US2361611A - Hydrocarbon conversion - Google Patents

Description

(kt-31, 1944. E. L D ouvlLLE ETAL HYDROCARBYON CONVERSION i Filed oct. (15, 1940 Patented ct. 31, 1944 UNITED STATES PATENT o1=1=1cE IIYDROCARBON CONVERSION i Edmond l..A dDuville and Bernard L. Evering,

Chicago, lli.. anaignors to Standard Oil Com- NIW. Chicalo. lll.. corporation oi' Indiana Application October l5. 1940, Serial No. 361,270

(Cl. 19d-48) 3 Claims.

This invention relates to certain new and useful improvements in the conversion of low antiknock hydrocarbon fractions to useful products such as high antlknock motor fuels and high antiknock aviation fuel. The invention has in view the provision of a combination process in which the low antiknock components of virgin petroleum naphthas aie treated in an integrated process which is highly efilcient in attaining the ultimate object of maximum yield of high antiknock naphthas.

Virgin naphthas of motor fuel distillation range contain variable proportions of straight- 'chain 'parailln hydrocarbons and naphthcnic hydrocarbons with minor proportions of aromatic hydrocarbons. For example. virgin naphthas from Pennsylvania crudes and Michigan crudes may contain as high as eighty to ninety per cent parafiins, the major part of which are straightchain parafllns and as low as five to fifteen per cent of naphthenic hydrocarbons. On the other hand, naphthas from Mid-Continent crude and from California crude contain as high as twenty per cent to thirty-live per cent of naphthenic hydrocarbons and from fifty to seventy per cent of paraffin hydrocarbons. Since the straight-chain paraiiin hydrocarbons and the naphthenic hydrocarbons are of low antiknock value it is highly desirable to convert the former to branched chain parafi'in hydrocarbons and the latter to aromatic hydrocarbons boiling within the gasoline boiling range in order to obtain maximum yields of high antiknock value naphthas. i

It is an vobject of this invention to provide an improved process for converting the straightchain parafiins in a parainic naphtha which also contains naphthenes and aromatics, to isobutane. Another object of this invention is to convert the straight-chain parain hydrocarbons of such a naphtha to condensible oleiins. Still another object of this invention is to react the isobutane and the condensible olens obtained from the straight-chain paraiins to form high antiknock branched chain parafnic naphthas which distill 'in the aviation gasoline boiling range. A further object of this invention is to raise the' antiknock properties of a low antiknock naphtha by removing therefrom, in a catalytic conversion process, a substantial part of the low antiknock straight-chain paraffin hydrocarbons. Another object of this invention is to convert the low naphthenic hydrocarbon components of a paraiiinic naphtha to high antiknock aromatics. A still further object of this invention is to provide a continuous process in which the catalysts employed may'be utilized eiectively for the production of high antiknock motor fuels v metallic oxide activated alumina catalyst.

the accompanying drawing which is a diagram-v matic representation of an apparatus for carrying out our invention.

We have found that straight-chain paraffnic naphthas can be catalytically converted to a gaseous product which is largely isobutane by passing the naphthaover a gel catalyst of-the alumina activated silica type at low space velocities o! the order of 0.01 to 1.0 volume of liquid per volume of gross catalyst space per hour and at temperatures withinV the range-of 575 F. to 850 F. This conversion process is carried out at absolute pressures within the range of fifteen pounds and sixty-live pounds per square inch. Moreover. we have found that by operating in the upper part of the above temperature range, say 800 F. to 850 F., the gaseous product consists primarily of isobutane and condensible oleilns in such proportion that the gaseous mixture may be alkylated to produce bz anched chain paraiiinic hydrocarbons boiling within the gasoline range. We have also found that if a paraflinic naphtha is passed over a solid catalyst comprising a mix2 ture of a cracking. type catalyst, such as alumina activated, silica and a dehydrogenation type catalyst, such as metallic oxide activated alumina, at temperatures within the range of 750 F. to 900 F., at space velocities within the range of 0.01 and 1.0 volume of liquid feed per volume of gross catalyst space per hour and at absolute pressures within the range of ten pounds and sixty pounds per square inch, relatively large volumes of condensible olens are produced. Under these operating conditions,any naphthenes present in the feed to such a catalyst bed are also dehydrogenated to high antiknock aromatic hydrocarbons.

In practicing the invention, we prefer to fractionate at least a major part of the naphtha feed stock in order that we may convert the paraffin -hydrocarbons of less than seven carbon atoms directly to the branched chain paralns. This lower boiling fraction of the naphtha feed stock will contain only limited amounts of aromatic hydrocarbons an'd hence We are able to isomerize it to branched chain paraffin hydrocarbons over aluminum halide type catalyst. Aromatic hydrocarbons tend to deteriorate the aluminum halide catalyst. Thus we reserve the C1 and higher fraction of the feed, which contains aromatics and naphthenes in addition to paraiilns, for conversion over the alumina activated silica gel cracking catalyst and/or for conversion over the mixed alumina activated silica gel catalyst and The isomerized Cs and C5 hydrocarbons are blended with the alkylated isobutane and utilized as aviation fuel or the blend may be added to the converted naphtha to produce high antiknock motor fuel. The C4 and lighter fraction of the virgin naphtha feed passes to the alkylation step which C4 and lighter fraction, a Cs-l-Cs fraction and a C1 and higher fraction. The Cv and higher fraction is heated to a temperature within the range of 575 F. to 850 F., preferably 750F. to 800 F., and is thenfpassed over a catalyst bed in a conversion zone which may be designated as a cracking zone in which 'the catalyst comprises an alu- The product from thisA mina activated silica gel. conversion step is fractionated into an overhead gaseous fraction which is extremely high in isobutane content. 'I'he major part of the liquid fraction is recycled to the heating step' and thence to the said cracking zone to convert more of the straight-chain paraiiins to isobutane. A part of the liquid fraction is returned to the furnace and reheated in a separate coil to a higher temperature, for example, 750` F. to 900 F., but preferably from 850 F. to 900 F., and then passed to a conversion zone containing a mixture of .a cracking catalyst and a dehydrogenation catalyst such as alumina activated silica mixed with a metallic oxide activated alumina catalyst. We have found that chromium oxide activated alumina is a suitable dehydrogenation catalyst for mixing with the alumina-silica catalyst in this conversion Zone which may be designated as a dehydrogenation zone` The product from this conversion step is fractionated and the high antiknock value liquid fraction from which a major part of the paramns have been removed, either by cracking to isobutane in the 'cracking zone or to condensible olefins in the dehydrogenation zone and in which the naphthenes have been converted to aromatics, is withdrawn for blending. y If desired the Cv and higher fraction may be sent to the dehydrogenation zone and a part of the liquid fraction from the fractionation of the dehydrogenation product may be recycled to the cracking zone, in which case the products from the two conversion zones may befractionated in a common fractionator, thus eliminating the necessity of a separate fractionation of the catalytically cracked product. The isobutane plus normal butane overhead product from the fractionation of the product of the cracking step is alkylated with the condensible oleiins from thefractionation of the product of the dehydrogenation step by intimately contacting the liquefied mixture with sulfuric acid of 95% to 103% concentration at temperatures chain C5 and Ce paramn hydrocarbons by intimately contacting this fraction and the normal butane -with aluminum halide or aluminum halide-hydrocarbon complex catalyst, preferably aluminum chloride or aluminum chloride hydrocarbon complex, at temperatures within the range of 250 F. to 400 F. The amount of catalyst used may be from ve to fty per cent of the Weight of the hydrocarbon feed. The catalyst is activated by the addition of hydrogen chloride to the reactor to the extent of from five-tenths to five per cent of the weight of the hydrocarbon feed. The isomerization reaction is carried out in the presence of hydrogen or a gas containing hydrogen. We prefer to use at least a part of the hydrogen and methane containing gas from the dehydrogenation step for this purpose. Pressures in the isom.

erization step may be within the range of 200 to 3000 pounds per square inch. The product from the isomerization step is passed to the dehydrogenation product fractionating step from which the isomerized C5+Cs cut is withdrawn as a high antiknock blending naphtha, isobutane being separately recycled to the alkylation step.

In converting low antiknock value naphthas, which contain very low percentages of naphthenes, to high antiknock blending naphthas by our process it is sometimes unnecessary to utilize the dehydrogenation step in which the feed is contacted with the mixed cracking and dehydrogenation catalysts. We have found (see Table I below) that whenoperating the cracking reactor at a temperature of 823 F. the gaseous product contains a weight ratio of isobutane to total gaseous olefins reactive in a sulfuric acid alkylation process of approximately 29.0 to 27.6. which corresponds to a mol ratio of approximately 22.0 to 25.6. Keeping in mind that one mol of olefin reacts with one mol of isobutane in the alkylation reaction, it is readily seen that within the range of 30 to 150 F. for periods of from five minutes to sixty minutes and at pressures sucient to keep the reactants in the liquid phase. The ratio of isobutane to condensible olens should be at least 2 mols of isobutane for one mol of oleiins. The alkylated product is separated from the unreacted gases by fractionation and used as a high antiknock value blending naphtha for aviation gasoline or it may be blended with the high antiknock value naphtha, from the cracking and dehydrogenation steps. Unreacted normal butane is passed to a catalytic isomerization step to produce additional isobutane for the alkylation reaction.

The Cs-i-Cs fraction from the virgin naphtha, feed is isomerized along with the recycle normal' butane from the alkylation reaction to produce simultaneously isobutane and branchedy by isomerizing the normal butane produced in the` cracking reactor and also the butane' in the overhead fraction of the fresh feed it is possible to maintain a properly proportioned feed to the alkylation reaction without using the dehydrogenation reactor to produce olens. -Any deficiency in hydrogen for pressuring the isomerization reaction may be furnished from an outside source for this type of operation. The nal products from this type of operation are the same as from the process when both the cracking and dehydrogenation reactors are used except lthat the product containsL a lower percentage of aromatic hydrocarbons. This type of operation may also be used when it becomes necessary to regenerate the catalyst in the dehydrogenation reactor.

` TABLE I CA'rALY'rIc CONVERSION or NORMAL HEPTANE CUT Space velocity 0.06 vol. of liquid feed per vol. gross Temperature F 76H SIR x03 Weight per cent reacted l. 4 3. 7 I4. 0 Product-Weight per cent:

H i. 4. 2 4. 8 4. 4 2. 3 4. li

l. l l. 6 2. 7

I3. 8 ll. 6 l0. 2

I8. 3 I7. 2 l5. 5

NC4Hm 7. 2 9. 0' 5. -l

' scribed below.

In a third embodiment of the invention the dehydrogenatlon reactor may be operated without the cooperation of the cracking reactor. Referring to Table II we have found that if a paraftln hydrocarbon fraction is passed over a mixed catalyst comprising an alumina activated silica gel and a chromium oxide activated alumina catmal butane is formed but the major part of the gaseous product consists of condensible olefins' which may be polymerized by either the cold sulfuric acid process or the hot sulfuric acid process. The polymerization process can be carried out in the alkylation reactor and the spent alkylation sulfuric acid catalyst can be reduced in concentration to a concentration within the range of 60% to 80% acid and used for this step.

The polymerized product4 can be fractionated to' The polymerization 4by the hot acid method isA alyst at 893 F.' no isobutane and very little norcarried out at temperatures within the range from about 160 F. to 200 F. at atmospheric pressure with contact time within the range of from thirty seconds to three hundred seconds.

'I'his type of operation is desirable when the feed stock consists of a naphtha containing a rela- I tively low percentage of paramos and a high percentage of naphthenes.; This type of operation may also be utilized during those periods when it becomes necessary to regenerate the catalyst in the cracking reactor. During operation according to this method the isomerizer may be used to isomerize only the Cs-l-Cs fraction of the fresh feed.

VTABLE -II CA'rALYrrc Conversion or NORMAL `Hari-Am: CU'r Space velocity 0.06 vol. of liquid gross catalyst space per hour. mospheric CATALYST-MIXED SILICA-ALUMINA ANT) 95% A1203.5% Creo;

feed per vol. of

Pressure-at- Other embodiments of our invention, such as passing fresh feed to the cracking reactor and to the dehydrogenation reactor in parallel with separate fractionation of the product and separate recycle to the respective reactors or with fractionation in a common zone and consequent recycle of mixed liquids to either or both reactors or a combination of any of themethods of operation described above are contemplated by this invention. For the purpose of illustrating the invention operation ,with a feed stock such as a Mid-Continent naphtha containing straightchain parafns, aromatics and naphthenes is de- Referring now to the drawing, which is a simpliiled ilow diagram illustrating one embodiment of our invention, a virgin naphtha boiling within the range from about 60 F. to about 400 F. is

the fresh feed through fractionator I6 in order to separate the maior part of the lower boiling fraction of the feed which may beisomerized directly to high antiknock value light naphtha suitable for blending in aviation gasoline. Fractionator I6 is provided with reux means and reboiler means I8. A side stream in line |33 containing primarily pentanes and hexanes is withdrawn from fractionator I6 by pump |34 as feed stock -for the isomerization step which is described in more detail below. The C4 and lighter overhead fraction from fractionator i6 is passed by,valved line 96 to alkylator |05 which is also described below.

The C7 and higher fraction of the naphtha is passed by means of pump 20 in line |9 to valved line 2| where it is mixed With non-fractionated virgin naphtha feed and the mixture is passed through valve 23 in line 22 ing coil 25 in furnace 24, valve 90 in line 89 being closed. The naphtha is heated in coil 25 temperature within the range of 575 F., preferably '750 F. to 800 temperature range it is passed 26 and valve 21 to lcatalyst chamber 28 which is lled with a refractive gel catalyst of the alumina activated silica type. This catalyst may F. to 850 F., within which be prepared by immersing commercial silica gel in an aqueous solution of an aluminum salt. draining ofi the supernatant solution, Washing the gel, and drying and heating the product for several hours at temperatures of from 800 F'. to '1100 F. The finished catalyst contains from 95% to 99.5% silica gel activated with from 0.5% to 5% alumina. We may also'employ an alumina-activatedv silica gel catalyst prepared by cogelling an alumina sol with a silica sol in such proportion as to produce a cogel comprising from about one to twenty parts by weight of alumina and from about to 99 parts by weight of silica in the dried gel product. Other refractory type cracking catalysts such as acid treated clay and silica gel activated with such metallic oxides as magnesium oxide and boron oxide may also be used incracking chamber 28. The space velocity in catalyst chamber 28 at absolute pressures within the range of l5 pounds per square inch and 65 pounds per square inch, is Within the range of 0.01 and 1.0 volume of liquid feed per volume of apparent catalyst space per hour based on the parain hydrocarbon content of the feed. We prefer to operate `at space velocities within the range of 0.04 and 0.1 volume of paraiin hydrocarbons per volume of catalyst per ho-ur or at such a space velocity that from about 5% to about 15% of the straightchain parafiin hydrocarbons in said naphtha are converted per pass through the reaction zone.

. The product from catalyst chamber 28 is with.

and it is passed which leads to heattoa` via transfer line which is operatedA product from reaction chamber 28 the product may be passed directly to fractionator ll via valved lines i2 and 53 with valve iid open and valve l5 closed. Fractionator lll with reflux producing means i6 and reboiler lil serves to separate the C4 and lighter -hydrocarbons which contain a large percentage of isobutane from the liquid product. The butanes pass via valve 95 in line 95 and line Tl to the alkylator feed line 98. The liquid product is withdrawn from' fractionator il through side drawoi line 55. Hydrocarbons boiling above normal motor fuel boiling range may be drawn off through bottom drawoi 58a. By proper adjustment of valve 50 in line 59 which joins line 58, valve 5i in line 52, valve 55 in recycle line 55 and valve 55 in nished product line 5l, the liquid product from fractionator lll may be sent in part to coil 5l in furnace 25 through lines t5 and 05 and thence by linel 92 to dehydrogenation reactor 50, and the major part of the product may be recycled by means of pump 53 through lines 52 and 55 to coi-l 25 in furnace 25 and thence to' reaction chamber 20- for the production of additional isobutane. 'We prefer to operate with4 valve 55 in line 5l closed when processing a vir-gin naphtha containing appreciable amounts of naphthenes in order that these be converted to aromatics in reactor 55 before blending the liquid product to form high antiknock gasoline.

Referring to Table I above, a conversion per pass of approximately eleven per cent by weight of-the paran hydrocarbons may be expected in reactor 20 when operating within the range of 750 F. to 800 F. In order to convert a major part of the paraln 'hydrocarbon components of the feed in reactor 28 we maintain a recycle ratio of from about eight to about twenty parts by weight of recycle to one part by weight of virgin feed. When processing feed stocks containing a relatively high percentage of paraffin hydrocarbons the recycle ratio will be inthe upper part of the above range while a feed stock containing a relatively low percentage of paraln hydrocarbons will not necessitate such high recycle ratios.

In order to more completely reduce the paraifin hydrocarbon content of theproduct in .side drawoff line 58 and thereby produce condensible olefins and also to dehydrogenate the naphthenes-therein contained to aromatics, valve 50 is partially opened and the stream is directed to furnace coil @l in furnace 20 via lines 85 and 89 by means of pump 53. This stream is heated to a temperature within the range from about '750 P'. to about 950 F., preferably 825 F. to 900 F., and passes by line 92 to catalytic reaction chamber 58 which is filled with a mixed catalyst comprising from about 1 to 10 partys by weight of alumina activated silica .gel catalyst intimately mixed with from about 99 to 90 parts by weight of a. metal oxide-activated alumina catalyst such as an activated alumina supported chromium oxide catalyst. We prefer to use in reactor 58 a catalyst comprising alumina impregnated with 5 to- 20 parts by weight of chromic oxide, preferably parts by Weight of chromic oxide, mixed with a second catalyst prepared by absrptolytically depositing A1203 on Si02 gel as described above. The chromic oxideactivated alumina catalyst may be prepared by immersing commercially activated alumina in an aqueous solution containing the desired amount of chromic oxide in the form of the acid or suitable salt, evaporating oi the water, drying the product and heating the same at 1200 F. for one hour. The powder is then pelleted to the desired particle size. Other metallic oxides. such as molybdenum oxide and vanadium oxide may also be used to activate the alumina by using the appropriate water soluble acid and/or water soluble salt. The space velocity in reactor 50 may be within the range of 0.01 to 1.0 volume of paraffin hydrocarbons (liquid) per gross volume of catalyst per hour. The operating pressure (absolute) in reactor 58 is within the range of from 10 pounds to 65 pounds per square inch. In reactor 55 the major part of the remaining paraffin hydrocarbon components of the naphtha are dehydrogenated and cracked to a gaseous product, a high weight per cent of Which product consists of condensible olens suitable for feed to the alkylation step which is described below. We have found (Table II above)r that a conversion of 14 per cent per pass of pararn hydrocarborm over the above mixed catalyst may be obtained at a temperature of 893 F., at space velocity of 0.06 volume of liquid feed per volume of catalyst space per hour and atmospheric pressure and hence, in order to reduce the parainic content of the residue naphtha to a minimum, we prefer to operate at a recycle ratio of from about five to twelve parts by weight of naphtha bottoms from the fractionation of the product from reactor 58 (fractionator l2) to one'part by weight of feed to reactor55 from line (i9. The naph- `thenes of4 the naphthaare converted to aromatics in reactor 58 thus further increasing the antiknocl: value of the residual naphtha.

The product from reactor 50 passes via line 59, valve 00, cooler 5i, compressor C32 and line 03 to hydrogen release drum 55 where hydrogen and non-condensible gases, such as methane, ethylene and ethane are separated and pass via lines-55 and 6l to the isomerization reactor Sill. These may be passed in part to the fuel burning line through valved line 55 which joins line G5 by adjusting valve 55 in line 55 and the valve in line 58, or any excess of these gases not required in isomerization reactor Hill, when reactor 50 is on stream in the reaction part of the cycle, may be passed to storage for use in reactor iS' when the catalyst in reactor 58 is being regenerated. Still another use for this hydrogen is for the hydrogenation of polymer produced when alkylator |05 is being used as a polymerizer. which operation is discussed below. Also, the hydrogen may be used to suppress carbon formation in reactor 58. Condensate from drum 50 to which product from line 53 is introduced at pressures within the range of 300 pounds to 400 pounds per square inch.' passes by line l0 and pump 'H to fractionator 72 which is provided with reboiler i3 and reflux means lil. Residual non-ccndensible gases pass overhead via valved linel'i5 to the fuel burning line. A side stream, predominantly C3 and C4 hydrocarbons, is withdrawn from fractionator 'l2 through line 16 and passes via line Tl to alkylator hydrocarbon feed line 98. This stream in line 16, which consists in largepart of condensible olelins, may be directed through line '18 when alkylator |05 is being used as a polymerization reaction chamber. Fractionator 'l2 is also used to separate the liquid isomerizer reaction products which are introduced to fractionator via line |55 and hence a Cs-l-C side stream is trapped out of fractionator 12 which is passed via lines 79 and l30 by pump to product line 51 for blending with the aromatic bottom product from fractionator 'l2 or this Cs-l-Cc stream may be passed via, lines 18 and 18a to an aviation gasoline blending tank (not shown). The aromatic naphtha bottoms are withdrawn from-fractionator 12 via line 8| and pump 82, a major part of this stream being .recycled to coil 8| in furnace 24 vialine 85, the Vremainder being eliminated as high antiknock naphtha via line 51. any hydrocarbons boiling above normal motor fuel boiling range may be withdrawn through bottom drawoff 8|a.

As stated above, we may operate our process using a common tower to fractionate the prod- 'ucts from the catalytic cracking step and the catalytic dehydrogenation step. By properly adjusting valve 23 in line 22 and opening valve .88 in line 88 the C1 and higher fraction may be passedl via heating coil 8|, line 92 and valve 83 to dehydrogenator 58 for the initial conversion Y step in which condensible oleflns and hydrogen are produced. The product is flashed in release drum 64 and fractionated in fractionator 12 and the major part of the liquid bottom fraction may be recycled via line 8|, pump 82, line 85 and lline 86 to coil 25 via line 26 and thence vto reactor 28 where isobutane is produced in the cracking operation. The product from reactor 28 is withdrawn via line 28 and passes to the dehydrogenator product line 58. By adjusting valve 84 in line 8|, valves 83 and 88 in line 85 and valve 81 in line 88, the liquid fraction from fractionator 12 may be proportioned into nished high antiknock fraction drawoif, recycle feed to the dehydrogenation step and recycle feed to the catalytic cracking step as desired. Virgin C7 and higher naphtha may be passed in parallel to reactors 28 and 58 by partially opening valve 23 in line 22. The chief advantage of passing the major portion of the virgin feed tothe dehydrogenation step with operation of the catalytic cracking reactor primarily on recycle from the dehydrogenation step is possible heat economy without sacrice of the desirable dehydrogenation of naphthenes to aromatics. 'Ihis type of operation-is desirable when the virgin feed stock is high in naphthenic hydrocarbon content. When the virgin feed stock is highly paraflinic with very low naphthenic hydrocarbons present in the feed, reactor 28 may beV operated at higher temperatures, say 800 F. to 850 F., and the product may be fractionated either in fractionator 4| or 12 as described above.

Referring now to the alkylation'step of our process, the butane overhead stream from fractionator I8 passes via line 88 and valve 81 as alkylator feed stream to line 88 where it is mixed with the isobutane rich gaseous stream from line 11 which connects with fractionator 4| overhead discharge line 85. Isobutane may be introduced to line 88 from an external source through valve line 88. condensible oleflns pass via fractionator 12 drawoi line 18 to line 11 and thence to line 88 or these oleilns may be added directly to the acid catalyst valved line |86 through line 18. Condensible olefins from an,external source may be added either directly to the acid catalyst by means of lines |88 and |8| or the olens may be' `passed through lines |88 and |82 to line 11. The mixed butane and olefin stream is compressed by compressor |83 and passes by line |84 to alkylation reactor |85 where the mixture is intimately contacted with sulfuric acid of 95% to 103% strength which is introduced to line |84 via line |86 and valve |81 by means of pump |88. Alkylation reactor |85 is equipped with a mechanical stirrer |88 and a temperature control jacket ||8 15 with contact times withink the range of from about five -minutes to sixty minutes. The alkylate-acid mixture is transferred to settler by means of pump ||4 in valved line H3. In settler-*H5 the acidseparates as a lower layer and it may be recycled through line H6, pump H1, line H8 and line |88 to line |84. Spent sulfuric acid may be withdrawn through valved line I8 to be either concentrated for reuse in the alkylation step or to be diluted for use in the polymerization step described below. The alkylate and -unreacted gases are passed vialine |28 and pump |2| to a caustic wash step and a water wash step, which steps are not shown in the flow diagram. The washed product is heated in heater |22 and passes by line |23 to fractionator 24 which is equipped with reboiler |25 and reflux producing means |26. Propane and lighter hy- .drocarbons pass from fractionator |24 through valved line |21 to the fuel burning line. A side stream consisting predominantly of unreacted isobutane and normal butane is withdrawn through valved line |28. This stream may be recycled to the alkylationstep feed line 88 via valved lines |28 and |28 or to isomerizer feed line |33 via line |28 for the production of additional isobutane. A second side stream consisting of alkylate having an octane number of to CFR-M and an end point suitable for blending in aviation gasoline orsuitable for blending in finished motor fuel is withdrawn through line |38 and passes to line |3| and thence to line 18a for blending with the isomerized Cs-i-Cs cut or the alkylate may be sent directly to finished motor fuel line 51 via line |38. Bottom drawoif from fractionator |24 consists of a small amount of alkylate distilling above the gasoline boiling range when operating this fractionator to produce 400 F. end point alkylate drawoff in line |88 for blending with motor fuel. When so operating these bottoms are withdrawn through valved line |38a and accumulated as additional feed to alkylator |85 which may be operated at higher temperatures, for example, from F. to F., as a depolymerization-alkylation reactor to produce a product boiling within the gasoline range when charging such a stock as the bottoms from tower |24 with condensible olens. When operating to produce an alkylate for blending in aviation gasoline thehigher -boiling fraction may be further fractionated to produce additional alkylate of 400F. maximum boiling point for blending in motor fuel and the fraction distilling above 400 F. may be accumulated and used as described above.

As stated above we prefer to isomeriz'e the C5+Cs` fraction of the virgin naphtha directly to highly branched parafns for`blending in the finished products. On the other hand, it is not desirable to crack this fraction in reactor 28 since the increased yield of non-condensible gases which represent a volume loss of product would be increased and moreover this light fraction is essential in blending to form a balanced aviation gasoline or motor fuel product to furnish desirable volatility characteristics and if isomerized the fraction contributes materially to the antiknock value of such products. Hence this frac tion is drawn oil from fractionator I8 through through which may be circulated a suitable coolline |33 by pump |30 along with recycle butane from line |28. A light naphtha fraction containing C4, C5 and Cs paraflins may also be introduced from an external source via line |3341. The mixed hydrocarbon stream is heated in heater |35 which may be a heat exchanger in which the heat is furnished by the product from reactor 20 or reactor 58. Aluminum chloride or aluminum chloride-hydrocarbon complex catalyst, which is` activated with hydrogen chloride, isintroduced to the hot stream in line |36 via valved line Mil and pump M2. Hydrogen chloride is added to the catalyst in line Uil through valved line 850 and the mixture passes tc isomerizer itl which is equipped with rapid stirring means i323. Reactor lill may be suitably jacketed and a heating uid is circulated therethrough by means of valved connecting lines i353 and i130. The isomerization reaction is carried out in the presence of hydrogen or hydrogen mixed with non-condensible gases supplied by the dehydrogenation product via line til and compressor M3. Hydrogen may be furnished from an external source through valved line and line l. We operate isomerization reactor itil Within the range of 250 F. to 400 F., preferably at about 330 F., and at pressures Within the range of 200 pounds to 3000 pounds per square inch. The contact time in reactor il may be :from to 100 minutes.

Part of the catalyst plus isomerized product mixture is withdrawn from reactor l'l by pump |05 through line lllrand passesl by line it to separator lill which operates as a catalyst settling Zone and also as a release drum in which the hydrogen and hydrogen chloride is separated from the liquid product. The pressure in separator Mill is lowered to the range of 150 pounds to 350 pounds per square inch by opening pressure release valve |53 in line |52 and the mixed hydrogen and hydrogen chloride is recycled via line |50 to catalyst feed line Ml and thence to reactor |3l 'through pump |02 and line |36. The aluminum chloride catalyst settles from the liquid hydrocarbons in separater li'l from which it is withdrawn by pump |69 in line |00 and it is recycled to reactor itl through lines ll, lill and |36. Spent catalyst may be withdrawn from the cycle by means of valved line |50. The hydrocarbon mixture comprising butane, isobutane and the isomerized Cs-l-Ce fraction passes via valved line |55 to line l0 which leads to fractionator l2 from which the butanes are recycled to the alkylation step and the isomerized Ct+Ce fraction of approximately 82 octane number (CFR-M) passes therefrom as blending naphtha to be disposed of as described above. If desired the hydrocarbon stream in line |55 may be subjected to a hydrogen chloride absorbent or neutralizing agent such as solid sodium hydroxide or other neutralizing material in order to remove the last traces of hydrogen chloride before introduction of the stream to linel.

Referring now to the regeneration of the catalysts in reactors 28 and 58, these may be regenerated by oxidation of the carbonaceous material deposited thereby means of an oxygen containing mixture of gases such as air which may be diluted with nue gas to reduce the oxygen percentage of the gas. Although we have described catalyst chambers as single reactors, `these may be provided as multiple units to be operated in parallel banks, either continuously or intermittently, the catalyst in one or more of the chambers of the multiple units being regenerated with'- out interruption of the continuous process. Re-

actors |05 and |31 may also be in multiple. However, the regeneration of the catalyst in reactors 23 and 58 may be accomplished without the use of multiple chambers by a slight change in the mode of operation when it becomes neces- Ysary to regenerate either of these catalyst beds.

For example, when the catalyst in reactor 20 becomes spent we may isolate reactor 28 by closing valves 2l and 30 and valves |57 and |59 are opened'to admit and discharge inert gas from reactor 28 through lines l5@ and |58 to remove hydrocarbon vapors. After the reactor is purged the inert gas is gradually enriched 'with oxygen or air and the carbonaceous material is removed as an oxidation product. The reactor is again purged with inert gas, after which reactor 28 is in condition for use in the hydrocarbon conversion process. During the regeneration of the catalyst in reactor 20 all of the virgin feed may be fractionated in ractionator iii in order to obtain maximum isobutane in the process. The supply of lsobutane for alkylation during this catalyst regeneration period may be augmented by introducing isobutane from an external source via valved line il@ or by introducing normal butane or a normal butane-isobutane stream to butane isomerizer i3? via valved line 20@ and lines l2@ and i3d. As an alternate method or `operation during the regeneration of the catalyst in reactor 28 we may operate alkylator it as a polymerizer using diluted spent allrylation acid in the hot acid process vto polymerize the olefins produced in. reactor Reactor 605 is operated at temperatures within the range of 165 F. to 212 F., the acid strength being Within the range of 60% to 80% sulfuric acid when polymerizing oleiins. rl`he polymer may be separated from the acid in settler liti and separated from unreacted gases in fractionator M0, polymerof suitable boiling range being with- 'drawn via line |30 for blending in motor fuel ork of suitable boiling range for hydrogenation over suitable catalyst and under conditions well known in the art, the product to be blended in aviation gasoline.

When it becomes necessary to regenerate the catalyst in reactor the reactor may be isolated by closing valves 03 and 60 in lines 02 and 59, respectively, and inert gas and regenerating gas is admitted via valved line i to reactor 08 and exits via valved line |02, valves 46| and M53 belng open. During the regeneration of the catalyst in reactor 58 reactor 20 may be operated at temperatures Within the range of about 800 F. to 850 F. within which temperature range sufflcient condensible olens are produced along with lsobutane to alkylate the isobutane in reactor Ill. Any make-up hydrogen required for the isomerization reaction in reactor itl during the regeneration of catalyst bed 5&3 is furnished from storage via valved line 69, which leads `to lines G5 and 51 and thence to reactor |311. During the regeneration of the catalyst in reactor 00 the product from reactor 28 may be sent to fractionator 0|, valve 50 in line 49 being closed and verted in reactor 58.

. 2,861,611 bimng the recycleboaoms with fresh feed in line 22.

We may also operate reactors28 and 58 in parallel as well as in series as described above. For example, virgin naphtha of wide boiling range or the C1 and higher fraction of virgin naphtha may be delivered to lines 22 and 89 from line 2l simultaneously and mixed therein with recycle from fractionators 4I and 'l2 by lines 55 and 85 respectively, the aromatic bottoms eliminated from each fractionator beingdischarged to finished product line -57 or the products from reactors 28 and 58 may be sent to a common frac-1 tionation system by passing the product from reactor 28 through line 29 to reactor 58 product line 59 viavalves 42 and 45, valves 3i, 36, 40, 50

the optimum amounts of isobutane and condensible olens.

By means of our process a typical Mid-Continent straight-run naphtha boiling between 60 F. and 400 F. and having an octane number of 45 to 50 can be converted into high grade motor fuel` of 85 to 95 octane number in yields of 80 to 90% on a, volume basis. The Cs-i-C cut of such a naphtha will generally be about to by volume and will have an octane number of 65 to 70. This cut is converted, at least partially, to isobutane in reactor itl, while the residue is isomerized to an octane number of 80 to 85 on a butane-free basis.. The volume recovery of this material, including the isobutane, is 95% to 105% by volume.

The C1 and higher cut of the naphtha generally contains 15 to 25% aromatics having an average octane number of 100, to 40% naphthenes having an average octane number of from to 65 and 40 to 50%k parafilns having an average octane number of zero to l5. Under normal operating conditions about to 70% of the paraiiins are vconverted togases in reactor 28 while to 90% of the remaining paralns are con- As previously described these gases are principally Ca 4and C4 hydrocarbons and are converted in reactor |05 to an alkylate of to 95 octane number in yields of 65 to 75% on the basis of the paramns converted in the condensiblegases may be separated from the l conversionA product and recombined to form naphthas having highantiknock value which can be blended baci:V with the residue naphtha to make'motor fuels of greatly increased antilcnock value. Although we have described 'our invention in detail and, therefore. utilized certain speciil'c terms and language herein, itis to be understood that the present disclosure is illustrative rather than restrictive and that changes and modifications may be resorted to without departing from the spirit or the scope of the claims appended hereto.

We claim:

1. The process oi.' producing isobutane which comprises contacting a petroleum naphtha with a solid, silica-alumina cracking catalyst in a reaction zone at a temperature above 575 F. but below 850 F., at a pressure within the approximate range of about 15 to 65 pounds per square inch and at a low space velocity of the order of 0.01o to 1 volume -of liquid naphtha per hour per volume of catalyst space in the reaction zone, removing the products from said reaction zone, fractionating the products to obtain a fraction rich in isobutane and a fraction containing hydrocarbons of the naphtha boiling range and recycling at least a substantial part of the lastnamed fraction to said reaction zone for obtaining further production of isobutane.

2. The process of claim 1 wherein the petroleum naphtha is rich in straight-chain parailin hydrocarbons and wherein the space velocity is such that from about 5% to about 15% of the straightchain parailln hydrocarbons in said naphtha is converted to normally gaseous products per pass through said reaction zone.

reactors 28 and 58. The naphthenes are substantially converted to 4aromatics in reactor 58 so that the liquid product from the bottom of fractionator 12 will be highly aromatic and have an octane number of to 100. As hereinbefore described these high octane number products can be blended together or used separately for special purposes.

We have described a process whereby the low 3. The process oi simultaneously producing isobutane and a motor fuel fraction having a higher antiknock value from a low antiknock naphtha rich in normal parailln hydrocarbons which process comprises contacting at least a part of said low knock rating naphtha with a solid, silicaalumina cracking catalyst in a reaction zone at `a temperature above 575 F. but below 850 F.

under an absolute pressure within the approximate range of 15 to 65 pounds per square inch and at such space velocity within the approximate range of about .0l to 1 volume of liquid feed per hour per volume of catalyst space in the reaction zone that the conversion of normal parafns per pass to normally gaseous products is only about 5% to 15%, fractionating the products to obtain an isobutane fraction and a normally liquid fraction rich in parafnic hydrocarbons and of higher octane number than the original charge, recycling a substantial portion of the normally liquid fraction to said reaction zone and removing the remainder 'of said normally liquid

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