US3097155A - Process for the conversion of paraffin hydrocarbons with isobutane utilizing hydrogen fluoride as a catalyst - Google Patents

Description

July 9, 1963 PROCESS FOR THE cdNv B S. FRIEDMAN ETAL ERSION OF PARAFFIN HYDROCARBONS WITH 25 26 STORAGE TANK 32 24 2? HEATER 3O {29 EXTRACTOR REACTOR 7 STRiPPER --l HEATER J i a DISTILLATION 35 C OR COLUMN 290 as 9 STORAGE 37 TANK HEATER DiSTiLLATION 38 lo COLUMN 4| l H REACTOR 41a 4 I6 l2 ,40 1 A H '3 f b REACTOR FL 5 SEPARATOR DRUM 44 48 '8 T l5 SEPARATOR :9 2O

DISTILLATION 43 46 COLUMN SEPARATOR ez DlSTlLLATiON COLUMN 23 l FRACTIONATOR INVENTOR IMITROVIS ATTORNEY;

United States Patent 3,097,155 PROCESS FOR THE CONVERSION OF PARAFFIN HYDROCARBONS WITH ISOBUTANE UTILIZING HYDROGEN FLUORIDE AS A CATALYST Bernard S. Friedman, Chicago, and Robert P. Zmitinvrs, Park Forest, 111., assignors, by mesne assignments m Sinclair Research, Inc., New York, N .Y., a corporation of Delaware Filed Apr. 3, 1959, Ser. No. 803,873 8 Claims. (Cl. 20866) The present invention relates to a method for the conversion of hydrocarbon fractions boiling in the motor fuel range to obtain products of higher octane value. More specifically, our invention is directed to the treatment of parafiinic materials boiling in the motor or spark-ignition engine fuel range in the presence of a hydrogen fluoride catalyst while charging to the reaction zone a substantial amount of isobutane. In one particular aspect the invention is concerned with the converstion of straight run petroleum fractions boiling in the motor fuel range to obtain good yields of higher octane gasoline by the use of a plurality of reaction stages employing catalysts differing in characteristics. In this particular system one of the stages includes the present hydrogen fluoride catalytic operation.

In recent years the automobile manufacturers have steadily increased the compression ratios of their sparkignition engines as a means of obtaining more power and greater efficiency. As the compression ratios of the engines increase, the hydrocarbon fuel employed must be of higher octane value to provide efficient knock-free operation notwithstanding that fuel octane can be increased through the addition of tetraethyl lead; and other undesirable aspects of engine operation, for instance preignition, can be overcome by the use of other additive components. Thus the problem remains for petroleum refiners to produce higher octane base hydrocarbon fuels under economically feasible conditions.

These refiners now have installed a substantial number of units for reforming straight run petroleum fractions in the presence of free hydrogen and over a platinum metalalumina catalyst to obtain relatively high octane products. Primarily these products, frequently called reformates, are blended with other gasoline components such as thermal and catalytically cracked gasolines, alkylates, etc., and additives such as tetraethyl lead in obtaining presentday motor fuels. The reforming operation has a number of disadvantages First, as the octane requirements of the blended engine fuels rise, the octane quality of the reformate must also increase if the blends be otherwise unaltered. This increase results in a substantial reduction in yield particularly when obtaining reformates having octanes (RON neat) of the order of 90 to 95 or above. As the severity of the operation is increased, the platinum metal-containing catalyst becomes fouled more often with carbonaceous deposits which requires more frequent regenerations and/or replacements. As the platinum metalalumina catalysts are relatively expensive, either replacement or withdrawal from use during regeneration materially increases the cost of providing a given volume of reformate. These and other factors affecting the yieldoctane number-cost relationship make it desirable for the refiner to consider various ways in which high octane hydrocarbon fuel components can be obtained by employing processing methods other than or in conjunction with the platinum metal-alumina catalyst reforming operation.

One method now under consideration by petroleum refiners for obtaining stocks of higher octane value involves the isomerization of parafiinic hydrocarbons boiling in the motor fuel range. In general, as the side chain branching of normal parafiins and of slightly branched isoparaffins increases, their octane ratings rise. A number of 3,097,155 Patented July 9, 1963 "ice catalysts are known as being useful in this type of operation and such catalysts include hydrogen fluoride (2,691, 688), hydrogen fluoride-boron trifluoride (2,405,996 and 2,583,740) and platinum-alumina. In the present invention we have devised a method which provides substantial improvements in the processing of normal paraffins, isoparaifins or normal paraffin-isoparaflin mixtures in obtaining components of higher octane value. These motor fuel range feedstoclts are essentially olefin-free, i.e. contain less than about 5 weight percent olefin, preferably less than about 1%, and boil primarily in the approximate C, to 425 F. range, frequently in the range of about 190 to 300 F. Particularly, by operating under given processing conditions and in an essentially olefin-free system containing less than about 5 weight percent, preferably less than about 1%, total olefins based on the 0 feedstock, we have obtained material advantages in terms of an increase in the yield-octane relationship of the gasoline hydrocarbons with a substantial portion being in the C to C range. Not only does the present system give materially advantageous results in the yield-octane relationship but also substantial economic savings are effected due to the exclusion of catalytic materials such as boron from the system and the fact that provision of an olefin promoter is not necessary. Although our system effects isomerization reactions, there is also considerable evidence at this time that other reactions such as disproportionation, polymerization and alkylation take place. By employing our improved process the operator can supply isoparaffin products for various uses and particularly as the yieldoetane relationship of the gasoline boiling range product is exceptional our method affords an advantageous means for obtaining higher octane gasoline blending components which permit a petroleum refiner to elevate the octane rating of his gasoline pool and with a corresponding decrease in cost due to the avoidance of the use of boroncontaining catalyst components and olefinic promoter.

The advantageous yield-octane relationships afiorded by our method are due among other things to the charging to the hydrogen fluoride reaction system of substantial amounts of isobutane. Thus we provide at least about 50% by weight of isobutane based upon the charge of 0 liquid boiling in the motor fuel range. The upper limit on the amount of isobutane charged is primarily an economic question involving factors such as the cost of distilling C s from the liquid product, the value of the liquid product, and the cost of increased consumption of butanes. However, in general there does not seem to be any advantage in employing more than about 600 to 1000 weight percent of the isobutane and preferably we charge about to 460 weight percent.

Also, it is important in our method that the hydrogen fluoride reaction system be operated so that preferably there is no net make of butanes although in some instances due to economical considenations some butane make can be tolerated. However, in those instances, when a not make of butane is tolerated there may be a decrease in ultimate yields of the higher octane product. In onder to obtain the best results from the present system, we prefer that there be at least a net consumption of butanes and advantageously there is consumed in the reaction at least about 5 weight percent of butanes based upon the 0 feed boiling in the motor fuel range. The isobutane charged to our system can be relatively pure or mixed with other materials such as the various refinery hydrocarbons. Frequently, as found in the refinery, the isobutane is in admixture lat least normal butane, and we can employ such mixtures. Of course, the isobutane can be recycled from the reaction system efliuent and usually at least a portion of any normal butane fed to the system is converted to isob-utane which can be recycled and considered as part of the isobutane requirement charged to the rcaction zone. The amount of isobutane in the light refinery streams could be increased as by contact in an isomen'zation reaction stage with a catalyst before passing to our hydrogen fluoride conversion zone. The effluent from this preliminary reaction is combined with our parafiinrich feedstock and charged to the principal hydrogen fluoride reaction system. The feed to the hydrogen fluoride reaction system can also include materials such as propane, hydrogen or other light diluents. However, it is desirable to keep the concentrations of these materials low since they are not as beneficial as isobutane and heat may have to be supplied to any subsequent sys tem to separate them from the reaction product. Also, these extraneous materials occupy space, in the reaction system which could necessitate an increase in equipment size, and they may in effect retard the desired reactions.

It may be that there are a number of reactions which are effected in our system. For instance, while the paraflinic C feed is undergoing isomerization, there may be a disproportionation of C to C and C hydrocarbons. Moreover, there is evidence that relatively high octane gasoline components are produced possibly by C polymerization and by alkylation. These and other reactions might be taking place in our system but in any event, we believe that the presence of the relatively large amount of isobutane and the C paraflinic materials in the overall hydrocarbon feed is necessary in obtaining the desired result. In general we prefer to obtain at least about 50% conversion of our feedstock.

The C7+ hydrocarbon feeds to our hydrogen fluoride catalytic process contain primarily normal paraflins and/or isoparaflins boiling in the motor fuel range. Such materials are found in various petroleum refinery streams and can be separated in more or less pure form or obtained in admixture with similar boiling materials such as olefins, naphthenes and aromatics. Although we can feed pure n-paraffins or isoparaffins, we advantageously employ mixtures of these materials containing at least about 10% of each component based on the 0 feed. Our preferred hydrocarbon C motor fuel range feedstocks contain about 15 to 35 weight percent of normal paraffin and about 65 to 85% isoparafiins based on their mixture in the reaction zone. Thus the feed can be n-heptane or other C gasoline boiling range normal parafiins either alone, mixed with each other or with their isoparaffins. Also the C7-I- hydrocarbon liquid feed can be any suitable narrow or broad range paraffinic material boiling in the motor fuel range and can contain, if desired, minor amounts of similar boiling, heavier or lighter components such as olefins, aromatics or naphthenes. However, the reaction system does not contain more than about of olefins; generally not more than about 15% of arcmatics, preferably more than about 1% aromatics, and not more than about 15 preferably not more than naphthenes. Usually, the nand isoparaffins will constitute at least about 85 weight percent of our 0 motor fuel range feedstock and preferably at least about 90%. The isoparaflins we have referred to are essentially or predominantly composed of branched chain structures which can be isomerized or converted in the reaction to structures having at least one additional branched chain.

A feedstock satisfactory for our hydrogen fluoride reaction system is a paraflinic concentrate derived from the 0 Liquid product contained in the efllucnt of reforming systems employing petroleum straight run naphtha feedstocks and platinum-alumina catalysts. As an example most if not all of the aromatics of the reformate can be separated as by adsorption, extractive distillation, or any other procedure desired. The paraffinic materials resulting are useful in our process particularly when in a feed containing about to weight percent of normal parafilns and these stocks will usually contain predominantly isoparafiins, and only small amounts of olefins and aromatics. To obtain the paraffin-rich feed from the reformate, the aromatics can be adsorbed on silica gel; or separated by solvent extraction through the use of a solvent selective for aromatics, e.g. phenol, or by any other desirable procedure. A particularly useful method for accomplishing this separation is employed commercially and includes the use of a glycol-water extraction medium. As commercially licensed, one such system is known as Udexing. By regulation of conditions such as the glycol to water ratio, the extraction and solvent stripping temperatures, and the character of the glycol, a Udex raflinate varying in paraflinicity can be obtained. The manner of controlling these factors is known in the art and it suffices to say that preferred glycol materials are the diglycols such as diethylene and dipropylene glyco-ls and their mixtures.

In the present hydrogen fluoride reaction system the conditions of treatment can vary widely depending upon fuel composition, octane number desired for the final product, economical yields, etc. We maintain the catalyst essentially in the liquid phase under the reaction conditions. Thus the pressure to be maintained must be sufficient to provide essentially liquid phase reaction conditions as determined by the vapor pressure of the hydrogen fluoride, the reactants and other materials such as hydrogen, and the reaction products present. Generally, the pressure will be at least about p.s.i.g. and preferably about 200 to 400 p.s.i.g. There seems to be little if any advantage in the pressure being above about 1500 p.s.i.g. and at higher pressures propane formation may be excessive. The reaction temperature and reaction contact time are interdependent factors with a lesser time being required to provide a given result as the temperature increases. The reaction temperature will usually be in the range of about 100 to 400 F., preferably about to 300 F., with the time required ranging from about 1 minute to 5 hours, preferably about 20 minutes to 2 hours. Times longer than about 5 hours can be employed; however, no particular advantage would be derived thereby which overcomes the obvious economic disadvantages. Although the contact time and temperature employed can be selected as desired and can even be dependent upon factors such as the type of reaction system employed, we believe the following temperature-time relationships provide the best results but We do not intend to be limited by them.

TABLE I Temperature, F.: Contact time, minutes 150 to 200 30 to 180 200 to 250 5 to 60 250 to 300 l to 20 In general, We prefer to select conditions which avoid production of substantial amounts of propane and long contact times at high temperatures.

In the reaction zone enough hydrogen fluoride is added so that a catalyst layer separate from the hydrocarbon layer can be obtained. Usually, this requires at least about 0.5 volume of hydrogen fluoride per volume of total paraffinic hydrocarbon feed. This ratio can be as high as 5 to 1 or more but preferably is not greater than about 2 to 1 as larger amounts of hydrogen fluoride require excessive handling facilities. The hydrogen fluoride can be added to the reaction mixture in any manner found most convenient. It is desirable that the hydrogen fluoride reaction system be conducted essentially in the absence of water although small amounts of water will not prove deleterious to the present system.

The hydrocarbon product and the hydrogen fluoride catalyst layers can be separated in any manner desired. When agitation of the reaction mixture is stopped it will separate into two phases in the reactor or in any other vessel into which it is transferred as in a continuous, semicontinuous or batch operation. These phases can be separated by simple decantation. The reaction mixture could be allowed to separate into a lower layer of catalyst containing an unsaturated oil including aromatics which can be recycled to the reaction system in whole or in part.

Usually, we keep the aromatic content of the C feed at less than about 15 weight percent, preferably less than about 1 percent. In designating amounts of aromatics, we refer to the values obtained by the Fluorescence Indicator Adsorption method commonly known as the F.I.A. method involving chromatography on silica gel. The upper hydrocarbon layer formed in our system could be freed from catalyst by distillation and/ or washing with water or passed through a column of basic ion exchange resin or other solid adsorbent such as charcoal, potassium sulfate, sodium sulfate, etc. The unsaturated oil (catalyst oil) appearing in the catalyst layer could be separated as by distillation of the catalyst, and the portion of the catalyst oil boiling in the gasoline range might then be combined with the hydrocarbons of the upper layer to provide a higher octane product. Small traces of fluoride remaining in the hydrocarbon material can be removed as by passage over aluminum or alumina at 200 to 500 F. Various drying procedures could be employed to separate water from the hydrocarbon materials and such materials could be stabilized, for instance by the removal of C s and lighter constituents.

In one particular aspect of the present invention we have devised a highly attractive method for obtaining higher octane fuels for spark-ignition engines which involves the use of separate reaction zones in which are employed catalysts of different properties. In this method a straight run gasoline or naphtha boiling range hydrocarbon is contacted with a platinum metal-alumina catalyst in the presence of free hydrogen under conditions which provide a substantial increase in the octane number of the petroleum hydrocarbon material. A paraffin-rich fraction of the resulting reformate and isobutane is then contacted with the hydrogen fluoride catalyst according to the above described procedure to give a product boiling in the motor fuel range which is of substantially increased octane quality.

The hydrocarbon feedstocks charged to the reaction system containing the platinum-metal-alumina catalyst are primarily the straight run petroleum fractions boiling in the gasoline and naphtha ranges, for instance in the range from about 175 to 450 F., but somewhat higher or lower boiling constituents can be included if desired. Preferably, the feedstock boils in the range of about 200 to 400 F. Although the hydrocarbons passing to the platinum metal-alumina catalyst reaction system are composed of predominantly straight run naphtha material, minor amounts of additional components can be included such as olefins, thermal and catalytically cracked stocks, recycled reformate and fractions of these cracked and reformed materials. The reaction conditions observed or maintained in the platinum-metal-alumina catalyst system include those suggested for present commercial reforming operations such as temperatures from about 750 to 1000 F., preferably about 825 to 975 F., and pressures from about 50 to 1000 p.s.i.g., preferably about 150 to 500 p.s.i.g. The free hydrogen supplied to this reaction system usually is in the form of hydrogen-rich recycle gases and generally provides about 2 to 20 moles of hydrogen per mole of hydrocarbon feed; preferably this ratio is about 4 to :1. The space velocity usually lies in the range of about 0.5 to 10 WHSV (weight of feed per weight of catalyst per hours) preferably about 2 to 5 WHSV.

The platinum metal-alumina catalysts employed in the method of this invention include a number of compositions. Generally, the platinum metal is a minor amount of the catalyst, e.g. about 0.1 to 1.5 weight percent of the final composition. Platinum is the most commonly employed metal present in these reforming catalysts although other useful platinum metals include rhodium, palladium, iridium which, along with platinum, are the face centered cubic crystallite types of the platinum family as distinct from the hexagonal types ruthenium and osmium which appear to be of lesser value.

These catalysts can be made by a number of procedures but a particularly effective catalyst is one in which the alumina is obtained through calcination of an alumina hydrate containing at least about 65 weight percent of trihydrate and about 5 to 35 weight percent of alumina monohydrate and/or amorphous alumina forms, and if desired having a surface area of about 350 to 550 square meters per gram (BET method) when in the virgin state. The minor amount of platinum metal in the catalyst is usually present in finely divided form and is not detectable by X-ray diliraction techniques. Also, these catalysts are advantageously prepared to afford about 0.10 to 0.5, preferably about 0.15 to 0.3 cc./gram of their pore volume in pores of about to 1000 Angstrom units in size. U.S. Patents Nos. 2,838,444 and 2,838,445 describe the preparation of such catalysts. If desired, the catalyst can contain minor amounts of additional materials, for instance promoting components particularly those acidic in nature, such as silica and fluoride. Such promoting components are usually less than 10 weight percent of the final calcined catalyst.

The platinum metal-alumina catalyst can be employed in any type of reaction system desired, for instance moving or fluidized bed, regenerative or non-regenerative, etc., but advantageously the catalyst is disposed as a fixed bed. In the latter type of operation the size of commercial units is such that essentially adiabatic reaction systems must be employed and in view of this and the endothermic nature of the reforming operation the catalyst is placed in fixed beds in a plurality of reactors, each of which is preceded by means for heating its charge. In fixed bed operations, the catalyst is in macrosize form, that is, particles generally at least about in length and diameter and preferably not exceeding about in diameter. Particularly when such particles are provided by extrusion, their length may be up to about 1" or more. If the platinum metal-alumina catalyst reforming system be of the regenerative type it can be arranged so that the catalyst of all of the reactors can be regenerated simultaneously or individually. Other variations in the platinum metal catalyst reaction system can be made according to the desires of the operator.

In this aspect of the present invention, the essential feed to our hydrogen fluoride catalyst system is a paraffinrich fraction of the liquid reformate from the platinum metal catalyst operation, and as noted above, it advantageously includes about 15 to 35 weight percent of normal parafiin constituents and usually less than about 15 weight percent of aromatics, preferably less than about 1 percent. This feedstock boils primarily in the motor fuel range, for instance, from C up to about 425 F., although heavier constituents can be included. The parafiin-rich feed can be obtained as a UdeX raflinate and the feed contains only a small amount of olefins. The character of this raflinate can be controlled by the boiling range of the reformate feed to the extractive distillation. As an example, if the reformate feed is of narrow boiling range, the rafiinate will also be of close boiling range.

In the drawing we have illustrated a simplified flow sheet of one operation conducted in accordance with our method.

In this system straight run naphtha is charged by way of line 1 to heater 2 and then through line 3 to the top of an initial reactor 4 which contains a fixed bed of platinum-alumina catalyst. The effluent from reactor 4 is passed by way of line 5 to heater 6 and then through line 7 to the top of a second reactor 8 containing a fixed bed of platinum-alumina catalyst. The platinum-alumina catalyst reaction section is of the adiabatic type and more than two reactors can be provided if desired and in fact usually at least three catalyst beds in separate reactors will be employed with each reactor having associated therewith a feed preheater. A third heater 9 and third reactor 10 containing a fixed bed of platinum-alumina catalyst are shown in the drawing.

The reformate from the bottom of reactor 10 is passed by way of line 11 to flash drum 12 which separates C s and lighter materials which are passed through line 13 to separator 14. The separator provides for removal of C to C hydrocarbon constituents through line 15 and hydrogen and methane are recycled by way of line 16 to line 1. Excess hydrogen and methane can be removed from line 16 by way of line 17.

The liquid reformate from flash drum 12 is passed through line 18 to an intermediate portion of distillation column 19 and a light gasoline is taken overhead through line 20. The bottoms fraction from column 19 is carried by line 21 to an intermediate portion of distillation column 22 and heavy gasoline is removed as bottoms from this column through line 23. The overhead from column 22 is passed by way of line 24 to storage tank 25. Liquid hydrocarbon is withdrawn from the storage tank through line 26 and passed to an intermediate portion of extractor 2 7. Entering near the top of extractor 27 through line 28 is a glycol-water extractive medium. The ralfinate produced in the extraction operation is taken overhead by line 29 and transported to storage tank 29a. The extract passes by way of line 30 to an intermediate portion of stripper 31. The stripped extractive medium then returns to extractor 27 through line 28. The overhead from stripper 31 is returned by line 32 to the lower portion of extractor 27.

A side stream from stripper 3-1 is charged to distillation column 34 and a toluene-containing overhead is removed by line 35. The bottoms from column 34 pass by way of line 36 to an intermediate portion of distillation column 37 from which xylenes are removed as overhead by line 38. The bottoms fraction from column 37 contains polymers and is removed by way of line 39.

The raflinate from storage tank 29a is charged through line 41 to reactor 40 after the addition of isobutane by way of line 41a. The hydrogen fluoride catalyst enters reactor 40 through line 42. The reaction efiluent is carried to separator 43 where a hydrocarbon phase and a catalyst phase are formed. The catalyst phase can be recycled to the reactor through lines 44 and 42 while the hydrocarbon phase is passed to the intermediate portion of fractionator 45. C minus overhead from the fractionator goes to separator 46. In this separator, the C hydrocarbons are obtained and then recycled by way of lines 47 and 41 to reactor 40. C and lighter materials are removed by line 48 from separator 46. A gasoline fraction is taken as a side stream from fractionator 45 by way of line 50 while heavier hydrocarbons are with drawn from the fractionator in bottoms line 49- Although the drawing provides an illustration of a typical process we can employ, it is not to be considered limiting; for instance, our paraifinic feeds can be charged to the hydrogen fluoride reaction system in admixture With relatively close out hydrocarbons such as n-pentane and n-hexane. Such mixtures would contain at least about 10% of our (3 motor fuel range hydrocarbon. in another system a catalytic reformate might be flashed to remove C and lighter hydrocarbons and a C to C fraction separated by distillation. The resulting C reformate can be treated to obtain a paraffinic-rich fraction which is then charged to the hydrogen fluoride reaction system. Also, the previously separated C to C portion of the reformate can be passed to this reaction system. The isobutane is provided by recycle from the reaction zone and in addition extraneous normal and isobutanes can be added to the reaction zone as desired. The motor fuel boiling range products would then comprise essentially the gasoline obtained from the paraffinrich portion of the reformate and isopent ane produced in the hydrogen fluoride system due to the charging of n-pentane in the C to C fraction of the reformate.

Example I A straight run naphtha is obtained by distillation from 8 crude oil, and the naphtha typically has an ASTM distillation boiling range of about 209 to 381 F., a RON (neat) of about 47.2, and a gravity API 60 F. of about 56.7. This naphtha is fed to a reforming unit containing three essentially adiabatic reactors each having a fixed bed of a platinum-alumina reforming catalyst. This system is equipped with means for heating the charge to each reactor and the heaters and reactors are arranged for serial flow. The catalyst employed is a platinumalumina reforming catalyst containing about 0.6 weight percent platinum, and manufactured in accordance with U.S. Patent No. 2,838,444 listed above. The inlet temperatures of the feed to each of the three catalyst beds are 940 F., while the pressure is about 500 p.s.i.g. Free hydrogen is supplied to the feed passing to heater before the first reactor and the hydrogen is obtained by recycle from the third reactor eflluent stream. The molar ratio of hydrogen-rich recycle gas (72.7% H to hydrocarbon feed is approximately 5.5 to 1, while the overall space velocity is about 2.34 WHSV. The eflluent from the last reactor is conveyed to a flash drum operating at 500 p.s.i.g. and is then treated or depropanized to remove C and lighter hydrocarbons by distillation. Inspection on the resulting reformate is as follows:

Gravity, API, 60 F. 53. ASTM distillation, F.:

IBP 112.

EP 397. Octane number (RON) 84.9 (neat); 95.3 (3 cc.

TEL added/gal).

Percent aromatics 51.2 (by F.l.A.). Percent olefins 0 (by F.I.A.).

To facilitate an understanding of this example the extractive distillation operation will be described with reference to the drawing. Thus, 6.71 parts by volume of the 112 to 397 F. boiling range reformate are passed at a temperature of 222 F. to an intermediate portion of distillation column 19. The column top temperature is 215 F. and the column bottom temperature is 336 F. In column 19, 2.54 parts by volume of light gasoline are separated as overhead and this gasoline or light naphtha has a gravity APl 60 F. of 71.1 and a boiling range of about 108 to 216 F. 3.17 parts by volume of the bottoms from column 19 are charged at 316 F. to an intermediate portion of distillation column 22. This column has a top temperature of 279 F. and a bottom temperature of 348 F. The overhead from column 22 is 2.24 parts by volume and the bottoms fraction is 0.93 part by volume of a heavier gasoline fraction. The overhead from column 22 has a gravity API 60 F. of about 46.2 and a boiling range of about 250 to 284 F.

The feed to the intermediate portion of extractor 27 is 1.622 parts by volume of the column 22 overhead. The tower top and bottom temperatures of the extractor are 280 F. and there results 0.838 part by volume of iaflinate overhead from the extractor. The bottoms from the extractor is passed to the intermediate portion of stripper 31 which has a top temperature of 231 F. and a bottom temperature of 297 F. 8.16 parts by volume of extractive medium are separated as bottoms from stripper 31 and passed to the top of extractor 27. This extractive medium contains about 17% by volume of dipropylene glycol, 75.5% by volume of diethylene glycol and 7.5% by volume of diethylene glycol and 7.5% by volume of water.

0.784 part by volume of a side stream from stripper 31 are charged at 292 F. to column 34. The top temperature of this column is 232 F. and the bottom temperature is 299 F. The overhead is 0.186 part by volume of a fraction consisting essentially of toluene. 0.598 part by volume are withdrawn as bottoms from column 34 and passed at 291 F. to column 37 which has a top temperature of 285 F. and a bottom temperature of 305 F. The overhead from column 37 is 0.597 part by volume of a fraction consisting essentially of xylenes 10 moved from the Dry Ice bath and allowed to warm up somewhat. The more dense catalyst layer is slowly discharged into a polyethylene receiver containing about 600 cc. of water and thence through a bubble counter,

and the bottoms is 0.001 part by volume of polymer. safety trap, drying tower, Dry Ice cooled traps, water bubbler and wet test meter. The receiver contaming the Example II o I catalyst layer 1s slowly warmed to about 130 F. and A Magnedzlsh autoclave 1 W so mold. i the residue remaining after this evaporation is extracted and a Hastenoy C pot LS evacuated i three times with pentane and the pentane evaporated to vacuum pump f about mom lifimlseramre' .Cooied hquld obtain the catalyst oil. The hydrocarbon layer from hydrogen fluoride (403 i ls dlschargitad mm the auto the pressure vessel is discharged into a glass bottle cooled clave. Isobutane (116 g.) is pressured into the autoclave to equipped with an exit tube leading to a Dry from a pressure cylinder. The autoclave and contents Ice cooled trap a are heated to a temperature of about 250 and under A Small Sample was submitted for analysis by vapor a pressure of about 525 p.s.1.g., wlth occas1onal st1rr1ng. phase chromatography and one for distillation through a A mlxture of grams of Tammie from eftractor low temperature Podbielniak column, to determine com- 27 plus 53.5 grams of lsobutane is then charged mm the ponents up to and including n pentane The main portion autoclave. The contents of the autoclave were stirred for of tha hydmcarbcm fraction is distilled through a high about 30 mmutes- The Talfinate fed from extractor 27 temperature Podbielniak vacuum jacketed distillation colanalyzed as follows: umn to debutanize and depentanize the material, and sub- Gravity, API, 60 F 67.4. sequently to separate fractions boiling 105 to 190 F., ASTM distillation, F.: 190-300 F., 300420 F. and an oil residue. The oil IBP 245. residue was considered part of the catalyst oil and added 50% 253. thereto. 90% 271. The IDS-190 F. cut contained 0.2% olefins and only EP 283. a trace of aromatics by F.I.A. This cut was blended Octane number (RON) 25.5 (neat); 60.1 (3 cc. with the requisite amount of isopentane and n-pentane TEL added/gal.). as produced, found to have a Research ON of 88 neat, Panafi'ins, vol. percent 95.5. and 100.9 with 3 cc. TEL. The yield of this blend was Olefins, vol. percent 3.1. 75 volume percent based on raflinate feed. The yield of Naphthenes, vol. percent 0.0. the 190-300 P. out was 34 volume percent. On recycling Aromatics, vol. percent 1.4. this cut to extinction, one would obtain an ultimate yield Composition, vol. percent: of 114%. The yield of hydrocarbons boiling 300 to C 1.4 420 F. is 1.1 volume percent and that above 420 F., i-C 31.5 35 10 volume percent. The yield of propane was 5.3 volume n-C 21.7 percent while the total butanes recovered in this example i-C, 40.6. are less than those charged to the hydrogen fluoride ren-C 4.9. action zone indicating a total consumption of butanes in With continuous stirring the entire contents of the autothe T935901 clave are rapidly discharged into a 2-liter stainless steel Th6 Y -P advantagc Obtalned y 0111 optifatlon pressure vessel which is cooled in a Dry Ice bath to about can be readily Seen ffOm t data f the tables below. 80 C. Th gom m f th a t l are di h d In the runs the Udex rafiinate was that of Example I. in about to seconds, The pressure vessel is re- The conditions of reaction are noted in the tables.

TABLE II Run 54 5 69 Temperature, F -1 200 250 300 300 Contact time, min 60 3O 60 60 Pressure, max. (p.s.i.g.) 325 525 875 860 A. Udex raflinate:

100.6 1s3 4.5/1) 0. 4.5 1 103.3 184 it; Product distribution, hcbn. m percent .I 92. s 32. 5 95. 5 9s. 9

Weight Vol. Weight Vol Weight Vol. Weight Vol. percent percent percent percent, percent percent percent percent 1. Condensible gas: I

c, 1.13 1. 59 2. 52 3. 54 10.81 15. 07 26.1 as. s

4. 45 5. 43 2s. 3 34. 5 66. 6 81. a 41. 47. 75 74. 2 s4. 6 s5. 4 10s. 5 2. 09 2. as 3.14 3. 55 13. 15. 7 0.13 0.14 .43 .47 1. 51 1. e4 5. 1o 5. 51 12. 0 12. 95 12. 34 13. 29 2.17 2. 211 4. s2 4. 81 4. 78 5. 04 (1.27 0.21 1.28 1.28 1.15 1.75 3.118 3. 8s 5. 82 5.82 6. s3 5. 93 0.18 0.18 .09 .09 0. a2 0. 32 0. 97 0. s7 1. as 1. 59 12.14 12. 74 55. 3 55. a 25. 0 25. 0

12 1 High octane gasoline, vol. percent/RON 6 63/851 114/159 Recycle gasoline, vol. percent/RON B 1 67/49 27/28 Total gasoline vol. percent/RON B Y /6 1 /77 Ultimate yield high octane gasoline} percent 175 See footnotes at end or table.

. 2 344586 03 g r 2 t .31.3 .1 3 m u t 4.162% mnmomms. u 0 0 0 mm 4 1 n n \W m t U W 5% mm 04 VM mow 187075 Qw 1 e w B1 wmml fll n u 7 (l .t. m h lkml l n 3 9 59 3 4 w u 1 a Mfl 1 2 5 35 3 .27 t 943 m T slmaluogfim 34 W w w ti "mu s H a m ww fi mm 62 4 a m 0 1 mm 3 d s aa 4. 12 QT S S 0 1 Wm M .1 w 1 m w I a .l n d 20 4054 57 .4. l y 864 m dwisaanat In a m w m n 9100512 wmmmm... mm A H 51 U a m Yr Lm a $2 13 3 L5 7 U vw n m o w 2 1. f 673 P m m 8)e 9 m. mmm om n we mu m Y m mla 2n nfln .mm 799 3 7 12 n 00 0% I H. 7 GM m 16 4 97 57 .171. S A. 55 7 572 l O 52 n 2 0 Q R S t flw 9 I w W 6 3 Ct %6317 .WWWWNWQ 5 mm 1 1 n n u c m 6834 2 6 2 3 151 n r e S 6 1% 212 3 W m. S H er n. v. e e Wm. n W mi 0 r t T n u S 1 6 7 0 u m B D W t %65 ZUQWMWMMNBZ mm +n a Lm 0 91 9 10 u& r A 00 1 6 G w V D. l I G 005 05 u no U9 9 005 HT 79 D. 017 T H1 3 B as m3 3 2 HUN 9 9 4i Mn U n U "H 884 3924 10 Won Mm H m Hm PU n tt 5262124040185 Wflu EU n" u 11 um l flnu or m h u to 1 G1 m s m cw H 1 A n W s u TABLE II-Continued Run....-....

reaction system, for instance compare the results with those of runs 35 and 38 of Table II. Also, n-butane cannot entirely replace isobutane, see run 40 and run 35,

TABLE III 7 Calculated.

2 Based on (A) corrected to 100% recovery.

hole number.

In run 69 Where a substantial amount '01. percent/RON 6 runs 54 and illustrate the exceptional 35 results of run 37 can apparently be attributed to the high Product distribution, liebn. recy, percent Run.....

Product: distribution, hcbn. rec'y, 1\ercer1t.....-...-......--......-.......

,Yol.percent/RON See lootnotes, Table II.

although a part can be replaced as shown by run 41.

of n-butane was charged with the isobutane there was a net consumption of both isomers. The contact time in runs 35 and 38 was greater than optimum at the temperature of 300 F. but the treatment was beneficial even in run 38 employing the lesser amount of isobutane. The

TABLE IV Run 71 78 75 Feed Recycle naphtha Hydrogenated Udex HasOs-treatcd Udex roilinate railinate Temperature, F 250 250 250 Time, minutes 30 30 30 Pressure, max. (psi. 525 515 A. Udex rciimate:

Weight Vol. Weight Vol. Weight Vol. percent percent percent percent percent percent Product di tribution: I 8

O s 3. 4 4. 77 2. B2 3. 67 2. 46 3. 45

Yield:

High octane gasoline, vol. percent (RON) 1 B7 1 Recycle naphtha, vol. percent 33 26 Ultimate yield:

Pentane, vol. percent 6 1 190 111 160 High octane gasoline, vol. percent 8 143 1 0 131 1 Based on A+B+G. 1 Based on (Al corrected to 100% recovery. 5 [m5 l gible. I Calculated. B Rounded on to nearest whole number. 6 Assuming same yields from Ce+ naphtha.

The recycle naphtha feed in run 71, was the 190 to 300 F. cut from previous runs such as run 54 of Table II which cut was treated with cold 96% H 80 to remove olefins and aromatics. The recycle naphtha reported in the product is the approximate 190 to 300 P. out from 1 Averaging about 95% iso.

5 Assuming same yields from recycle naphtha.

platinum-alumina catalyst. The product contained 7.1% naphthenes. To obtain the feed for run 75, the Udex raflinate (1 liter) was treated with 100 m1. of 96% H 80 at room temperature and then with 150 ml. of fuming H 80 in 3 batches at to C. The product was washed consecutively with 96% H H O,

run 71. The feed for run 78 was made by hydrogenat- N OH and 0 and dned and comamcd 4 mg the Udex raifinate at 450 F. and 600 p.s.1.g. over a naphthcnes.

TABLE V Run 73 74 79 Temperature, F 250 250 250 250 Contact time, min 30 3D 30 30 Pressure, max, (p.s.i.g.) 555 555 520 540 A. Udex raffinatc:

G l 36. 5 1 35.4 :1 51 49.8 B. Other [eed n-lcntnne n-Pentane mHexane G 73 67 33. 2 1M1 .1 115.5 100 50 (J. Isobutenc vol. ratio, C/A'i-B" 273/1 211 3/1 G F8 168. 5 169. 2 h e 316. 5 299 300 D. Hydrogen fluoride, ratio 1 1.0 1 G 402 449 406 400 Product distribution, hebn. reey, percent 99. 2 97. 3 100.6 99.8

Weight V01. Weight Vol. Weight Vol. Weight Vol. percent 6 percent percent 1 percent percent 1 percent percent 5 percent 1. Condensiblo gas:

5. 6 -15.1 --99. 2 121.4 2. 4 15. 71 17. 72 20. 0S 6. 04 61. 99. 65 104. 2 S3. 0 (70.3% 6. 13 1. 01 1. 81 1. 50 1. 67 9.15 25. 74 20.8 3. 83 10. 9 11. 11 1. 73 (49% conv. of n-Ct) 7. 06 15. 72 15. 21

Il-Cr, Cl 1 s+.- (Udcx 52.6% conv.) (51% eonv. oiUdex) Catalyst oil 1.4 7. '76 0.8 23. 2 17.7 est. Ultimate high octane gasoline, vol. percent /RON 11 88/88 160 89 1 Added with some l-C during 20', stir 10' longer. 1 Vol. retio DJA+B+ O. 1 Based on (Al corrected to recovery. 4 Based on A+B+O. 5 Based on 11-C5 feed. 5 Based on Urlcx plus n-C consumed. Based on IJ-CA. 6 Based on Udcx plus n-C' consumed. p Blending O.N. 10 Calculated. H Assuming some yields from recycle naphtha.

1 5 The data of Table V show that without the C7+ feed (Udex raffinate), n-pentane and n-hexane are not advantageously treated in our process. On the other hand, runs 73 and 79, illustrate that these normal hydrocarbons can be converted to higher octane materials when mixed with the 6 feedstock in our system.

low octane number Udex raffinate and should be recycled, only the i-Cq alkylate will go to gasoline, and the i-C alkylate will have value only as recycle feed. Thus, concurrent alkylation would be at a disadvantage.

It is seen that addition of propylene to the feed leads to a greater production of propane (8.7 and 20.9 vs. 3.1).

TABLE VI Run 20 14 54 Temperature, F 200 200 200 Contact time, min. ,0 Pressure, max. (p.s.i.g.). 3G0 300 325 A. Udex ralfinate:

52 51. 0 73 73 71. 5 102. 5 B. Propylene: 7 5

Mi. (percent of A) 3 9 (5. 4) 14 5 C. Isobutane:

G 124.5 119.5 171.5 M 221 212 304 1). Hydrogen fluoride, ratio 1.0 1.0 1.0 G 205 301 405 Product distribution, hcbn. rec'yfi percent 02. 0 100 92. 8

Weight Vol. Weight Vol. Weight Vol. percent percent percent percent percent percent ondenslble Gas. 4 1 C 5. 60 7. 85 10.15 1.13 1. E5

16. 8 20. 1. 70 20. 8 4. 45 5. 43 61. 7 70. 20 58. 4 66. ['1 41. 85 47. 75 t. 2.84 3. 21 2. 55 2. 88 2.09 2. 36 2,2-DMB, 9 .44 .48 .85 .93 0.13 0.14 2,3-DMB+2-MP, 88 L. 0. 70 1D. 42 10. 57 11. 4 6.10 6.57 3MP, 74.5 i c. 2. 73 2. 93 4. 22 4. 45 2. 17 2. 29 5 20 20 85 85 0. 27 0. 27 3. 52 3. 52 4. 64 4.64 3.88 3. 88 64 0. 18 0. 18 s 0. 97 0. 97 c. 65. 3 65. 3

High octane gasoline, vol. percent/RON Recycle gasoline, vol. percent/RON Total gasoline, vol. percent/RON Ultimate yield gasoline, percent D/(A-i-13+C). 2 Based on (A) corrected to 100% recovery. Calculated. 1 Assuming same yields from recycle naphtha.

The data of Table VI establish that the provision of excessive olefin in our process is uneconomic. The pertinent data of Table VI are gathered in Table VII.

TABLE VII Bun 20 14 54 Products 3 no. 2, 3- Ult. 1 2,33 Ult. 1

(1) Propane, wcight,per-

cent 5 6 8.7 10.15 20.9 1 13 3.1 (2) Catalyst oil, weight,

percent 15.7 24. 4 16. 7 34. 3 7. 7 23.0 (3) Gasoline, vol.percent. 91 141.7 01. 75 188. 5 63.3 189 (RON) (89. 5) (SB. 6) (89.3) Gasolinefbyproduct, ratio fl/(l-l-Z) 4.28 3. -12 7. 23 (4) Recycle naphtha, vol.

I: 35.7 51.4 66.5 (RON).. (a? 49) so 5 (5) Corrected for aikylate, gasoline, v percent 81. 55 127 56. 75 118.8 63.3 189 Ultimate yield based on recycling #4 to extinction. 1 Allrylate equals 1.75 times propylene volume and has RON or 9091 see article on HF Alkylation (U.O.P.), page 163, Oil and Gas Journal; Mar. 25, 19 7 about 190 F. (n-C is contaminated with unconverted 3 Based on A-l-B+C.

The gasoline in all cases had about the same research octane number; however, only run 14 (20% propylene added) produced as high an ultimate yield (188.5 vs. 189 vol. percent (as run 54 (no propylene added). Referring to the propane and catalyst oil by-products for these two runs, we see that the ratio of gasoline to these lay-products is much more favorable when no propylene is fed (7.23 vs. 3.42).

Run 14 consumes 20% of propylene once-through at a 48.4% conversion level. On a recycle basis, to produce 188.5% ultimate yield of gasoline, the process would consume 40.5% of propylene. In a commercial isobutane HF alkylation plant this amount of propylene would produce 40.5 l.75 or 72.5% of alkylate (RON=91). Instead, in run 14 this amount of propylene produced only 29.2% of by-produots (17.8% more propane and 11.3% more catalyst oil than in run 54), which is, of course, a poor trade for the potential 72.5% of highoctane alkylate.

The data of Table VIII below illustrate that our process can be used to convert relatively pure n-paraffins, providing they have at least 7 carbon atoms, into higher octane materials such as isopentane. 2,4-dimethyl pentane underwent a similar reaction even in the presence of a naphthenic impurity while the latter apparently inhibited the conversion of n-octane into lower boiling products.

TABLE VIII 1 Run 030-10 84 930-7 88 89 90 93 Temperature, F.- 251 253 Contact time, min. 30 30 Pressure. max. (p.s.i.g 510 475 A. Liquid hydrocarbon. 2, 4-DMP n-Octane G 67. 2 95 Q5 MeCyclo Cs MeCyclo Cu 3. 9 3. 9 5. 5. 0

169 171. 4 300 304 Hydrocarbon recovery, percent 100 98. 15

Product distribution, weight S,3-DMP+MCP 0. 52 0. 27 0.16 2,3I)MP+2- and CyCu 43. 51 1B. 23 7.1 5. 45 4.10

1 Hydrogen fluoride catalyst 400 :l: 5 g. 1 Based on A+B+O. Based on (A) corrected to nodes basis. 1 Ca to Cu. 5 Cm.

We claim:

1. In a method for converting parafiinic hydrocarbon boiling in the 0 motor fuel range, the steps which comprise contacting in the liquid phase isobutane and the parafiinic hydrocarbon with a catalyst consisting of hydrogen fluoride at a temperature of about 100 to 400 F. under a pressure sufficient to maintain the liquid phase in an essentially olefin-free reaction zone, said contacted isobutane being at least about 50 weight percent of said paraffinic hydrocarbon and separating a hydrocarbon boiling in the motor fuel range.

2. The method of claim 1 in which the 0 feed contains about 15 to 35% normal paraflins and about 65 to 85% isoparaffins.

3. In a method of converting paratfinc hydrocarbon boiling in the 0 motor fuel range, the Steps which comprise contacting in the liquid phase isobutane and the paralfinic hydrocarbon, said hydrocarbon containing about 15 to 35% normal parafiin and about 65 to 85% isopraflin, with a catalyst consisting of hydrogen fluoride at a temperature of about 175 to 300 F. and at a pressure suflicie-nt to maintain the liquid phase in an essentially olefin-free reaction zone with there being a net consumption of butanes of at least about 5 weight percent based upon said paratfinic hydrocarbon, said contacted isobutane being about 75 to 400 weight percent of said paraifinic hydrocarbon and separating a hydrocarbon boiling in the motor fuel range.

4. The method of claim 1 in which the hydrocarbon feedstock consists essentially of at least 10% of said C7+ motor fuel range material and a member selected from the group consisting of n-pentane and n-hexane.

5. The method of claim 1 in which the temperature is about 175 to 300 F.

6. In a method of converting a straight run hydrocarbon petroleum fraction boiling in the motor fuel range, the steps which comprise contacting said hydrocarbon fraction with a platinum metal-alumina catalyst in the presence of free hydrogen at a temperature of about 750 to 1000 F. and a pressure of about 50 to 1000 p.s.i.g. to provide a product boiling in the motor fuel range with increased octane value, separating from this product paralfinic hydrocarbon boiling in the 0 motor fuel range, contacting in the liquid phase iso-butane and the parafiinic hydrocarbon with a catalyst consisting of hydrogen fluoride at a temperature of about to 400 F. and under a pressure sufficient to maintain the liquid phase in an essentially olefin-free reaction zone, said contacted isobutane being at least about 50 weight percent of said parafiinic hydrocarbon and separating the hydrocarbon boiling in the motor fuel range.

7. 'lhe method of claim 6 in which the (3 feed contains about 15 to 35% normal paralhns and about 65 to 85% isoparafiins.

8. In a method of converting a straight run hydrocarbon petroleum fraction boiling in the motor fuel range, the steps which comprise contacting said hydrocarbon fraction with a platinum metal-alumina catalyst in the presence of free hydrogen at a temperature of about 750 to 1000 F. and a pressure of about 50 -to 1000 p.s.i.g. to provide a product boiling in the motor fuel range with increased octane value, separating from this product paraflinic hydrocarbon boiling in the 0 motor fuel range, contacting in the liquid phase isobutane and the paraffinic hydrocarbon, said hydrocarbon containing about 15 to 35% normal paraffin and about 65 to 85% risoparatfin, with a catalyst consisting of hydrogen fluoride at a temperature of about to 300 F. and at a pressure suflicient to maintain the liquid phase in an essentially olefin-free reaction zone with there being a net consumption of butanes of at least about 5 weight percent based upon said paraflinic hydrocarbon, said contacted isobutane being about 75 to 400 weight percent of said parafiinic hydrocarbon and separating a hydrocarbon boiling in the motor fuel range.

References Cited in the file of this patent UNITED STATES PATENTS 2,220,092 Evering et al. Nov. 5, 1940 2,349,458 Owen et a1. May 23, 1944 2,370,144 Burk Feb. 27, 1945 2,405,996 Burk Aug. 20, 1946 2,527,529 Cede Oct. 31, 1950 2,540,379 Ridgway et a1. Feb. 6, 1951 2,592,740 Ridgway Apr. 15, 1952 2,767,124 Myers Oct. 16, 1956 2,946,833 Kimberlin et al. July 26, 1960 2,965,693 Kramer Dec. 20, 1960 2,971,037 Gilbert et a1 Feb. 7, 1961

Claims (1)

1. 6. IN A METHOD OF CONVERTING A STRAIGHT RUN HYDROCARBON PETROLEUM FRACTION BOILING IN THE MOTOR FUEL RANGE, THE STEPS WHICH COMPRISES CONTACTING SAID HYDROCARBON FRACTION WITH A PLATINUM METAL-ALUMINA CATALYST IN THE PRESENCE OF FREE HYDROGEN AT A TEMPERARTURE OF ABOUT 750 TO 1000*F. AND A PRESSURE OF ABOUT 50 TO 1000 P.S.I.G. TO PROVIDE A PRODUCT BOILING IN THE MOTOR FUEL RANGE WITH INCREASED OCTANE VALUE, SEPARATING FROM THIS PRODUCT PARAFFINIC HYDROCARBON BOILING IN THE C7+ MOTOR FUEL RANGE,

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