,3-dimethylbutane, 1068 barrels per day of Z-methylpentane, 675 barrels per day of 3methylpentane, 658 barrels per day ofA n-hexane, 30 barrels per day of methylcyclopentane, 10 barrels per day of dimethylpentanes, 70 barrels per day of cyclohexane, and 30 barrels per day of benzene.
  • Thecombinedfeed to fractionator 4 is fractionated therein under a 4:1 molal reflux to feed ratio and a substantially naphthene free stream is separated overhead therefrom in an amount of 3260 barrels per day.
  • This 3260 bar,- rels per day contains 30 barrels per day of cyclopentane, 5,25 barrels per day of 2,2-dimethylbutane, 304 barrels per day of 2,3-dimethylbutane, 1068 ⁇ barrels per day of 2methylpentane, 675 barrels per day of 3-methylpentane, and 658 barrels per day of n-hexane.
  • This fractionator 4 ⁇ overhead is Withdrawn therefrom through line 5, is condensed in cooler 6, ,and is .passed through line 7 to receiver 8. From ⁇ receiver 8 the liquid is withdrawn through line 9 by pump 10 and supplies feed to fractionator 21, described hereinafter, via line 13.
  • Fractionator4 bottoms are withdrawn from line 14 through line 18 and have a composition as follows: 30 barrels per day of methylcyclopentane, 10 barrels per dayof dimethylpentanes, 70 barrels perV day of cyclo-V hexane, and 30 barrels per day of benzene.
  • the F-1-l-3 cc. octane number of this stream is 103.7.
  • Fractionation zone 21 separates a substantially monomethylpentane and n-hexane free dimethylbutane stream overhead therefrom in an Iamount of 859 barrels per day. This 859 barrels per day contains 30 barrels per dayof cyclopentane, 525 barrels vper day of 2,2-dimethylbutane, and 304 .barrels per day ⁇ ofQ-dimethylbutane.
  • This stream has an F-1-i-3 cc.' octane number of 109.8.
  • This fractionator overhead is withdrawn from receiver 25 through line 26 the net product is pumped through lines 30 and 31 to storage.
  • this stream with its F1+3 cc. octane number of 109.8 is blended with the aforementioned bottoms stream from fractionator 4, said latter stream having lan F-l-l-3 cc. octane number of 103.7, the combined pool thus having an F-l-i-S cc. octane number greater than 100.
  • the substantially dimethylbutane free bottoms stream from fractionator 21 in the quantity of 2401 barrels per day is the reactor feed.
  • the composition of this stream is as follows: 1068 barrels per day of 2-methylpentane, 675 barrels per day of S-methylpentane, and 658 barrels per day of n-hexane.
  • This bottoms stream is pumped by pump 37 through lines 38 and 39, where it is joined With hydrogen from line 54, ⁇ through heat exchange zone 1 l 40 and line 41- to heater 42. which is continuously supplied is sufficient to maintain a hydrogen to hydrocarbon mol ratio of 1:1 in the reaction zone. This is accomplished by recycling hydrogen from the high pressure separator to the reaction zone.
  • a small amount of substantially pure makeup hydrogen is added through line 53 to make up for the hydrogen consumed in the reaction and for that which is dissolved in the separator liquid ⁇ eiuent.
  • the combined feed to the reaction zone is heated to a temperature of 350 F.
  • the reaction is carried out at a pressure of about 300 p.s.i.g. in vapor phase at an hourly liquid space velocity of 0.5.
  • isomerization of the reaction zone feed is accomplishedA in reaction zone 44, in the presence of the hereinabove described catalyst and ⁇ results in the productionof 2400 barrels per day of reaction zone effluent containing 495 barrels per day of 2,2-dimethylbutane, 254 barrels per day of ⁇ 2,3-dimethylbutane, 818 barrels per day of Z-methylpentane, 475 barrels per day of ⁇ 3methylpentane, and 358 barrels per day of nhexane.
  • reaction zone efuent after heat exchange with incoming reaction zone feed is passed via lines 4S and 46 to condenser ⁇ 47, ⁇ and the liquid product and hydrof gen gas are then passed through ⁇ line 48 to high ⁇ pressure separator 49.
  • the hydrogen gas and minor amounts of low boiling hydrocarbons formed by cracking in the reaction zone areseparated in high pressure separator 49 and passed through line 50 to compressor 51.
  • High pressure separator liquid ⁇ which dischargestthrough line 55 is passed to debutanizer 56 to further separate the small amount of cracked products producedV during isomeriza- The quantity of hydrogen tion and to preparethe liquid recycle stream as herein- ⁇ above set forth.
  • the liquid recycle stream passes from debutanizer 56 through lines 58 and 62 to pump 63 wherein it is recycled to fractionation zone 4 via line 2.
  • the liquid volume yields based on the feed are in the order of 98 volume percent or higher.
  • the F-1- ⁇ 3 cc. octane number of the dimethylbutane fraction is 109.8 and the F-l-l-B cc. octane number of the cyclics fraction is 103.7,
  • octane number of at least 100 saidisomerization being carried out inthe which comprises passing to a iirst fractionation zone said irst mentioned hydrocarbon fraction in combination with liquid ⁇ phase isomerization zone effluent produced as hereinafter described, fractionating said hydrocarbons tol produce a substantially naphthene-free overhead hydroi carbon fraction and to produce a substantially pure cyclics hydrocarbon bottoms fraction, tally pure cyclics hydrocarbon bottoms fraction having an F-1+3 cc.
  • octane number of at least 100 as one product from the process passing said substantially naphthene-free overhead hydrocarbon fraction from said first fractionation zone to a second fractionation zone as freed therefor, fractionating said hydrocarbons to produce an overhead fraction characterized by a major proportion of hydrocarbons with at least two methyl substituents per molecule and an F-l-l-3 cc. octane number of at least 100 and to produce a bottoms hydrocarbon fraction, removing said overhead hydrocarbon fraction as a second product from the process, passing said last mentioned bottoms hydrocarbon fraction to an isomerization zone wherein low octane number hydrocarbons are presence of anv isomerization catalyst,

Description

May 31, 1960 D. H. BELDr-:N
ISOMERIZATION OF SATURATED HYDROCARBONS Filed May 13, 1957 Mm m@ IsoMERIzAnoN F SATURATED nYDRocARBoNs Donald H. Belden, North Rivmida'ul., assigner, by mesne assignments, to Universal Oil Products Company, Des Plaines, lll., a corporation of Delaware Filed May 13, 1957, ser. No. 658,871 s claims. (ci. 26o-683.68)
process high yields of high antiknock hydrocarbon fractions are produced with minimum loss to byproducts such as dry gas land `highboiling materials. Simultaneously with the maximum utilization of available relatively straight chain paraffnic hydrocarbons, this invention relates` to a particular series of process steps, the use of which result in maximum catalyst life and optimum product quality. This invention is particularly directed to an improved process for, the production of dimethylbutanes by the isomerization of relatively straight chain hexane hydrocarbons. These objectives are accomplished by the unique combination process of the present invention as will be set forth hereinafter.
Production of highly branched chain parainic hydrocarbons `having high antiknock properties and therefor suitable for use 4in automotive and aviation fuels is of considerable importance in the petroleum refining industry. Furthermore, the recent introduction of automobile engines of highcompression ratio has necessitated the utilization of high` antiknock fuels in these engines to obtain maximum horsepoweroutput therefrom. Thus the demandfor higher and higher octane number fuels has led tothe need for increased quantities of highly branched chain paraiiinic hydrocarbons for use as blending agents in gasoline. A convenient source of highly branched chain parainic hydrocarbons is the catalytic isomerization of less highly branched chain parallinic hydrocarbons. n-Butane and npentane have been isomerized to isobutane and isopentane `respectively by various prior art processes utilizing either liquid or vapor phase. However, it is, well known in the prior art that cracking occurs along with isomerization, and that this cracking increases with increasing molecular weight of the hydrocarbon reactant. Thus hexane fractions have never been successfully isomerized on a commercial scale. A process for such isomerization is particularly attractive when it is realized thata hexane fraction can be converted by proper isomerization and fractionation into a pool of motor fuel having an F-l +3 cc. octane number of over 100. It is there fore an object of this invention to provide a process which will yield these desired high octane hexane isomers.
Prior are processes for the isomerization of saturated hydrocarbons have taught the utilization of various catalytic agents to accelerate the desired molecular re-I arrangement at the conditions selected. Ordinarily, the catalytic agents utilized have comprised metal halides `such as aluminum chloride, aluminum bromide, etc.,
which were activated by addition of the respective hydrogen halide thereto. y These catalytic agents are very -wh1ch compnses passing to a first fractionation zone 2,938,936 Patented May 3l, 1960 ice ac tiveand effect high conversions per pass. However this high activity is accompanied by many disadvantages. One of the greatest disadvantages is the fact that these .catalytic materials not only accelerate isomerization reactions, but they also induce decomposition reactions. These `decomposition reactions are particularly detrimental to the overall economics of an isomerization process since they cause a loss of a portion of the charging stock as well as increasing catalyst consumption by the reaction of fragmental material with the catalytic agent to form sludge-like materials. The process of the present invention overcomes these disadvantages `by utilization of more recently developed catalysts and thus the use of the process along with these catalysts results in the attainment of isomerization reactions which have hereinbefore been available to the petroleum industry.
As stated hereinabove the process of the present invention is particularly directed to the isomerization of hexane hydrocarbons. Hexane hydrocarbonsboil from about 40 to about 90 C. (m4-185 F.) and analyses of typical hexane fractions show that they contain varying quantities ofcyclopentane, 2,2-dimethylbutan e, 2,3- dimethylbutane, `Z-methylpentane, 3-methylpentane, nhexane, methylcyclopentane, dimethylpentanes, cyclohexane, and benzene. The boiling points `and octaneunurnbers vof these six carbon atom hydrocarbons are given in the following table:
By the process of 'this invention the high octane number components are not passed to the isomerization` reactor but are fractionated from the fresh feed. Such cornponents include the 2,2-dimethylbutane, `2,3-dimethylbutane, methylcyclopentane, benzene, and cyclohexane. When processing hexanes in the process of this invention, the reaction zone feed will comprise 2-methylpentane, 3-methylpentane, and n-hexane. reaction zone feed will be converted therein to an equilibrium mixture of hexane isomers including the dimethylbutane cornpounds. In this'manner a total gasoline pool havingan F-l-l-3 cc. octane number of greater than 100 is produced. v
In one embodiment the present invention relateshto a process for the isomerization of an isomerizable saturated hydrocarbon fraction, characterized hy'an average molecular weight greater than about 80, to produce a composite hydrocarbon fraction having an F-l-l-B cc. octane number of at least 100, said isomer-ization being carried out in the presence of an isomerizationcatalyst,
said rst mentioned hydrocarbon fraction `in combination with liquid phase isomerization zone eluent produced as hereinafter described, fractionating said hydrocarbons to produce a substantially naphthene free overhead `hydrocarbon fraction and to produce a substantially pure cyclics hydrocarbon bottoms fraction, removing said substantially pure cyclics hydrocarbon bottoms fraction having an F-l-l-S cc. octane number of at least as one product from the process, passingf said substantially naphthene free overhead hydrocarbon fraction from said firstY fractionation zone to a second fracjdi''rrailonsarefisoinerizecl to higher octane number 'hydroc'arbons, recycling tlieisonieriz'ation'zone eliuent to 'thefrstnientioned fractionation zone as aforesaid to'recoverthe *hig'h octane number hydrocarbons therefrom rs vproduct'andto separate low"octanejnumber hydroey'arbgns'as" aforesaid for reuse fandv internal recyclel in jsaid procesa "and v'compositiilg 'the f aforesaid i product streams. l Afspecicernbodiment ofthe invention relates to a ce'ss uiforth'e"isor'ne'rrization lof an isomerizable hexane "friction to 'proluce `a Composite hexane fraction having jf'niFi-'l-l-:Bcc octane number of'at "least 100, said isoime'riza'tio'n lbeing carried-out in the presence of hydroge'n and nisomeriz'ation catalyst comprising platinum,
`lfal'uini'na,"and combined halogen, which comprises pass- "ingitoa 'first vfractionation zone said first mentioned Yhex-` `ane'"hydro'carbon fraction in combination with liquid phaseisomerization zone eluent producedl as hereinafter escribed, ylfra'ctionating -said hydrocarbons to produce lsiibstantiallyr naph'tliene` free yoverhead hydrocarbon frac- 'i-tin and tofproducea substantially pure cyclics-hydrocarbon bottoms fraction containing methylcyclopentane and cyclohexane, removing said vcyclics hydrocarbon bot- `toms fraction containing methylcyclopentane and cyclo- 'hexane as one product from the process, passing said ,substantially naphthene free overhead hexane fraction from'lsaid rst fractionation zone to a lsecond fraction- .,ationzone as feed therefor, fractionating said hydrocarbons to produce an overhead hydrocarbon fraction containing 2,2 and 2,3-dimethylbuta'ne's and to produce a -bottoms hydrocarbon fraction containing methylpentanes Vand n-hexane, removing saidov'e'rhead hydrocarbon fraction as a second product fromv the process, passing said last mentioned bottoms hydrocarbonfraction containing -methylpentanes and n-hexane to an isomerization zone ti'oned fractionation'zone as aforesaid to recoverdimethylbutanes -therefrom as one product and to separate methylpentanes andn-hexane for reuseand internal recycle `in said process, and compositing the aforesaid v i product streams.
' The combination process ofgthis invention has diverse advantages, all interrelated. The process prevents a buildup of heavier products in the combined feed to the reaction zone since these heavier products are eliminated by removal in the bottoms from the first fractionation zone-"along with methylcyclopentane and cyclohexane. k"11n-a preferred embodiment the isomerization process of "thefpresent invention-is carried out'in the presence of fhydogen. "In'the Vpresence ofhydrogen, cracking of yfnaphthenic hydrocarbons occurs readily due to the acid )function ofthe preferred catalysts. Thus the process of libe-present"inventionA preventsdestruction of C6 naph- "-thenichydrocarbons"by removal'thereof prior to isomeri- "zationfin thelow'erportion of the lfirst fractionation u Y By elimination of the naphthenes from the` reractiomzonecombined feed, conversion Yof `methylcyclo- -pentaxleA to"cyclohexane due to equilibrium considerationsfis prevented. This :is a: -denite advantage in the "overall pool octane numberl sinceV methylcyclopentane "hasta" 'high'er octane number than cyclohexane. f Benzene "which"is"naturally occurring in 'many CffhydrocarbonY ffractinsis hydrogenated tocyclohexane inL the presence typical catalysts `utilized Vin tliisv process. The benfzele content of itheractinf z'o'e Teeds'substantilly fpfes'entin 'substntial quantities in the'iea'c'tion zone deed. 'Another advantagefofE the processof-the present invention? is that @catalystV deactivation lor destruction' is i'ninim'izedfor vsubstantially'"eliminated'by distillation dryfof 1vvate'r V-`:'r"1tained '5 therein along fw'ith the vvdimethylbutanesinthe-'overheadfrom `the second fractionation relativelyfminor 'q1.iantit'ies,I andv thus arev utilized at inter- Ttions. l'Ih'esehex'ane fractionsmaybe obtainedr fromV -of Athe -naturally pr'esentaphtlenes i-f zone feed-fand by. reduction of the'sbenzeiiecontent of this f feed, -hydrogen 4661is'tlrztp't.io'n "-initle'prcessf isminimized. This yresults in ang-economicadvantage since large quantities ofwhydrogn` do not have to be furnished to the process. VIVMajorthy'drogen "consumption is that due to hydrogen solubility in the reaction zone eluent in the high pressure separator, described hereinafter. The Vprocess v prevents *iso-called reverse Zisomerization of 'high octane nurnber "dimethylbutanes to other"hexane Visomers by separation of these high Aoctane number cornponentsiprior to'isomeriz'ation. Atany particular isomeiization'reaction temperature,the formation of hexane 'isomers 'is limited by theequilibrium distribution-of these isomers. `By yhaving the fhighoctane number dimethyl- Vbutanes present 1 in rthe l reactionA zo'nef feed, the 'conversion of fmonomethylpentanesand n-hexane to these compounds cannot'proced-toftlje snie extent/as whenthese isomers are absent.v Furthermore, the equilibriumarnong hexane isomers H'is dynamic sothat'tl'e desired dimethylbutane isomers "can :undergo l'everse isomerization if ingwof therea'ctr` fee'd'a'nd'rejection of srnall quantities zone.v @Thusthe Vreaction zone'- feed is ra fractionation ',zon'eb'ottoms productand'by this means full advantage Mis: taken of'fmaxinium distillation-drying. `Al1 isomeriza-Y *tioncatalysts depend `ponfzin acid function to accom- !plish `he :desiredrreacti'o'm lThis acidfunetion is de- ,Stmmw ,t 4wherein an equilibrium, mixture of `hexane hydrocarbons i ye 'or' decreased by 'Contact Wlth water" The pre vincluding `2,2- and'2,3dimethylbutanes is produced,re V'cycling the isomerization zone eflluent to the first menferred l catalysts fof" the 2pre'ser'lt invention which operate at elevated temperatures contain jthis acid function in mediate"v processing temperatures.. Thus small amounts of 'waterfwhich can be'brought in'with the feed stock, 'cause'.relatively/"rapid catalyst "deactivation Utilization vof the'u'process 'of they present inventionj prevents this deactivation. By' eliniinationof dimethylbutanes, methyl-V `cyclopentarne, and cyclohexane from'` the reaction zone 'fjeed; the `qant'i'ty of feed is' reduced-as well as the coml binedifeed ratio Vand thusY 'the' total investment' necessary 55 for"ca'talyst"'canbe"'substantially reduced. These and other advantages willfbe "eiiplained`-more fully in the vfollowing detailed Vdes'cript'icmof the' process of the presi ent invention.
As set vforth 4hereinabove Vthis invention relates to `a process V:forl the isomerization of "an isomerizable saturated hydrocarbon 'fraction characterized by'ran average moleculanweight-greater than about'SOQ `Hydrocarbons within'the scope 'of thel-aboveflimitation and'utilizable in` vthe process of this invention include methylcyclopentane,` cyclohexane, n-hex'ane, Y 2m`e`thylpentane, 3-methyl ticulafrlyy applicable to'V the isomerization off-*hexane -fracvarioussonrcesincluding fractionationlfrom straight run 'gas'olinaA straight rim1 naphtha, natural gasoline, catalyti- 76 eally'refr'f'ied hpltha, andfcatalytic'ally reformedl'gasol. Analyses of typical hexane fractions from various crude sources are given in the following Table 1I:
Table II.-Analyses of typical Cs fractions [Fractlonatlon out 40-74" O. (10i-165 F.)
- Barnlnus Pennsyl- South Gulf Okla- Natl N atl Composition, Vol. Percent Crude vanta Louisiana Coast homa Wyoming Arabian Gasol Gasol Texas Texas Cyclopentane 2 2 3 3 4 2 2 5 2,2dimethylbutane. Tr 3 2 3 1 1 1 2 3 2,3-dlmethylbutane 5 5 5 5 2 3 3 4- 6 2-methylpentaue. 20 27 25 25 22 22 22 28 28 3-methylpentane- 19 18 18 20 15 14 16 17 15 n-Hexane 32 36 31 30 40 35 44 34 23 Methylcyclopentane.- 16 6 14 3 14 18 7 9 14 Dimethylpentanes-- Tr Tr 1 1 1 1 Tr Tr Oyelohexane 1 Tr 1 7 1 1 2 1 2 Benzene 2 2 1 3 2 1 2 3 4 :Total 100 100 100 100 100 100 100 100 100 Os Oyclics 19 8 16 13 17 20 11 13 20 Octane N o. F-l +3 co 88 86 88 90 85 86 82 86 92 A detailed description of the processing of one of these hexane hydrocarbon fractions by the process of the present invention will be given hereinafter.
Various isomerization catalysts are utilizable Within the generally broad scope of the process of the present invention. These catalysts include a support, an acidacting function, and a hydrogenation component. The support may be selected from diverse refractory oxides including silicia, alumina, silica-alumina, silica-aluminamagnesia, silica-alumina-zirconia, silica-zirconia, etc. Depending up'on the method of preparation and upon the treatment of the support thereafter, these various supports will have surface areas ranging from about 25 to about 500 square meters per gram. In some of these supports the acid-acting function is inherently present, as when silica-alumina is used as the support. The amount of effectiveness of this acid-acting function is then controlled by the quantity of silica which is combined with the` alumina, and by the treatment of the silica-alumina, particularly by calcination, prior to or after compositing the hydrogenation component therewith. Of the various supports, alumina is preferred, and particularly gammaalumina having a surface area of from about 150 to about 450 square meters p'er gram. When gamma-alumina is utilized as the support, the acid-acting function can be added to the catalyst by the incorporationtherein of what is known in the art as'combined halogen. The amount of combined halogen can be varied from about 0.01 to about 8% by weight based on the alumina. Of the various halogens which may be utilized, both uorine and chlorine can be used satisfactorily. Thus in an' alumina type catalyst to be utilized -at reaction temperatures of from 750 to about 850 F., about 0.3% by weight of uorine and about 0.43% by Weight of chlorine may be incorporated therein. When it is desirable to utilize the catalyst at lower temperatures, for example, from about 500 to about 750 F., as is the case in the preferred embodiment of the present invention, the combined halogen which will be utilized along with the alumina support is uorine, and this fluorine will be utilized in an amount of from about 2.5% to about 4.5% by weight. The composite will then have the desired hydrogenation component combined therewith. This hydrogenation component will normally be selected from groups VI(B) and VIII of the periodic table or mixtures thereof. Such hydrogenation components include chromium, molybdenum, tungsten, iron, cobalt, nickel, and the socalled platinum group metals including platinum, palladium, ruthenium, rhodium, osmium, and iridium. Of the various hydrogenation components which may be utilized, those of the platinum group metals are preferred, and of these platinum group metals, platinum itself is particularly preferred. The hydrogenation component ofthe catalysts of the present invention will normally be utilized inan amount of from about 0.01% to about` thequantity utilized will range from about 0.01% to about 2% by weight. A particularly preferred catalyst comprising platinum, combined halogen, and alumina will contain 0.4% platinum, 4.0% iluorine, and alumina. Because of equilibrium considerations and because it is often desirable and/or advisable to carry out the isomerization reaction at the lowest possible temperature, for example, from about 300 to about 500 F., catalysts may also be prepared by impregnating` composites such as described hereinabove with a metal halide of the Friedel-Crafts type. For example, an excellent low temperature isomerization catalyst can be prepared by impregnating from about 5 to about 20% aluminum chloride onto a composite of platinum, alumina, and combined halogen.
The isomerization process of the present invention can be carried out `at varying conditions of temperature, pressure, space velocity, and combined feed ratio. The temperature utilized will particular catalyst selected. Thus the temperature may range over a wide range of from about 300 to about 800 F. The pressure will be selected so as to insure vapor phase operation and depending upon theparticular temperature utilized will range from about to about 1000 pounds per square inch or more. Liquid hourly space velocity will range from about 0.1 to about 10 or more, the only limitation being that equilibrium mixtures of isomerized hydrocarbonsl or a close approach thereto shall be obtained in the reaction zone etluent. The combined feed ratio which is defined as the total amount of fresh feed entering the reactor and recycle divided by the quantity of fresh feed will range from 1 up to about 5 or more. Values in bet-Ween these two limits are determined by the concentration ofthe desired high octane number hydrocarbons in the reaction zone elluent. In the preferred embodiment of the present invention, hydrogen is utilized to minimize cracking and to maintain the surface of the catalyst in a carbon free condition. The quantity of hydrogen utilized will range from about 0t25 to about 10 mols or more of hydrogen per mol of hydrocarbon. Hydrogen consumption will be exceedingly small, in the range of from` about 30 to about 100 cubic feet per barrel of hydrocarbon feed.
The process of the present invention can perhaps be best understood by reference to the accompanying draw-l generally be dictated by the` assegnare 8 'The higher boiling saturated hydrocarbons, substantially free from lnapht-henes, as hereinabove described, and-sub ystantiallyu free from; :dimethylalkane hydrocarbons, are
be produced during processing by hydrogenation of aromatic hydrocarbons,
Ordinarily, thesev naphthenie hydrocarbons occur in the feed and their destruction via hydrocrackingis prevented 'by removal at this point.
Furthermore, atpcertain temperatures of operationv methylcyclopentane .is isomerizedto cyclohexane. This is undesirable since methylcyclopentane normally occurs in larger quantities than does cyclohexane, and since methylcyclopentane has a higher 'octane number than does cyclohexane. Thus the hydrocarbon stream from line'3 is fractionated ,in fractionator 4 and the substantially naphthene free portion :thereof separated overhead. This naphthene free hydrocarbon ifractiongpasses through line 5, is condensed in condenser L6, andrrjpasses through line 7 into receiver 8. The ysubstantially;naphthene free hydrocarbons from receiver S are withdrawn therefrom through line 9 by pump `10 vwhich discharges into line 11 and supplies refluxV topfractionationzone 4 by means of liner12. VThe.netsubstantially naphthene free Vhydrocarbon stream :from fractionation zone 4 passes from line 11 `through line 13 to fractionation zone 2l as hereinafter described. i
Fractionation zone 4 is heated by reboiling agportion of the 'higherboiling hydrocarbons which are withdrawn through lines 14 and 15, and are passed through heat vexchangerjlr. The lheated hydrocarbons are thenreturned to a'lowfer portion of fractionation zone 4 through line 17. 'Ihe higher boiling substantially 4pure cyclics hydrocarbon bottoms fraction having an F-l-Jf-K cc. octane number of atleast -100 fis withdrawn from the bottom of fractionation zone 4 through Vlines Vll4 and 18, and is pumped by with the other high octane number fraction hereinafter described.V l
The purpose of fractionationzone 21 is to fractionate the feedvthereto to produce a hydrocarbon Vfraction characterized by a major proportion of the hydrocarbons with at least two methyl substituents per molecule and and F-l-l-3 ce. octane number of at least 100 and to produce `reactionzone feed. The 'overhead product from this fractionation zone 21 thus comprises dimethylalkanes naturally occurring in the feed and dirnethylalkanes from thereactionrzone effluent which have been formed as a result ofthe isomer'iza'tion-of riparans and vrn lonomethylalkanes as setv forth hereinafter. For example, when processing a hexane fraction, the overheadkproduct which is separated Vfrom the process at this point comprises 2,2- and 2,3Ldimethylbutanes naturallyroccurring in the feed and Y2,2- and 2,3-dimethylbutanes from the reaction zone effluent which have been formed as a result of the isomerization'of -n-hexaneV and 2- and 3-methylpentanes as set forth hereinafter. Also, any water which is dissolvedin-the feed or produced during the processing separates overhead atthispoint due to distillation drying. Thus' the bottom'sffrom fractionation zone 21 which pass to the reaction zone Vare for all Vpractical purposes substantiall'yV lwater free.V The above described high octane number dimethylalkane hydrocarbon stream passes from fractionator 21 overhead through line 22, is condensed in condenser-25,'and passes through line 24 to receiver 25. From receiver 25, the hydrocarbon fraction passes through line 26 to pump 27 and into line 28 which provides reflux ,forY fractionator Z1 by means of line 29. The net dimethylparafn hydrocarbon stream is withdrawn throughline 50 vand Vis combined with the other high octane num-ber fraction hereinabove described and both fractions are then withdrawn throughline 31.
- ,P ractionator 21 is heated by reboiling a portion of the higher boiling hydrocarbons, which are withdrawn through linesZ Vand 33 and passed through heatexchanger 34 and the heated hydrocarbons are, then returned into a lower portionof fractionation zone 21 through `:line '35.
pump 19 through line 20 for combination withdrawn from fractionation zone 21 throughV lines 32 .and 36 by Vpump A37. Y.
The `net hydrocarbon stream from fractionator 21 is pumped bypump 37 through line 38 'and Aline 39.V From "line 39 thisrhydrocarbon stream isrheat exchanged with the hot reaction Azone effluent in heat exchanger 40` ,and then passes through line 41 to' heater 42. 'In heater .42,
Vthe stream is heated to the `desired reaction temperature, and then passes through lineA V43 `to reaction zone'44l It is in this reaction zone 44 thatthe-less highly lzyrarrched Chain and Straight chain' hy droiciarbolns,Y are isomerized. to an equilibrium mixture ofi-saturated hydrocarbonfiso.- mers. As will be set forth hereinafter, the hydrocarbons are processed inthe presence of hydrogen whichfis introducedvia lines 52 and 5 4iintoi1ine39. ThisV hydrogen also -hydrogenates any aromatic'hydrocarbonsin'reaction zone 44 and `the thus produced naphthenie `hydrocarbons arethen isomerized to an equilibriumrmixture of the sameduring isomerization of the remaining saturated hydrocarbons. l
As stated hereinabove, the conditions utilized in reaction zone 42 will depend upon the particular isomerization catalyst utilized therein. It was pointed out above that a preferred catalyst in the process of this invention 'is one comprising va platinum group metal, particularly platinum, combined halogen, and alumina. With such a catalyst, the pressure utilized in the reaction zone will range from about to about 1000 pounds per square inch, the temperature will range from about 300 to about 800 FL, and the hourly liquid space velocity will range from about 0.1 to about l0. The hydrogen to hydrocarbon ratio in the reaction zone will range from about 0.2.5 to about 10 mois of hydrogen per mol of hydrocarbons. When the catalyst utilized comprises platinum, alumina, and combined iluorine in an amount of from about 2.5 to about 4.5 weight percent of the latter', the reaction temperature will be lower than the higher part of the above set forth range, for example, from about 500 to about 750 F. When the catalyst comprises platinum, alumina, and lcombined halogen onto which has been impregnated a metal halide of the Friedel- Crafts type, such as aluminum chloride, the temperature will still be lower, for example, from about 300` to aboutr50 F. In some instances it is desirable vand/or advisable to utilize hydrogen halide along with these catalysts and thus the use of hydrogen chloride, for example, is within the generally broad scope of the present invention.
VThe reaction zone eiuent passes from reactor 44 through `line'f45 in indirect heat exchange with `the reaction zone feed through heat exchanger 40 and then passesthrough line 46, is cooled in condenser 47, and passesthrough line `48 to high pressure separator 49. The Yhigh pressure separator 49 is utilized for separating hydrogen from the reaction rzone eflluent, which hydrogen is separated and passed through line Y50 to compressor 51 where its pressure is increased to the desired number of pounds per square inch, and then the hydrogen is discharged through lines 52 and 54 described hereinabove. When a hydrogen halide'is utilized along with the hereinabove described catalystV as a catalyst promoter, it will pass through line 50 to compressor 51 as hereinabove described. Hydrogen and/ or hydrogen-halide for startup and makeup are added through line 53.
The cooled reaction zone effluent after high pressure separation is discharged from separator '49 through line 55 to debutanizer 56.V This debutanizer is a conventional fractionation zone by means of which a light hydrocarbon gas or low boiling cracked products and dissolved H2 is removed from the process. This removal is accomplished through Vline :57 whichwilll includere;
assenso .9 ceiver-and reflux means .not shown. =Debutanizer`.56 is heated by passing 'a portion Aof the` higher 4boiling hydrocarbons through lines '8 and 59 to heat exchanger 60 from which the heatedihydrocarbons are discharged into a lower portion of the debutanizer via line 61. 'I'he net reaction zone etliuent is `withdrawn from debutanizer 5,6 through lines 58 and 62 by means of pump 63 which discharges the liquid reaction zone eflluent through lineZ to fractionator 4 for fractionation as described hereinabove.
EXAMPLE I One specific example of .the operation of the process with a platinum-aluminacombined halogen catalyst and as the process is carried outsimilar to that set forth hereinabove with reference to thedrawing'is described herewith nowin connectionwith the drawing. The catalyst utilized comprises alumina containing 0.4% platinum and about 0.5 %v combined halogen. The combined halogen comprises about 0.3% uorine and about\0.2% chlorine. This example` illustrates the isomerization of a Gulf Coast hexane fraction, boiling point 40-74 C. The composition of the Gulf Coast hexane fraction is as follows: cyclopentane, 3%, `2,2-dimethylbutane, 3%; 2,3- dimethylbutane, 5%; Z-methylpentane, 25%; S-methylpentane, 20%; n-hex-ane, 30%; methylcyclopentane, 3%; dimethylpentanes, 1%; cyclohexane, 7%; andiA benzene, 3%. Referring again to the drawing this hexane fraction having an average molecular weight greater than 80 inthe quantity of 1000 barrels per day is passed as a liquid under pressure through line l1 and through line 3 to an upper portion of fractionator 4. This fresh feed stream has combined therewith in line 1, `3830 barrels per day of recycle reaction zone elfluent from line 2 as hereinafter described. 'This recycle hexane stream consisting of the liquid reaction zone eiuent contains 452 barrels per day of 2,2-dimethylbutane, 298 barrels per day of 2,3-dimethylbutane, 1310 barrels per day of Z-methylpentane, 8,62 barrels per'day of B-methylpentane, and 908l barrels per lday of n-hexane. Thus the total combined `feed to fractionator 4 amounts to 4830 `barrels per day comprising 30 barrels per day ofcyclopentane, 482 barrels per day of 2,2-dimethylbutane, 348 barrels per day of 2,3-dimethylbutane, 1560` barrels per day of 2-methylpentane, 1062 barrels per day of S-methylpentane, 1208 barrels per day of nhexane, 30 barrels per day of methylcyclopentane, barrels per day of dimethylpentane, 70 barrels per day of cyclohexane, and 30 barrels per day of benzene. The combined feed to fractionator 4 is fractionated therein under a 4:1 molal` reflux to feed ratio and a substantially naphthene free stream -is separated overhead therefrom in an amount of 4690 barrels per day. This 4690 barrels per day contains 30 barrels per day of cyclopentane, 482 barrels per day of 2,2-dimethylbutane, 348 barrels per day of 2,3- dimethylbutane, 1560 barrels per day of 2-methylpentane, 1062 barrels per day of 3methylpentane, and 1208 barrels per day of nhexane. This tfractionator` 4 overhead is withdrawn therefrom through line 5, is condensed in cooler 6, and is passed through line 7 to receiver 8. From receiver 8 the liquid is withdrawn through line 9 by pump10 and supplies feed to fractionator 21, described hereinafter, via line 13. i
Fractionator- 4 bottomsare withdrawn from line 14 through line .18 andhave a compositori as follows: 3Q barrels' per day of methylcyclopentane, 10 barrels'per day of dimethylpentanes, 70 barrels per day of cyclohexane, and 30 barrels per day of benzene. The F1+3 cc. octane number of this stream is 103.7. These bottoms are then pumped by to storage. q
The feed stream to `fractionation `zone 21 in the quantity of 4690 barrels per day and of the composition describedv hereinabove is fractionated therein molal redux to feedjratio` and `the overhead passed therepump19 through lines 20 and 31` per day of n-hexane.
as a liquid through line24 to receiver 2S. Fractionatio zone 21 `separates a substantially monomethylpentane and n-hexane free dimethylbutane stream overhead therei from in an amount of 860 barrels per day. This `860 barrels per day contains 30 barrels per day of cyclopen tane, 482 barrels per day of 2,2-dimethylbutane, `and 348 barrels per day of` 2,3-dimethylbutane- This stream has an F-ll3 cc. octane number of 112.0. This fractionator overhead `is withdrawn from receiver 25 through line 26 by pump 27 andthe net product is pumped through lines 30 and 31 to storage. In storage this stream with its F-14-3 cc. octane number of 112.0 is blended with the aforementioned 'bottoms stream fromfractionator 4, said lattery stream having `an F-14-3 cc. octane number of 103.7, the combined pool thus having an F-l--S` cc. octane number of greater than 100.
The substantially dimethylbutane free bottoms stream from fractonator 21 in the quantity of 3830 barrels per day is the reactor feed. The composition of this stream is as follows: 1560 barrels per day of 2methylpentane, `1062 barrels per day of-3-methylpentane, 1208 barrels This bottoms stream is pumped by pump 37 through lines 38 and 39, where it is joined with hydrogen from line 54, through heat exchange zone 40, and line 41 to heater 42. The quantity of hydrogen which is continuously supplied is sufficient to maintain a hydrogen to hydrocarbon mol ratio of 2:1 in the reaction zone. This is accomplished by recycling hydrogen from the high pressure separator back to the reaction zone. A small quantity of substantiallyk pure makeup hydrogen is added through line S3 to make up for the hydrogen con sumed in the reaction and for that which is dissolved in the separator liquid efuent. By means of heat exchanger 40 and heater 42, the combined feed to thereaction zone is heated to a temperature of 800 F. The reaction is carried out at a pressure of about 500 p.s.i.g. in vapor phase at a liquid hourly space velocity of about 2.0. As set forth hereinabove, isomerization of the reaction zone, feed is accomplished in reaction zone 44 in the presence of the hereinabove described catalyst and results in the production of 3830 barrels per day of reaction zone effluent containing 452 barrels per day of 2,2- dimethylbutane, 298 barrels per day `of 2,3-dimethylbutane, 1310 barrels per day of 2-methylpentane, 862 barrels per day of 3-methylpentane, and 908 barrels per day of n-hexane.
The reaction zone effluent after heat exchange with incoming reaction zone feed is passed via lines 45 and 46 to condenser 47, and the liquid product and hydrogen gas are then passed through line 48 to high pressure separator 49. The hydrogen gas and minor amounts of low boiling hydrocarbons formed by cracking in the reaction zone are separated in high pressure separator 49 and passed through line 50 to compressor 51. High pressure separator liquid which discharges through line 55 is passed to debutanizer 56 to further separate this small amount of cracked products produced during isomerization and to prepare the liquid recycle stream as hereinabove set` forth. The liquid recycle stream passes from debutanizer 56 through lines 58 and 62 to pump r63 wherein it is recycled to fractionation zone 4 via line 2.
In the above example, the small amount of cracking which is observed during processing has not been taken as the process into account. This Yis due to the fact that liquid volume yields based on the percent or higher. As stated hereinabove, the .F-l-f-3 cc. octane number ofthe dimethylbutane -fraction is `112.0 Iand the F1+3 cc. octane number of the cyclics fraction is 103.7, resulting in a pool F-l +3 cc. octane number of greater than 100.
` EXAMPLE II i Another specific-example `ofthe operation of the process with a" platinum-aluminafcoinbinedshalogen i catalyst i an carriedsft out` t that yfeed are in the order of 98 volume.
. hereinabove with reference to the drawing is described lagain'herewith inconnection With the drawing. The catalyst utilized comprises alumina ,containing 0.4%` platinum and about 4.0% combined iluorine.
This example'illustrates the isomerization of the Gulf Coast hexane fraction of the first specific example. This hexane fraction in they quantity of 1000 barrels per day isfpas'sed as a Vliquid under pressure through line 1 and through line 3 to Yan upper portion of fractionator 4. This' l000barrels per day is made up of 30 barrels per day of cyclopentane, 30 barrels per day of 2,2-dimethylbutane, 5,0 barrels per day of 2,3-dimethylbutane, 250 barrels per day of 2 rnethylpentane, 200 barrels per day of 3-meth`ylpentane, 300 barrels per day of n-hexane, 30 barrels per day of methylcyclopentane, barrels per day of dimethylpentanes, 70 barrels per day yof cyclohexane, and 30 barrels per day of benzene. This fresh feed stream has combined therewith in line 1, 3290 barrels per day of recycle reaction zone efuent from Iline 2 as hereinafter described. This recycle hexane stream consisting of liquid reaction zone eiuent contains 494 barrclsper day of 2,2-dimethylbutane, 257 barrels per day of 2,3- dimethylbutane, 1085 barrels per day of 2-methylpentane, 727 barrels per, day of 3-methylpentane, and 727'barrels per day of n-hexane. Thus the total combined feed to fractionatorl amounts to 4290 barrels per day consisting of 30 Abarrels per vday of cyclopentane, 524 barrels per day of 2,2-dimethylbutane, 307 barrels per day of 2,3- dimethylbutane, 1335 barrels per day of 2-methylpentane, 927 barrels per day of 3methylpentane, 1027 barrels per day of n-hexane, 30 barrels per day of methylcyclo-V pentane, 1,0 barrels per day of dimethylpentanes, 70 barrels per day of Vcyclohexane, and 30 barrels per day of benzene. The combined feed to fractionator 4 is fractionated thereinl under va 4:1 molal reux to feed ratio and a substantially naphthene Vfree stream is separated overhead therefrom in an amount of 4150 barrels per day. This -4150 barrels per day contains 30 barrels per day vof cyclopentane, 524 barrels per day of 2,2-dimethylbutane, 307 barrels per day of 2,3-dimethylbutane, 1335 barrels per day of Z-methylpentane, 927 barrels per day of S-methylpentane, and 1027 barrels per day of nhexane. This fractionator 4 overhead is withdrawn therefrom through lineY 5, is condensed in cooler 6, and is passed throuh line 7 to receiver 8. From receiver 8 the liquid is withdrawn through line 9 by pump 10 and supplies feed to fractionator 21, described hereinafter, via lne'13`.
'Fractionator 4 bottoms are withdrawn through line 14 andthrough line 18 and have a composition as follows: 30 barrels per day of methylcyclopentane, 10 bar-A rels per day of dimethylpentanes, 70 barrels per day of cyclohexane, andv 30 barrels per day of benzene. The F-1+3 ,octane number of this stream is 103.7. These bottomsare then pumped by pump 19 through lines 20 and 31 to storage.
lThe feed stream to fractionation zone 21 in the quantity of 4150 barrels per day and of the composition describedhereinabove is fractionated therein under a 4:1 molal redux to feed ratio and the overhead passed therefrom through line 22, condensed in cooler 23, and passed asa liquid through line 24 to receiver 25. Fractionation zone 21 separates a substantially monomethylpentane and n-hexane free dimethylbntane stream overhead therefrom in an amount of 86,1V barrels per day. This 861 barrels per day contains30`ibarrels penday of cyclopentane, 524 barrels per day of 2,'2-'dimethylbutane, and 307 barrels per day' of 2,3-dimethylbutane. This Vstream has an F-1+3 cc. octane number of 109.8. This fractionator overhead is withdrawn from receiver 25 through line 26 by pump 27 and the net product'is pumped through lines 30 Aand 31 to storage. In storage, this stream with its F.l+3 cc..octane.number of 109.8 is blended with the aforementioned vbottomslstream for fractionator .4, said 'latter ystream havingV an F-l-PS ce. octane '1.1wr1br"' f v '103.7,`fthe combined pool thus having an 'F-l-t-B ce. octanenumberof greater than 100'. f'
Thefsubstantia-lly Y 1 Y from `fractonator '21l in the quantity of 3289 barrels-'per day is thelreactor feed. The composition of this s'tream is as `followsfl335 barrels Vper day'ofV 2-methylpentane, V927 barrels-per day of 3methylpenftane,"and I027"ba'rrels per day of n-hexane.Y 'This bottoms stream is pumped by pump 37 through lines 3 8 and 39, where it; is? joined with hydrogen from line 54, through heat exchange zone 40, and line 41 to heater 42. `T he quantity of hydrogen which is continuously suppliedis-suicient tomaintain-a hydrogen toA hydrocarbon mol ratio of 2:1 inthe reacl tion zone. This is' accomplished by recycling'hydrogen from the high pressure separator back to the yreaction zone. -A small amount of substantially 'pure'makeup hydrogen is added through line '53 tei-make upj-for-the hy'- drogen consumed in the reaction and for that whichlis dissolved in theseparator liquid ethuent. By 'means of heat exchanger 40 and heater 42, the combined feed to the reaction zone is heated to a temperature ofV 600 F. TheY reaction is Vcarried out at av pressure' of about 400 p.s.i.g. in vaporphase at an hourly liquid space'velocity of 2.0. As set forth hereinabove',isomerization of the reaction zone feed is accomplished in reaction zone 44 in Vthe presence of the hereinabove described catalyst and results in the production of 3289 barrels per day of reaction zone effluent cont, 'ng 494 barrels per vday of 2,2-dimethylbutane, 257 barrels per `day of '2,3-dimet`hyl butane, 1085 barrelsper day of Y2-rnethylpen'tane, V727 barrels per day of B-methylpentane, 'and 727. barrels per day of n-hexane. Y f
The `reactionf-'zone -efuentafterV heat exchange with the incoming reaction zone feed is passed via lines'45 and 46 to condenser 47, and the liquid `product and;hydrog en gas are then'passedthrough line 48 to high pressure sepa,- rator 49;- k'The Vhydrogenl gas andjminor amounts of iovv boiling hydrocarbons formed by cracking in the reaction ozone areV separated in high pressure separatorV 49 and passed through vliue50 to compressor (5,1. High pressure separator liquid v which discharges through line 55 is passed to .debutanizer Y56 to furtherseparatethe small amount ofunreacted products produced duringisomerization and to prepare the liquid recycle stream jas hereinabove set forth. The liquidrecycle stream passes from debutanizer ,'56 through lines 58 and 62 to pump 6,3 wherein it is recycled to fractionation zone 4 via line-H2.
Inft'hi's-seco'nd example, as previously, the small amount of cracking which is observed during processing has not been taken into account. Here again the liquid volume yields basedV on the feed are in the order of v98 volume' percent or higher.V As stated, the F-lft-VSV cc. octane number Vof the dimethylbutane fraction isV 109.8 and the F-1}3 cc. octane number of the cyclicsfra'ctioriisy 1013.7, resulting in a pool llF-1l3 cc. octane Ynumber of greater' than 100.Y
EXAMPLE III' A still further specific example of the operation ofthe process with a platinum-aluminafcombined halogen catalyst containing aluminum chloride and las theproess is carried out similar to that set `forth hereinabove 'with reference to the drawing is describedherewith.Y 'The catalyst utilized comprises alumina containing Y0.4% platinum,
about 0.3% uorine, and about 0.2% .hlOrina said oomposite having been impregnated with about 17% Aby weight of'aluininum chloride.
In this example the sameGulf VCoast'htptane fraction is isomerized. Referring again to the drawing, thisheirane fraction Ahaving anV average vmolecular ,weight greater than ,i,n the quantity of 100,0,.barrels per day lis passed as a liquid under pressure through line 1 Aand through line 3 to an upper portion of fractionation ,zone 4Q This 1,0 00 barrels per day of .Gulf Coast hexane'fnactioncontams 30 barrels per day of cyclopentane, 30 barrels per dimethylbutanefree bottoms streaml d by pump 27 and day of 2,2dimethylbutane, SObarrels per day of 2,3-dimethylbutane, 250 barrels per `day of Q-methylpentane, `200 barrels per day of S-methylpentane, 300 barrels per `day of n-hexane, 30 barrelsper day of' methylcyclopentane, barrels per day of dimethylpentanes, 70 barrels per day `of cyclohexane, and 30 barrels per day of benzene. This feed` stream has combined therewith'in |line 1, 2400 barrels p er day of recycle reaction zone effluent from line 2 as hereinafter described. This recycle `hexane stream consistingy of the liquid reaction zone effluent contains 495 barrels per day of 2,2-dimethylbutane, 254 `barrels per day of 2,3-dimethylbutane, 818 barrels per day of 2-methylpentane, 475 barrels per day of 3methylpentane, and 358 barrels per day of n-hexane. Thus -the total combined feed to fractionator 4 amounts to 3400 barrels Vper. day comprising 30 barrels per day of cyclopentane, 525 barrels per day of 2,32-dimethylbutane, 304 barrels per day of 2,3-dimethylbutane, 1068 barrels per day of Z-methylpentane, 675 barrels per day of 3methylpentane, 658 barrels per day ofA n-hexane, 30 barrels per day of methylcyclopentane, 10 barrels per day of dimethylpentanes, 70 barrels per day of cyclohexane, and 30 barrels per day of benzene. Thecombinedfeed to fractionator 4 is fractionated therein under a 4:1 molal reflux to feed ratio and a substantially naphthene free stream is separated overhead therefrom in an amount of 3260 barrels per day. 'This 3260 bar,- rels per day contains 30 barrels per day of cyclopentane, 5,25 barrels per day of 2,2-dimethylbutane, 304 barrels per day of 2,3-dimethylbutane, 1068 `barrels per day of 2methylpentane, 675 barrels per day of 3-methylpentane, and 658 barrels per day of n-hexane. This fractionator 4` overhead is Withdrawn therefrom through line 5, is condensed in cooler 6, ,and is .passed through line 7 to receiver 8. From` receiver 8 the liquid is withdrawn through line 9 by pump 10 and supplies feed to fractionator 21, described hereinafter, via line 13.
Fractionator4 bottoms are withdrawn from line 14 through line 18 and have a composition as follows: 30 barrels per day of methylcyclopentane, 10 barrels per dayof dimethylpentanes, 70 barrels perV day of cyclo-V hexane, and 30 barrels per day of benzene. The F-1-l-3 cc. octane number of this stream is 103.7. These bottoms are then pumped by pump 19 through lines 20 and 31 topstorage. The feed stream to fractionation zone "21 in the quantity of 3260 barrels per day and of the composition described hereinabove is fractionated therein under a 4:1 molal reux to feed'iratio and the overhead passes theref fron1`.through line 22,condensed in cooler 23, and passed as a liquid through line 24 to receiver 25. Fractionation zone 21 separates a substantially monomethylpentane and n-hexane free dimethylbutane stream overhead therefrom in an Iamount of 859 barrels per day. This 859 barrels per day contains 30 barrels per dayof cyclopentane, 525 barrels vper day of 2,2-dimethylbutane, and 304 .barrels per day` ofQ-dimethylbutane. This stream has an F-1-i-3 cc.' octane number of 109.8. This fractionator overheadis withdrawn from receiver 25 through line 26 the net product is pumped through lines 30 and 31 to storage. In storage this stream with its F1+3 cc. octane number of 109.8 is blended with the aforementioned bottoms stream from fractionator 4, said latter stream having lan F-l-l-3 cc. octane number of 103.7, the combined pool thus having an F-l-i-S cc. octane number greater than 100.
The substantially dimethylbutane free bottoms stream from fractionator 21 in the quantity of 2401 barrels per day is the reactor feed. The composition of this stream is as follows: 1068 barrels per day of 2-methylpentane, 675 barrels per day of S-methylpentane, and 658 barrels per day of n-hexane. This bottoms stream is pumped by pump 37 through lines 38 and 39, where it is joined With hydrogen from line 54,` through heat exchange zone 1 l 40 and line 41- to heater 42. which is continuously supplied is sufficient to maintain a hydrogen to hydrocarbon mol ratio of 1:1 in the reaction zone. This is accomplished by recycling hydrogen from the high pressure separator to the reaction zone. A small amount of substantially pure makeup hydrogen is added through line 53 to make up for the hydrogen consumed in the reaction and for that which is dissolved in the separator liquid` eiuent. By means of heat exchanger 40 and heater 42, the combined feed to the reaction zone is heated to a temperature of 350 F. The reaction is carried out at a pressure of about 300 p.s.i.g. in vapor phase at an hourly liquid space velocity of 0.5. As set forth hereinabove, isomerization of the reaction zone feed is accomplishedA in reaction zone 44, in the presence of the hereinabove described catalyst and `results in the productionof 2400 barrels per day of reaction zone effluent containing 495 barrels per day of 2,2-dimethylbutane, 254 barrels per day of `2,3-dimethylbutane, 818 barrels per day of Z-methylpentane, 475 barrels per day of `3methylpentane, and 358 barrels per day of nhexane. i
`The reaction zone efuent after heat exchange with incoming reaction zone feed is passed via lines 4S and 46 to condenser `47, `and the liquid product and hydrof gen gas are then passed through `line 48 to high` pressure separator 49. The hydrogen gas and minor amounts of low boiling hydrocarbons formed by cracking in the reaction zone areseparated in high pressure separator 49 and passed through line 50 to compressor 51. High pressure separator liquid` which dischargestthrough line 55 is passed to debutanizer 56 to further separate the small amount of cracked products producedV during isomeriza- The quantity of hydrogen tion and to preparethe liquid recycle stream as herein-` above set forth. The liquid recycle stream passes from debutanizer 56 through lines 58 and 62 to pump 63 wherein it is recycled to fractionation zone 4 via line 2.
In the above example the small amount of cracking which is observed during processing has not been taken into account. Here again, the liquid volume yields based on the feed are in the order of 98 volume percent or higher. As stated hereinabove, the F-1-{3 cc. octane number of the dimethylbutane fraction is 109.8 and the F-l-l-B cc. octane number of the cyclics fraction is 103.7,
resulting in a pool F-l-l-S cc. octane number of `greater` than 100.
-I claim as my invention:
1. A process for the isomerization of an `isomerizable saturatedh'ydrocarbon fraction, characterized by an average molecular weightgreater than" aboutr80,` to produce a composite hydrocarbon fraction having an F-H-S cc. octane number of at least 100, saidisomerization being carried out inthe which comprises passing to a iirst fractionation zone said irst mentioned hydrocarbon fraction in combination with liquid `phase isomerization zone effluent produced as hereinafter described, fractionating said hydrocarbons tol produce a substantially naphthene-free overhead hydroi carbon fraction and to produce a substantially pure cyclics hydrocarbon bottoms fraction, tally pure cyclics hydrocarbon bottoms fraction having an F-1+3 cc. octane number of at least 100 as one product from the process, passing said substantially naphthene-free overhead hydrocarbon fraction from said first fractionation zone to a second fractionation zone as freed therefor, fractionating said hydrocarbons to produce an overhead fraction characterized by a major proportion of hydrocarbons with at least two methyl substituents per molecule and an F-l-l-3 cc. octane number of at least 100 and to produce a bottoms hydrocarbon fraction, removing said overhead hydrocarbon fraction as a second product from the process, passing said last mentioned bottoms hydrocarbon fraction to an isomerization zone wherein low octane number hydrocarbons are presence of anv isomerization catalyst,
removing said substan-l absence isomerized to higher `octane number hydrocarbons, r'ecycling the isomeriz'ation zone -eiuent tothe nrst mentioned fractionation zone as aforesaid to recover the high octane number hydrocarbonstherefrom as product and to separate low octane number hydrocarbons as aforesaid for reuse and internal recycle in said process, blending said secondproduct with said cyclics hydrocarbony bottoms fraction and recovering the resultant blend as said compositefraction.
2. A process for the isomerization of an isomerizable hexane hydrocarbon Ifraction --toproduce a composite hexane hydrocarbon fraction having an F-l-i-S cc. octane number of at least 100, said'isomeriza'tion being carried V4out in the presence lof hydrogen-.and Visornerization catalyst comprising platinum, alumina, and rcombined halogen, Vwhich comprises passingtol airs'tffractionation zone saidV rst mentioned heira'nehydrocarbon fraction in combination with liquid -phase isomerization zone effluent produced as hereinafter`describcd, fractionating said hydrocarbons to produce a substantially naphthene-free overhead hydrocarbon fraction and 'toproduceasubstan- L, tially pure cyclicslhydrocarbon'bottoms fraction containing methylcyclopentane and cyclohexarie, removing said cyclicsV hydrocarbon bottoms fraction containing methylcyclopentane and cyclohexane as one' product from the process, passing said substantially naphthene-free overhead hexane fraction from `said rst fractionation zone to a second fractionation zone as feed therefor, fractionating said hydrocarbons to produce an overhead hydrocarbon fractionvcoritaining 2,2- and 2,3-dimethylbutanes and to producea-bottoms hydrocarbon fraction containing methylpentanes and n-hexane, removing said Voverhead hydrocarbon fraction as a-second product from the process, passing said last mentioned bottoms hydrocarbon fraction containing methylpentanes and n-hexane to an isomerization zone 4wherein an equilibrium vmixture of hexane hydrocarbons including 2,2- and 2,3-dimethylbutanes is produced, recycling the isomerization VZone efuent to the tirst mentionedfractionation zone as aforesaid to recover dimethylbutanes therefrom as one product and to separate methylpentanes vand n-hexane as aforesaidy forreuse and internal .recycle in said process7 blending said ysecond product VWith said cyclics hydrocarbon bottoms"fractionV and recovering the resultantlblend as saidcomposite fraction. Y v v 3. A process for the isomerization of an isomerizable hexane hydrocarbon fraction to produce a composite hexane hydrocarbonpfraction'having :an Ffle'l-Scc'. octane number ofat. le'ast 100, said isomerization being vcarried out the presence vof lhydrogen and #an 'isomerization catalyst comprising, platinum, alumina, and combined halogen, which comprises passingto a first yfractionation .zonef said iirrst mentioned-hexane hydrocarbon fraction in combination'withliquid phase isomerization 'zone' eluent v produced as hereinafter described, `fractio'nating said hydrocarbons to produce a substantially naphthene-free overhead hydrocarbon vfraction and to produce a substantially pure cyclics hydrocarbonbottorns fraction containing methylcyclo'pentaneand cyclohexane, removing said substantially pure ,methylcyclopentanc and cyclohexane i176 i fraction as one product vfrom the hydrocarbon bottoms process, passing said substantially naphthenefree pvercontaining methylpentanes andn-hexane, removinggsaidoverhead hydrocarbon fraction as a second product from the process, passingvsaid 'last mentioned bottom-s hydrocarbon fraction containing methylpentanes and Ynv-herrane to an isomerization zone maintained at a temperature Y of from about 300 Fito about 800 F. `and at a pressure of from about to about 1000 pounds per supiare inch wherein methylpentanes Yand n-hexane are isornen ized in the presence of said catalyst andhydrogen at a liquid hourly fspacefvelocity of from about 0.1 to about 10 to an equilibrium mixture of hexane hydrocarbons vin cluding 2,2- and 2,3-di-methy1but'anes, recycling the `isomerization zone eluent to lthe iirst mentioned fractionation zone as aforesaid torrecovler dimethylbutanestherefrom as oneproduct andato-separate methylpentanes and n-hexane as aforesaid for reuse and/internal recycle in said process, and combiningsaid 'secondproduct vVith said cyclics hydrocarbon bottoms fraction to yield ahexane fraction having an F-1+3 octane' number of at least 100. Y
4. The process of claim 3 lfurther characterized in that the catalyst comprises alumina, from Yabout 0.01% to about 2% by weight thereofof platinum, and fromabout 0.1% to about 8%' by weight thereof of combined Vhalogen.
5. The process of claim 4 further characterizedy in that the combined halogen is a )mixture of chlorine and fiuorine in an amount of from about-'0.3% to about 0.7%
7. The process of `claim 4 further vcliaract'erizedin that the catalyst composite has impregnated thereon fromV about 5% to about 20% b yiweightpf a Friedel-Crafts metal halide and the .isomerization is 4carried opt at a 'temperature of from 3003 .toV *about50p."V F,
8. T he process of claim 7 f urtjherharacterizedin that .the Friedel-Crafts metal halideis aluminum chloride.
References Cited inthetlefof this vpatent UNITED lsri-*frias -PArnrrrsV v Suiten et ai. f r v Feb. 19, 1194s 2,440,751 Legatski Y. May .4, :194.8 2,443,607 Evering June 22, .1948
2,766,302 Elkins Oct.V 9, 1956 2,805,269 Carter et al. Sept. 3, 195-7 2,834,823 Patton et al May 13, 19518

Claims (1)

1. A PROCESS FOR THE ISOMERIZATION OF AN ISOMERIZABLE SATURATED HYDROCARBON FRACTION, CHARACTERIZED BY AN AVERAGE MOLECULAR WEIGHT GREATER THAN ABOUT 80, TO PRODUCE A COMPOSITE HYDROCARBON FRACTION HAVING AN F-1+3 CC. OCTANE NUMBER OF AT LEAST 100, SAID ISOMERIZATION BEING CARRIED OUT IN THE PRESENCE OF AN ISOMERIZATION CATALYST, WHICH COMPRISES PASSING TO A FIRST FRACTIONATION ZONE SAID FIRST MENTIONED HYDROCARBON FRACTION IN COMBINATION WITH LIQUID PHASE ISOMERIZATION ZONE EFFLUENT PRODUCED AS HEREINAFTER DESCRIBED, FRACTIONATING SAID HYDROCARBONS TO PRODUCE A SUBSTANTIALLY NAPHTHENE-FREE OVERHEAD HYDROCARBON FRACTION AND TO PRODUCE A SUBSTANTIALLY PURE CYCLICS HYDROCARBON BOTTOMS FRACTION, REMOVING SAID SUBSTANTIALLY PURE CYCLICS HYDROCARBON BOTTOMS FRACTION HAVING AN F-1+3 CC. OCTANE NUMBER OF AT LEAST 100 AS ONE PRODUCT FROM THE PROCESS, PASSING SAID SUBSTANTIALLY NAPHTHENE-FREE OVERHEAD HYDROCARBON FRACTION FROM SAID FIRST FRACTIONATION ZONE TO A SECOND FRACTIONATION ZONE AS FREED THEREFOR, FRACTIONATING SAID HYDROCARBONS TO PRODUCE AN OVERHEAD FRACTION CHARACTERIZED BY A MAJOR PROPORTION OF HYDROCARBONS WITH AT LEAST TWO METHYL SUBSTITUTENTS PER MOLECULE AND AN F-1+3 CC. OCTANE NUMBER OF AT LEAST 100 AND TO PRODUCE A BOTTOMS HYDROCARBON FRACTION, REMOVING SAID OVERHEAD HYDROCARBON FRACTION AS A SECOND PRODUCT FROM THE PROCESS, PASSING SAID LAST MENTIONED BOTTOMS HYDROCARBON FRACTION TO AN ISOMERIZATION ZONE WHEREIN LOW OCTANE NUMBER HYDROCARBONS ARE ISOMERIZED TO HIGHER OCTANE NUMBER HYDROCARBONS, RECYCLING THE ISOMERIZATION ZONE EFFLUENT TO THE FIRST MENTIONED FRACTIONATION ZONE AS AFORESAID TO RECOVER THE HIGHER OCTANE NUMBER HYDROCARBONS THEREFROM AS PRODUCT AND TO SEPARATE LOW OCTANE NUMBER HYDROCARBONS AS AFORESAID FOR REUSE AND INTERNAL RECYCLE OF SAID PROCESS, BLENDING SAID SECOND PRODUCT WITH SAID CYCLICS HYDROCARBON BOTTOMS FRACTION AND RECOVERING THE RESULTANT BLEND AS SAID COMPOSITE FRACTION.
US658877A 1957-05-13 1957-05-13 Isomerization of saturated hydrocarbons Expired - Lifetime US2938936A (en)

Priority Applications (1)

Application Number Priority Date Filing Date Title
US658877A US2938936A (en) 1957-05-13 1957-05-13 Isomerization of saturated hydrocarbons

Applications Claiming Priority (1)

Application Number Priority Date Filing Date Title
US658877A US2938936A (en) 1957-05-13 1957-05-13 Isomerization of saturated hydrocarbons

Publications (1)

Publication Number Publication Date
US2938936A true US2938936A (en) 1960-05-31

Family

ID=24643079

Family Applications (1)

Application Number Title Priority Date Filing Date
US658877A Expired - Lifetime US2938936A (en) 1957-05-13 1957-05-13 Isomerization of saturated hydrocarbons

Country Status (1)

Country Link
US (1) US2938936A (en)

Cited By (4)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US3215753A (en) * 1963-08-15 1965-11-02 Universal Oil Prod Co Selective isomerization of neohexane
US3392212A (en) * 1964-12-21 1968-07-09 Standard Oil Co Process for producing dimethylbutane from pentane
US3852372A (en) * 1970-06-25 1974-12-03 Texaco Inc Isomerization with fluorided composite alumina catalysts
WO2007059873A1 (en) * 2005-11-22 2007-05-31 Haldor Topsøe A/S C7 isomerisation with reactive distillation

Citations (6)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US2395022A (en) * 1943-03-31 1946-02-19 Standard Oil Co Aluminum halide hydrocarbon conversion system
US2440751A (en) * 1945-11-05 1948-05-04 Phillips Petroleum Co Conversion of hydrocarbons
US2443607A (en) * 1943-03-31 1948-06-22 Standard Oil Co Heptane isomerization
US2766302A (en) * 1952-01-17 1956-10-09 Sinclair Refining Co Isomerization of alkanes and cycloalkanes
US2805269A (en) * 1951-05-22 1957-09-03 Phillips Petroleum Co Process for isomerization of liquid hydrocarbons
US2834823A (en) * 1955-03-07 1958-05-13 Kellogg M W Co Isomerization of hydrocarbons

Patent Citations (6)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US2395022A (en) * 1943-03-31 1946-02-19 Standard Oil Co Aluminum halide hydrocarbon conversion system
US2443607A (en) * 1943-03-31 1948-06-22 Standard Oil Co Heptane isomerization
US2440751A (en) * 1945-11-05 1948-05-04 Phillips Petroleum Co Conversion of hydrocarbons
US2805269A (en) * 1951-05-22 1957-09-03 Phillips Petroleum Co Process for isomerization of liquid hydrocarbons
US2766302A (en) * 1952-01-17 1956-10-09 Sinclair Refining Co Isomerization of alkanes and cycloalkanes
US2834823A (en) * 1955-03-07 1958-05-13 Kellogg M W Co Isomerization of hydrocarbons

Cited By (5)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US3215753A (en) * 1963-08-15 1965-11-02 Universal Oil Prod Co Selective isomerization of neohexane
US3392212A (en) * 1964-12-21 1968-07-09 Standard Oil Co Process for producing dimethylbutane from pentane
US3852372A (en) * 1970-06-25 1974-12-03 Texaco Inc Isomerization with fluorided composite alumina catalysts
WO2007059873A1 (en) * 2005-11-22 2007-05-31 Haldor Topsøe A/S C7 isomerisation with reactive distillation
US20100145128A1 (en) * 2005-11-22 2010-06-10 Sven Ivar Hommeltoft C7 isomerisation with reactive distillation

Similar Documents

Publication Publication Date Title
CA2038824C (en) Combination process for hydrogenation and isomerization of benzene- and paraffin-containing feedstocks
US3788975A (en) Simultaneous production of aromatic hydrocarbons and isobutane
EP0321141B1 (en) Dual function process for isomerization and ring opening of a feed stream containing paraffins and cyclic hydrocarbons
US2651597A (en) Process for improving the octane number of light naphthas
EP0400228B1 (en) Simultaneous ring opening and isomerization of cyclic-containing n-paraffin feed
US2834823A (en) Isomerization of hydrocarbons
US3933619A (en) Gasoline production process
US2698829A (en) Two-stage process for the catalytic conversion of gasoline
US2689208A (en) Hydrocarbon conversion process
US5227554A (en) Isomerization process
US5498810A (en) Selective isoparaffin synthesis from naphtha
US2915571A (en) Isomerization of saturated hydrocarbons
US2946736A (en) Combination process for high-octane naphtha production
US5235120A (en) Selective isoparaffin synthesis from naphtha
US2953612A (en) Catalytic hydrogenation of dripolene
US2938936A (en) Isomerization of saturated hydrocarbons
US2908735A (en) Isomerization of normal butane
US5198097A (en) Reformulated-gasoline production
US3392212A (en) Process for producing dimethylbutane from pentane
US3974061A (en) Isomerization of C5 and C6 isomerizable hydrocarbons
US3650943A (en) High octane unleaded gasoline production
US2349821A (en) Catalytic conversion process
US2349516A (en) Treatment of paraffins
US3153627A (en) Catalytic process
US5401385A (en) Selective upgrading of naphtha