Search Images Maps Play YouTube News Gmail Drive More »
Sign in
Screen reader users: click this link for accessible mode. Accessible mode has the same essential features but works better with your reader.


  1. Advanced Patent Search
Publication numberUS3050456 A
Publication typeGrant
Publication dateAug 21, 1962
Filing dateOct 15, 1958
Priority dateOct 18, 1957
Publication numberUS 3050456 A, US 3050456A, US-A-3050456, US3050456 A, US3050456A
InventorsMelchior Jan G
Original AssigneeShell Oil Co
Export CitationBiBTeX, EndNote, RefMan
External Links: USPTO, USPTO Assignment, Espacenet
Conversion process for the production of high octane number gasoline components
US 3050456 A
Previous page
Next page
Description  (OCR text may contain errors)

2 Sheets-Sheet 1.

MELCHIOR S FOR THE PRODUCTION OF Aug. 21, 1962 CONVERSION PROCES HIGH OCTANE NUMBER GASOLINE COMPONENTS Filed on. 15, 1958 $25 mowmu x5636 1 P 3 1| f 25 220 46 93 23: r

5 w w Y 3 m 3 E C N L R E n O M T n T 2 G A w P F R w E mm mm wmmwwn :0 m3 563 0 N m 3 Al 2.5 .53 i T H 5 N W J 5; rmm w2wm r m B 2. wm u w 1 ll-llllll I I l I I I I I I Ill l How |:l Z E g IIIIFI q l 1% o M536 mo: 9 zo; vj M U33 -mww ow mm mm vm m w m m mm 52%;? ME 1: L 1 m 3 mm 25 220 miwma Sills P o? L mm zkwzwm k Q3 8\ m x H Al w J 6 5 Q Aug. 21, 1962 J. G. MELCHIOR 3,050,455



INVENTOR JAN G. MELCHIOR HIS ATTORNEY United States Patent Ofiice 3,050,456 Patented Aug. 21, 1962 3,050 456 CONVERSIGN PROCESS I OR THE PRODUCTION OF HIGH OCTANE NUMBER GASOLINE COM- PONENTS Jan G. Melchior, The Hague, Netherlands, assignor to Shell Oil Company, a corporation of Delaware Filed Oct. 15, 1953, Ser. No. 767,424 Claims priority, application Belgium Oct. 18, 1957 8 Claims. ((11. 2ti8-67) This invention relates to a process for the alkylation of isoparaffins with olefins for the production of high octane number gasoline blending components.

It is well known to produce high octane number gasoline blending components by alkylating isoparafiins of from four to five carbon atoms per molecule with olefins of from three to five carbon atoms per molecule. A typical alkylation reaction is the addition of butylenes to isobutane to produce saturated branched octanes. The reaction is catalyzed by acids. Concentrated sulfuric acid or anhydrous hydrofluoric acid are the preferred commercial alkylation catalysts.

It is an object of this invention to provide an improvement in alkylation processes. It is a more specific object to provide a method for increasing the total production of alkylate from a fresh feed which is deficient in isoparaffins. It is a further object to provide a method for producing alkylate of improved octane number. Further objects will become apparent from the following description in which reference is made to the drawing, wherein:

FIG. 1 is a schematic representation of an arrangement of equipment used for carrying out -a preferred mode of the process, and

FIGS. 2 and 3 are schematic representations of arrangements of equipment used in the process according to the prior art, and according to this invention, respectively.

In view of the fact that isoparaffin is commonly recovered from alkylation product and recirculated to the alkylation zone it should be understood that in this text fresh feed to the alkylation zone means the alkylation reactants before addition of isoparafiin recycle. The feed which is actually alkylated, which includes isoparafiin recycle, may be designated total alkylation feed. The ratio of isobutane to olefin in the total alkllation feed is known as the external isoparaffin-to-olefin ratio. The ratio of the amount of isoparaffin which is present at the point at which olefin is injected into the reaction mixture to the amount of olefin injected is known as the internal isoparafiin-toolefin ratio.

Saturated alkylation feed compounds are predominantly isobutane and in some cases isopentane. Unsaturated alkylation feed compounds are olefins having from three to five carbon atoms per molecule. The alkylation feed streams can consist of a single olefin or of a mixture of olefins on the one hand and of a single isoparafiin or of a mixture of isoparafiins on the other hand, or of mixed isoparaffins and olefins. Other components may be present, e.g., unbranched parafiins, which do not participate in the alkylation reaction to any appreciable extent but act merely as diluents.

Branched olefins, e.g. isobutylene, react readily with isoparafiins under alkylation conditions. However, in the sulfuric acid-catalyzed process, isobutane-isobutylene alkylate is of lower quality and is produced in lower yield than the alkylate produced from unbranched butylenes. In the HF-catalyzed process, isobutane-isobutylene alkylate is inferior to that produced from butene-Z.

Mixtures of light branched and unbranched hydrocarbons are produced in the petroleum industry by the distillation of crude petroleum and by thermal or catalytic cracking of certain petroleum fractions.

In the distillation of crude petroleum, one can separate a fraction boiling below gasoline which consists mainly of saturated C C hydrocarobns.

In the work-up of thermally or catalytically cracked petroleum fractions one can obtain a light fraction which consists mainly of saturated and olefinic C to C hydrocarbons.

In the stabilization of the product obtained by catalytic reforming of naphtha fractions, one can obtain a light hydrocarbon fraction which consists mainly of saturated C and C hydrocarbons, including substantial amounts of isobutane.

C C, hydrocarbons may also be obtained by other refinery processes.

The predominant sources of isoparafiins for alkylation are isobutane and possibly isopentane which are recovered from the distillation of crude petroleum; further amounts of isobutane and possibly isopentane can be produced by the catalytic isomerization of the corresponding saturated normal paraffins.

The predominant sources of olefins are the light fractions recovered from the work-up of thermally and catalytically cracked petroleum fractions. These fractions are not in themselves suitable as complete fresh alkylation feed stocks because they are deficient is isoparaflin hydrocarbons. They are, therefore, preferably first admixed with other light fractions which are low in olefins, such as the straight-run light fractions. It is often preferable for economic reasons to admix all available light hydrocarbon fractions and fractionate the mixture to recover a C fraction as alkylation feed, rather than fractionating each light hydrocarbon fraction separately to recover a C fraction from each.

It is important in alkylation processes that the hydrocarbon mixture in the alkylation zone contain a substantial excess of the isoparafiin. The external is-oparafiinto-olefin ratio is usually maintained between 3.5 :1 and 8:1 in sulfuric acid alkylation and even higher, e.g. between 6:1 and 12:1 in HF alkylation. The internal isoparaffin-to-olefin ratio is preferably at least 300:1 and may be as high as 800:1 or more. A substantial amount of the isoparaffin present in the alkylation zone remains unconverted while practically all the olefin is combined with isoparaffin.

It isespecially important that the molecular ratio of isoparafiins-to-olefins in the fresh feed to the alkylation zone, including both normal and branched olefins, remain above a specified lower limit. The molecular ratio of isoparafiins to total olefins in the fresh alkylation feed should in every case be greater than 1.0:1 and preferably should be about 1.1:1. Ratios of 1.05 to 1.20 are satisfactory and ratios in the range from 1.1 to 1.15 are especially preferred. When the value of the ratio is less than 1:1, the yield and antiknock quality of the alkylate are substantially lower than at higher isoparafiin-to-olefin ratios, and greater acid losses are observed. This is especially objectionable in the case where hydrogen fluoride catalyst is employed because substantially greater amounts of the rather expensive HF are carried off in the form of alkylfiuorides in the reaction product and are thus lost. When the molecular ratio exceeds 1:1 these undesirable side effects quickly decrease and in each case are substantially negligible at ratios of 1.1:1. Substantially higher values of the ratio may be employed; however, these lead, especially with total recycle of unconverted branched paraflins, to unnecessarily high isoparaffin-to-olefin ratios in the total alkylation feed and thus require an excessively large alkylation zone or reduce the capacity of a given system. This difficulty can be overcome by recirculating only a portion of the total isoparaffins; an excessively high isoparafiin-to-olefin molar ratio in the fresh feed to the alkylation zone is in that case not objectionable.

The present invention provides not only a means for improving the quality of total alkylate but also for increasing the output of total alkylate in a refinery in which the amount of available alkylating olefins is greater than the amount of alkylat-able isoparafiins.

It has been proposed in the past in similar refinery situations to polymerize a portion of the available alkylatable olefins and thus bring the isoparaflin-to-olefin ratio into a desired balance. Obviously, the amount of alkylate produced by this expedient is only that which corresponds to the isoparaflin originally available in the refinery. By the process of this invention, however, a total amount of alkylate can be produced in a given refinery situation which is substantially greater than the amount of alkylate that could be produced by employing the prior art expedient of merely removing some of the olefins by polymerization. This will be further illustrated below by means of numerical examples.

The present invention is concerned with an improvement in the production of alkylate in refineries in which the availability of isoparaflins limits the production of total alkylate. In the past this has not been a serious problem because in the nation-wide economy isobutane has been available in excess and thus a refiner could purchase isobutane on the market if in his refinery sufficient isobutane was not available. It is not expected, however, that this situation will prevail much longer. In Petroleum Refiner, 37, No. 7, 119-123 (1958), Sutherland and Belden show that by 1960 there is expected to be a substantial nation-wide deficiency of isobutane and even a deficiency of n-butane which could be isomerized to isobutane for the production of alkylate.

The process of this invention is particularly adapted for improving the alkylate yield for the refinery in which an olefin fraction is available which has a relatively high isobutylene-to-normal ratio. Catalytic cracking processes, and particularly the so-called two-stage catalytic cracking process which has recently come into commercial use, provide such a C fraction. The two-stage process is de scribed, i.a. in copending patent application Serial No. 436,004, filed June 11, 1954, now abandoned. Suitable apparatus is shown in US. 2,798,795 to Rehbein et al. A typical ratio of branched-tonormal butylenes produced in such a process is between 0.8 and 09:1. This compares with ratios in the range from 0.5:1 to 0.65:1 observed in conventional commercial catalytic cracking processes and from 0.3:1 to :1 in conventional commercial thermal cracking processes. A C; fraction from two-stage catalytic cracking is a preferred feed in the process of this invention; olefinic fractions from other catalytic cracking processes and including a substantial proportion of isobutylene are also suitable.

The present invention provides a method which makes it possible to increase the molecular ratio of branched paraflins to olefins in the fresh feed to an a-lkylation zone to exceed 1:1 to any desired extent. It thus becomes possible, according to the process of this invention, to produce a larger amount of better quality alkylate from a given refinery feed stock having an isoparafifin-to-olefin ratio below 1:1 than would otherwise be possible.

The process of this invention is a method for the product-ion of high octane gasoline components from a fresh feed containing one or more olefins, one or more isoparaflins and possibly one or more unbranched paratfins at an isopar-afiin-to-olefin molar ratio of less than 1:1 by alkylating the feed under the influence of an acid catalyst after a suit-able part of the total fresh feed stream or of at least one of the components of the total fresh feed stream has been taken as a side stream and catalytically hydrogenated, preferably in a reaction zone wherein at the same time normally liquid hydrocarbons are catalytically treated in the presence of hydrogen, the light components of the hydrogenated side stream being recovered and recombined with the remainder of the feed to produce a fresh feed in which the molecular ratio of branched paraffins to olefins is greater than 1:1.

There are several well-known commercial processes for carrying out the alkylation step of this invention. Typical processes for alkyl-ation using concentrated sulfuric acid catalyst are described in Petroleum Refiner, 34, N0. 9, 148 (1955) and Petroleum Refiner, 35, No. 9, 251 (1956). Typical processes for alkylation using concentrated hydrofluoric acid are described in Petroleum Refiner, 34, No. 9, 126 (1955); in Trans. Am. Inst. Chem. Engrs. 39, 793 (1943), and in the book Hydrofluoric Acid Alkylation, Phillips Petroleum Company, 1946.

The sulfuric acid alkylation may be advantageously carried out by the method described in the Petroleum Refiner, supra, or by the methods described, for example, in the United States Patents 2,232,674, 2,260,945, 2,283,603 and 2,370,164; other methods of batch, intermittent or continuous operation may also be used.

The titratable acidity of the sulfuric acid employed as catalyst in the alkylation reactor is generally in the range from 85% to 100% H and preferably between 88% and 94% H 50 It is general practice to charge to the process sulfuric acid having between 96% and 100% concentration and to use it until its titratable acidity has dropped to a lower value, e.g., about to The alkylation reaction is conventionally carried out at temeratures in the range from about 0 C. to about 22 C. and preferably from 4 C. to 16 C. and pressures in the range from atmospheric to 135 p.s.i.g., but sufiiciently high to maintain the reactants in liquid phase. It is generally desirable to employ a volume of liquid acid catalyst phase equal to from about 75% to about 250% of the volume of hydrocarbon phase used. An acid-tohydrocarbon volume ratio of about 1:1 is generally preferred.

In the continuous alkylation processes the fresh feed is admixed with isoparafiin recycle and the total mixture contacted with concentrated sulfuric acid catalyst. The hydrocarbon product, separated from the acid catalyst, is passed to a fractionation zone in which at least some of the following are separated: unconverted branched paraffins; normal paraifins, if present, and crude alkylate. In a typical C alkylation process the total alkylate is deisobutanized and isobutane recycled to the reaction. The cleisobutanizer bottoms is debutanized to remove normal butane and obtain a total alkylate fraction as bottoms and the alkylate is re-run to separate alkylate bottoms, or it may be fractionated into a light alkylate and heavy alkylate fraction.

The alkyl-ation employing concentrated hydrofluoric acid as catalyst may also be carried out in different types of process equipment, typically in the manner described in the publications, supra. The alky-lation may be carried out at a temperature between 0 C. and 65 C. and preferably between 25 C. and 45 C. Acid-to-hydrocarbon ratios are not critical but may be in the range mentioned above for sulfuric acid or lower, down to about 1:10. The concentration of the hydrofluoric acid is in the range between 80% and and suitably between 86% and 90%; care is taken to keep water out of the reaction system.

The total hydrocarbon layer is separated from the HF acid layer in a separator. Part of the acid is recycled to the reaction zone and part is regenerated prior to being returned to the reaction zone. The hydrocarbon product is fractionated to recover isoparaflin for recycle, normal parafiin for removal from the system and alkylation product. Means are provided for separating hydrogen fluoride present in the hydrocarbon phase.

The catalytic treatment in the presence of hydrogen to which the normally liquid hydrocarbon fractions are subjected is preferably a catalytic desulfurization or a catalytic reforming process.

The catalytic desulfurization of normally liquid hydrostraight-run crude oil fractions or fractions which have been subjected to one or more thermal or catalytic treatments are passed at elevated temperatures and pressures in the presence of hydrogen or of a hydrogen-containing gas over a suitable catalyst.

Depending on the kind of material to be desulfurized and on the circumstances, the starting material together with hydrogen or with the hydrogen-containing gas is passed over a catalyst either in vapor phase, in mixed phase, in liquid phase or in super-critical condition. The so-called trickle technique, described in US. 2,608,521 and in Petroleum Refiner, 32, No. 5, 137 (1953), and 34, No. 9, 155 (1955), is particularly suitable for such desulfurization. In this trickle technique the desulfurization feed, partly in liquid phase and partly in vapor phase, is passed co-currently with a hydrogen or a hydrogencontaining gas downward over a fixed bed of catalyst particles whereby the unvaporized portion of the feed passes over the catalyst in a thin film and the vaporized portion diffuses through the thin film to contact the catalyst. In the trickle technique, relatively low hydrogen-tooil ratios can be employed. The following conditions are suitable: pressures of -100 atmospheres, temperatures of 300 C.-500 C., space rates of 0.5 kg. of oil per liter of catalyst per hour, and gas-to-oil ratios of 50500 1. gas per kg. oil.

Suitable desulfurization catalysts are those which contain one or more elements of the sixth and/or eighth group of the periodic system, either as metals or in the form of one or more compounds with one or more other elements and suitably supported on a catalyst carrier. Particularly suitable catalysts are those containing one or more of the elements iron, nickel, cobalt, chromium, molybdenum, or tungsten as metal or in the'form of one or more compounds with one or more other elements, e.g. oxygen or sulfur, suitably on an alumina carrier.

Especially preferred desulfurization catalysts comprise aluminum oxide as catalyst carrier and supported thereon 5%-l5% by weight cobalt or molybdenum as metals and/ or in the form of one or more compounds thereof with one or more of the elements oxygen, sulfur or aluminum, and in which the atomic ratio of cobalt to molybdenum is in the range between 1:20 and 18:20 and preferably between 1:10 and 9:10.

The gas is suitably hydrogen or a hydrogen-containing gas mixture, preferably a mixture of hydrogen and light hydrocarbons; particularly suitable is the hydrogen-rich gas produced in the catalytic reforming of hydrocarbon fractions. It is suitable although not necessary to recirculate the hydrogen from the effiuent, suitably after removing undesired components.

When the catalytic treatment of the normally-liquid hydrocarbons is catalytic reforming this is carried out in the well-known manner. Reforming processes are described in Petroleum Refiner, 34, No. 9, 232-259 (1955). The catalytic reforming treatment, which is endothermic, can be carried out in a reaction vessel or pipe which is provided with means for addition of heat in order to maintain the total reaction space at a suitable temperature. It is also possible to carry out the reaction in a series of separate unheated reactors whereby intermediate streams passing from one reactor to the next are heated to the required temperature. Combinations of these two methods can also be employed. It is an advantage of the process of this invention that the hydrogenation reaction of the light hydrocarbons is exothermic so that a part of the heat of reaction of the catalytic reforming process can be supplied thereto internally by charging unsaturated light hydrocarbons thereto in accordance with the process of this invention.

Catalysts in the catalytic reforming process are employed in the form of fast-flowing or fluidized beds of catalyst powders or they may be employed in fixed-bed 6 form. The naphtha to be treated is passed in vapor phase through the catalyst bed.

Reforming temperatures are in the range between about 420 C. and about 560 C. and preferably between about 470 C. and 540 C.

It is preferable to have relatively high partial pressures of hydrogen in the reaction zone. These partial pressures may be about 2 atmospheres, gauge, or higher, e.g. up to about 50 atmospheres, gauge. The total pressure in the reaction zone may be still higher.

Reforming catalysts are preferably one or more of the metals of the sixth and/ or eighth group of the periodic table as metals and/ or in the form of compounds with one or more other elements such as sulfur or oxygen. They may be supported on a suitable carrier such as aluminum oxide. A preferred type of catalyst consists of molybdenum oxide and/ or chromium oxide supported on an aluminum oxide-containing carrier. At present the most widely used catalysts are those containing platinum. Such catalysts contain a small amount, eg between about 0.1% and 2% by weight of platinum supported on an acidic carrier. The acidic carrier my consist of compounds of silica and alumina or of a non-acidic carrier, such as aluminum oxide, which has been made acidic by applying a small proportion of an acidic material such as boria, phosphoric acid or a halogen. Thus, porous aluminum oxide containing a small proportion of fluorine and/ or chlorine, e.g. 0.1% to 2% by weight, is a suitable acidic carrier for the platinum.

When using such platinum catalysts the partial hydrogen pressure is suitably between 15 and 50 atmospheres and preferably between 18 and 40 atmospheres for the temperatures between 440 C. and 540 C. and preferably between 470 C. and 540 C.

A preferred method of carrying out the process of this invention is illustrated by means of FIG. 1. Crude oil is passed through line 11 into fractionating zone 12, which may consist of one or a multiplicity of fractionating columns. Among the fractions produced is a C -and-lighter straight-run fraction in line 14, a naphtha fraction in line 15, a light gas oil in line 16 and a heavy gas oil in line 17. The naphtha is desulfurized in desulfurizer 19, which may be hydrogenative or nonhydrogenative, and the desulfurized naphtha passed to catalytic reformer 20. The catalytic reformer is suitably of the platinum reforming type, employing catalyst and conditions mentioned above. The reformate is separated into a gaseous and a liquid fraction in separator 21, the gaseous fraction being taken through line 22 which may contain an H 5 removal unit. The fraction may be desulfurized, partly desulfurized and partly by-passed, or completely by-passed. This fraction, consisting essentially of hydrogen and some C and C hydrocarbons, is returned at least in part to furnish hydrogen to the catalytic reformer via line 23 and may be passed in part to a gas oil catalytic desulfurizer through line 24. The total hydrocarbon fraction recovered from separator 211 is passed to fractionator 26in which naphtha is taken as a bottoms fraction and a gas fraction, e.g. C to C is taken through line 28.

The light gas oil from line 16 passes to catalytic desulfurizer 30 in which it is contacted with hydrogen-rich gas from line 24 at the above-mentioned conditions suitable for catalytic desulfurization in the presence of hydrogen. The reaction product passes to separator 31 in which a light gas fraction, consisting mainly of hydrogen and C and C hydrocarbons, is taken through line 32 containing an H 8 removal unit, the desulfurized hydrogen fraction being sent to line 24 for return to the desulfurizer. The hydrocarbon product from separator 31 is passed to fractionator 35 in which a desulfurized gas oil is taken as bottoms, an intermediate fraction may be taken if desired, and a gas fraction, eg. a C -C cut, is taken overhead through line 36. It will be understood that the separation is indicated more or less schematically and that fractionation zone 35 may consist, for example, of a steam stripper 7 in which the separator liquid product is contacted with steam to strip therefrom all components lighter than gas oil, the overhead being recovered in an additional separator in which a liquid gasoline fraction is separated 8 alkyl-ate, if desired. The fraction in line 68 is fractionated in column 74 to take overhead a fraction lighter than isobutane, e.g. propane present in the system, and to take as bottoms an isobutane concentrate which is returned via from a gaseous C -C -containing fraction. 5 recycle line 75 to alkylation feed line 56. If desired, some The heavy gas oil in line 17 is passed to catalytic crackof the product in line 75 may be Withdrawn as a bleed ing unit 40 which may be of the fluidized or of the riser stream through line 76. type or may be a fixed-bed catalytic cracking unit or a In the above-described process it is essential that the moving bed catalytic cracking unit. The cracked prodisobutane-to-butylene ratio in line 56 ahead of the addiuct is passed through line 411 to fractionation zone 42 tion of recycle isobutane via lines 75 and 64' be at least in which a number of fractions may be separated, includ- 1:1 and preferably in the range from about 1.05:1 to ing a light hydrocarbon fraction such as a C -C conabout 1.15:1. This composition is generally the same as centrate taken through line 44. that in surge drum 55. When the said ratio in surge drum The light hydrocarbon fractions from lines 14, 28, 36 55 falls below the desired value and when the ratio in and 44 are combined in line 50 and passed through line line 54- is not sufliciently high to permit prompt adjustment 51 to a gas separation unit 52, which may be of the conto the desired range in line 56 then at least a substanventional type in which the light gases are compressed tial portion of the C fraction in line 54 is Withdrawn and the compressed gases are contacted countercurrently through line 80, shown as a dashed line, and is passed With an absorber oil in one or more stages. The gas sepatherethrough either to catalytic desulfurizer 30 via line ration unit may further contain an absorber oil debutan- 2O 81 or to catalytic reformer via line 82. Addition to izer, an H S removal unit for the debutanizer overhead the catalytic desulfurizer is ordinarily preferred. The and a butane depropanizer so that the final product leavbutylenes present in the fraction added via lines 81 or ing gas separation zone 52 via line 54 consists essentially 8 2 are completely hydrogenated in units 30 or 20; the only of C hydrocarbons. The products taken from gas resulting saturated C mixture is ultimately returned to separator 52 other than the C, out in line 54, are thus, surge drum 55 via the fractionation and gas separation for example, a C and lighter fraction, a propane-propylsystem. ene fraction and a gasoline fraction. The C cut in line The following examples are presented for illustrative 54 passes to surge vessel 55 which provides feed to purposes only and are not to be considered as limiting the alkylation zone 58 via line 56. Alkylation zone 58 is, for invention. example, a reactor system in which butylenes are con- EXAMPLE I tacted with isobutane in the presence of concentrated an The proportion f il bl l fi hi h i to b 1 hydrous HF Catalyst added Via line P acid drogenated in a given situation is easily determined. The is added through line oi. The 0.; fraction in line 56 is f ll i is a typical calculation; admixed with recycle sobutane from line 64 before being lsopamffim Olefin ratio desired for charge to alkyhr passed to the alkylation zone. tiOnzlL The temperature in the alkylation zone is suitably be- Available feed (moles): tween 30 C. and C. and the pressure suitably about lsopamfiqn A 10 atmospheres. The alkylation zone may be divided into Normal fii several reactors, preferably substantially agitated and Isoolefin B cooled to remove heat of reaction. The reaction mixture 40 Normal olefin C which leaves alkylation Zone 58 is first passed through a m pressure separator 59 wherein it is separated into two 10ml liquid layers, the lower layer consisting mainly of HF; Condition: A 1.1 (8-!- C) this is recycled to the alkylation zone via line 60 which Let the amount of isoolefin to be hydrogenated by X may contain an HP regeneration zone for removal of moles. L impurities. The hydrocarbon phase which, besides hy- Eubsumte the reiatlonshlpz drocarbons, also may contain about 1% HF, passes from Total f f a i fresh. feed to alkylatwn the e ar t r t fractionat r 62-. in which th maor fl O .Qlefins m fre-Sh feed to alkylatlon) s P a o e POT =l.1 (originally available olefinshydrotron of remaining HP is removed in admixture with some genated 1 5 light hydrocarbons and recycled to line 56 via line 64-. The bottoms from column 62 are passed via line 65 to A+X=1,1 (3+ C (1 column 66 in which isobutane and lighter hydrocarbons B are taken overhead in line 68 and heavier product, mainly 1 1 (13+ 0) A alkylate, taken as bottoms in line 69. The alkylate is W defluorinated in 'defluorinator 70, which may be, for ex- T ample, a vessel containing alumina; the defluorinator X product is distilled in column 71 to remove remaining The proportion of total feed to be hydrogenated is butane overhead and take alkylate product as a bottoms B fraction. The alkylate may be re-run to remove heavy The results, expressed in general terms, areas follows:

Available Sent to hydro- Hydrogenated Direct to alkylation Total to genation alkylation Isoparafiins A 5A %A+X A%A A+X X X Normal paraflins EC -0 Isoolefins B X B-X B-X Normal olefins 0 5-0 C%O C C +B+ +B+0 +B+ 1%) +B+0 A+B+o 9 Without hydrogenation, the amount of alkylate produced is With hydrogenation, it is 11 moles EXAMPLE II Sent to hydrogenation Direct to alkylation Total to alkylation Hydrogenated Isoparaflins Normal paraffins Isoolefins Normal olefins.

By hydrogenating 36% of the total feed, the yield of alkylate is increased from 31.8 moles to 41.6 moles.

EXAMPLE 1H 40 Reference is made herein to the flow lllustrated in FIG. 1 of the drawing.

In a petroleum refinery, crude oil is distilled into several fractions including a light fraction (A) (line 14) consisting predominantly of C -C hydrocarbons, and several other fractions referred to hereinafter. Fraction (A) contains 13.6 t./d. (tons per day) isobutane and 59.9 t./d. n-butane.

A heavy naphtha fraction is taken via line 15, desulfurized and catalytically reformed over a supported acidic platinum alumina catalyst. Part of the resulting hydrogen is recycled and part passed to catalytic desulfurizer '30. The hydrogen fraction passing via line 24 to desulfurizer '30 contains 1.8 t./d. isobutane and 2.2 t./d. n-butane as well as other saturated light hydrocarbons. From the fractionation of the catalytic reformate there is recovered a light fraction (B) (line 28) which contains 24.3 t./ d. isobutane and 38.0 t./d. n-butane.

In desulfurizer '30 gas oil is desulfurized by means of cobalt oxide-molybdena-alumina catalyst, using the trickle 0 technique. The pressure is about atmospheres, the temperature about 380 C., the flow velocity about 2.5 v./v./hr. and the gas-to-oil ratio about 125 l. gas/kg. oil. From the product work-up there is recovered a light fraction (C) Which contains the above-mentioned 1.8 t./d. 5 isobutane and 2.2 t./d. n-butane.

Heavy gas oil is cracked in cracking unit 40. In the distillation of its reaction product there is recovered a C -C hydrocarbon fraction (D) (line 44). This fraction contains 75.5 t./d. isobutane, 29.8 t./d. n-butane, 28.8 t./d. isobutylene and 67.2 t./ d. normal butylenes.

The light fractions (A), (B), (C) and (D) are combined in line 50 and passed through gas separation zone 52. The effluent (line 54) contains 115.2 t./d. isobutane, 129.9 t./d. n-butane, 28.8 t./d. isobutylene and 67.2 t./d.

n-butylene. When the system is in steady state, the butane composition in surge drum 55 has the same composition. The molecular ratio of isobutane to total butylenes is thus 1.16, which is in the correct range for satisfactory operation of the alkylation unit. Since even lower ratios are satisfactory, some isobutane may be recovered for other uses from line 76 when the system is in operation.

When fraction (A) is not available however, i.g. because of difliculties in the crude distillation while the reforming, desulfurizing and cracking units are still operating satisfactorily on stored feed stock, then valve 13 in line 14 would be closed and the butane feed Would consist only of fractions (B), (C) and (D). These combined fractions contain 101.6 t./d. isobutane, 70.0 t./ d. n-butane, 28.8 t./d. isobutylene and 67.2 t./d n-butylenes The molecular ratio of isobutane to total butylenes in line 54 is thus 1.02, which is undesirably low.

In order to remedy this, a portion of the fraction in line 54 is passed to line 80 to be hydrogenated, for example, by passing it through line 81 to catalytic desulfurizer .30. For steady-state operation it is suitable to pass, for example, 283.5 t./d. of the C fraction to surge drum 55; this fraction would contain 109.3 t./ d. isobutane, 78.2 t./d. n-butane, 28.8 t./d. isobutylene and 67.2 t./d. normal butylenes. Of the total passing through line 54 one recirculates via line 80, for example to desulfurizer 30, a total of 15.7 t./d. which leaves for alkylation a net of 267.8 t./d. of a C fraction consisting of 103.2 t./d. isobutane, 73.9 t./d. n-butane, 27.2 t./d. isobutylene and 63.5 t./d. normal butylenes The molar ratio of isobutane to total butylenes in the alkylation feed is thus 1.10, which is a value in the desired range.

The passage of the small amount of C fraction through the catalytic desulfurizer has no undesirable effect on the desulfurization of the gas oil. There is, of course, a slight increase in hydrogen consumption (circa 7.2 t./d. instead of 7.0 t./d.). This effect is so small as to have no substantial influence on the operation of the process.

The present invention is further illustrated by the following example in which reference is made to FIGS. 2 and 3. FIG. 2 is a schematic representation of an alkylation and fractionation system and FIG. 3 a similar schematic representation of an alkylation and fractionation system including hydrogenation in accordance with the process of this invention.

EXAMPLE IV The present example illustrates how a substantially greater total production of alkylate can be achieved in a refinery in which the total butylenes available have a relatively high ratio of isobutylene to normal butylenes and in which isobutane is not present in the amount re quired for alkylation. Such a situation is found, for example, in a refinery in which all of the cracking is carried out in a catalytic cracking unit which produces a high ratio of isoto normal butylenes.

Referring to FIG. 2, the alkylation unit is a conventional one, e.g. a sulfuric acid alkylation unit. The total feed fraction (1) passing to the alkylation unit has the following composition: isobutane, 210 b./d. (barrels per day); normal butane, 244 b./d., isobutylene, 226 b./d.; and normal butylene, 320 b./d., amounting to a total of 1000 b./d. This is passed to the alkylation unit which contains the usual reactors and associated equipment. The total effluent passes to a deisobutanizer in which isobutane is removed overhead for recycle, the bottoms passing to a debutanizer in which n-butane is removed as overhead and total alkyl'ate is taken as the bottoms stream.

The compositions of the various streams are shown in Table 1, below. For convenience of presentation, volume ratios. are employed in the calculation, although this causes a slight loss of precision, which is not of significant magnitude.


i-O4H 210 70 140 105 II-CiHro. 244 81 163 119 1-C4Hs 226 75 151 110 n-O4Ha 320 107 213 157 Total C4Hs 546 182 364 267 Total 04 1, 000 333 667 608 491 Alkylate. 312

The ratio of isobutane to total butylenes in the fresh feed is only about 0.385:1. There is, thus, a very substantial excess of butylenes since isobutane is limiting and since a ratio of 1.15:1 is desired in this example. The amount of butylenes that can be alkylated is only 182 b./d. It is, therefore, necessary to split the total feed and pass only fraction (II), which contains 182 b./d. of butylenes, to the alkylation zone while fraction (III) is bypassed and sent to the alkylation fractionation system so that the isobutane content thereof can be recovered and recycled back to the alkylation zone while the normal butane and butylenes content passes with the alkylate to the alkylate debutanizer in which fraction (IV) is taken as overhead. This contains 364 b./d. of unconverted butylenes. The amount of alkylate produced in this system is 312 b./d.

Instead of passing fraction (III) to the alkylate fractionating system it is, of course, also possible to split it separately because it is only necessary to recover its isobutane content for alkylation.

The improvement obtainable by treating the same feed fraction (1) in accordance with the present invention is illustrated by reference to FIG. 3.

The proportion of fraction (I) which must be hydrogenated in order to provide a complete balance between isobutane and butylenes is readily calculated.

If a portion of the total fraction (I) is to be hydrogenated to produce suflicient isobutane to alkylate the remaining butylenes, the required volume of isobutylene to be converted is 110 b./d., calculated as illustrated in Examples I and II. Thus fraction (VI) passing to the dehydrogenator must contain 110 b./d. of isobutylene. The hydrogenation may take place in the manner illustrated previously, e.g. by combining the fraction with the feed to a reforming or hydro-genative desulfurization unit or by a separate hydrogenation step. The hydrogenated C fraction returned to feed to alkylation has composition (VII); the total fresh feed to alkylation thus has composition (VIII), the debutanizer overhead has composition (IX) and the total alkylate production (X) is 460 b./d. Thus, by the mere expedient of hydrogenating a portion of the total fresh feed, alkylate production has been increased from 312 to 460 b./d., or by approximately 50%.

Process variations within the scope of this invention will occur to those skilled in the art. For example, where separating means are available the isoolefin may be concentrated relative to normal olefins prior to the hydrogenation step.

I claim as my invention:

1. A process for increasing the yield of alkylate o tained from a feed fraction containing isoolefin and normal olefin and containing less than one mole of alkylatable isoparafiin per mole of alkylating olefin which comprises separating said feed fraction into a first portion and a second portion, hydrogenating in the presence of free hydrogen the first portion to convert the isoolefin and normal olefin therein to the corresponding isoparafiin and normal parafiin, recombining the hydrogenated first portion with the unhydrogenated second portion and alkylating said recombined portions by contact with a mineral acid alkylation catalyst at alkylating conditions,

the relative proportions of said first portion and said second portion being adjusted so that the recombined first and second portions have a molar ratio of isoparaffin to olefin of at least 1.05:1.

2. A process according to claim 1 in which the mineral acid is concentrated hydrofluoric acid.

3. A process according to claim 1 in which the mineral acid is concentrated sulfuric acid.

4. A process according to claim 1 in which said first portion of the feed fraction is hydrogenated in the presence of free hydrogen by being admixed with the naphtha feed to a reforming process and at least the isoparafiinic components of the reformed product in the molecular weight range of the feed fraction are recovered from the reformed product and passed to the alkylation with said second portion of the feed fraction.

5. A process according to claim 1 in which said first portion of the feed fraction is hydrogenated in the presence of free hydrogen by being admixed with the gas oil feed to a desulfurization process and at least the isoparaifinic components of the product in the molecular weight range of the olefinic fraction are recovered from the desulfurization product and passed to the alkylation with said second portion of the feed fraction.

6. A process for the production of high octane number gasoline components which comprises catalytically cracking a hydrocarbon oil to produce a cracked product including a light hydrocarbon fraction having an isoto normal butylene ratio of at least about 0.8:1 and having an isobutane to olefin ratio of less than 1:1, separating said fraction into a first portion and a second portion, hydrogenating in the presence of free hydrogen said first portion to convert the isoand normal butylene therein to isobutane and normal butane, respectively, recombining the hydrogenated first portion with the unhydrogenated second portion and alkylating said recombined portions by contact with a mineral acid alkylation catalyst at alkylating conditions, the relative proportions of said first portion and said second portion being adjusted so that the recombined first and second portions have a molar ratio of isobutane to olefin of at least 1.05:1.

7. A process for the production of high octane number gasoline components which comprises catalytically cracking a hydrocarbon oil to produce a cracked product including a light hydrocarbon fraction having an isoto normal butylene ratio of at least about 0.8:1 and having an isobutane to olefin ratio of less than 1:1, separating said fraction into a first portion and a second portion, hydrogenating in the presence of free hydrogen said first portion of said fraction to obtain isobutane and normal butane and recombining the hydrogenated first portion with said tmhydrogenated second portion so that the recombined first and second portions have a molar ratio of isobutane to olefin of at least 1.05:1, recombining the portions and alkylating in an alkylation zone the recombined fraction by contact with a mineral acid alkylation catalyst at 'alkylation conditions whereby alkylate is formed, and recovering normal butane from said alkylation zone.

8. A process in accordance with claim 7 and further characterized by isomerizing in an isomerization zone in the presence of an isomerization catalyst under isomerization conditions the normal butane recovered from said alkylation zone, whereby isobutane is formed, separating said isobutane from said iso-merization Zone and recycling said isobutane into said alkylation zone.

References Cited in the file of this patent UNITED STATES PATENTS 2,172,146 Ruthruff Sept. 5, 1939 14 Roetheli Oct. 17, 1944 Evering et a1. May 28, 1946 Munday Dec. 17, 1946 Porter Feb. 11, 1947 Lee Mar. 11, 1947 Skelton Apr. 5, 1949 Voge et a1. Jan. 24, 1950 Krebs et a1. Feb. 23, 1954

Patent Citations
Cited PatentFiling datePublication dateApplicantTitle
US2172146 *Jul 20, 1935Sep 5, 1939Standard Oil CoManufacture of iso-butane
US2360622 *Apr 30, 1943Oct 17, 1944Standard Oil Dev CoMethod of producing aviation gasoline
US2400922 *Aug 13, 1942May 28, 1946Standard Oil CoHydrocarbon conversion system
US2412645 *Oct 14, 1943Dec 17, 1946Standard Oil Dev CoTreating hydrocarbon fluids
US2415530 *Mar 8, 1943Feb 11, 1947Pure Oil CoIsobutane production
US2417308 *Apr 12, 1943Mar 11, 1947Union Oil CoDesulphurization and hydroforming
US2466334 *Aug 10, 1944Apr 5, 1949Texas CoMethod of producing synthetic fuel
US2495648 *Aug 24, 1946Jan 24, 1950Shell DevHydrocarbon treating process
US2670322 *May 1, 1951Feb 23, 1954Standard Oil Dev CoNaphtha reforming process
Referenced by
Citing PatentFiling datePublication dateApplicantTitle
US3138645 *Aug 21, 1961Jun 23, 1964Pure Oil CoAlkylation process and stabilization of product
US3170002 *Jun 19, 1963Feb 16, 1965Exxon Research Engineering CoReduction of acid consumption in alkylation
US3250822 *Jan 17, 1963May 10, 1966Shell Oil CoHydration-alkylation process
US3714022 *Sep 22, 1970Jan 30, 1973Universal Oil Prod CoHigh octane gasoline production
US7939953 *Apr 16, 2008May 10, 2011Schlumberger Technology CorporationMicro scale fischer-tropsch and oxygenate synthesis process startup unit
US8293805May 29, 2008Oct 23, 2012Schlumberger Technology CorporationTracking feedstock production with micro scale gas-to-liquid units
U.S. Classification208/67, 585/719, 585/833, 585/253, 585/850, 585/730, 585/841, 585/256, 585/723, 585/717, 585/250, 585/251
International ClassificationC07C9/00, C07C9/16
Cooperative ClassificationC07C9/16
European ClassificationC07C9/16