US3364132A - Hydrocarbon conversion process to produce gasoline from high boiling hydrocarbon oils by hydrocracking and reforming - Google Patents

Description

Jan. 16, 1968 R SCHOENFELD 3,364,132

HYDROCARBON CONVERSION PROCESS TO PRODUCE GASOLINE FROM HIGH BOILING HYDROCAHBON OILS BY IIYDROCRACKING AND REFORMING Filed Sept. 19, 1966 Q R m K N v, o Q n w "a m a M Fracf/onalar N N "a 4 m u v \1 I Q O "Q 4 Depentam'zer i J Q i Compressor R 0 E -1 g Reactor f i w a N N a N VE N T0,?-

:: E h g Jack R. Schaenfald Q? y. 4, W %m g 3 g 3 A. T [.0 RAIELSI.

United States Patent 3,364,132 HYDROCARBON CONVERSION PROCESS T0 PRO- DUCE GASOLINE FROM HIGH BOILING HY- DROCARBON OILS BY HYDROCRACKING AND REFORMING Jack R. Schoenfeld, Oak Park, Ill., assignor to Universal Oil Products Company, Des Plaines, lll., a corporation of Delaware Continuation-impart of application Ser. No. 384,112, July 21, 1964. This application Sept. 19, 1966, Ser. No. 589,155

4 Claims. (Cl. 208-60) This invention is a continuation-in-part of my copending application Ser. No. 384,112 filed July 21, 1964, now abandoned. This invention relates to the production of high octane gasoline. More specifically, this invention relates to the conversion of heavier hydrocarbons into lighter gasoline boiling range hydrocarbons having a high octane number without excess production of light hydrocarbons, such as butane. Still more specifically, this invention relates to a process for the conversion of heavy hydrocarbons into high octane gasoline which process comprises hydrocracking the heavy hydrocarbons into gasoline boiling range, reforming said gasoline boiling range components to a moderate octane number and contacting the normally liquid reformate with a molecular sieve sorbent selective for normal aliphatic hydrocarbons thereby producing a relatively less sorbed fraction having a high octane number.

In one of its embodiments this invention relates to a process for the production of high octane gasoline which comprise contacting a heavy feed stock with a hydrocracking catalyst in a reaction zone maintained at hydrocracking conditions; withdrawing a hydrocracking reactor effluent and separating said efiiuent to produce a gasoline boiling range fraction suitable for reforming; introducing said fraction into contact with a reforming catalyst in a reaction zone while maintaining said reforming zone at mild reforming operating conditions; withdrawing from the reforming unit a normally liquid reformate of moderate octane number; introducing said reformate into a sorption unit containing a solid molecular sieve sorbent selective for normal aliphatic hydrocarbons; withdrawing and recovering from said sorption unit a relatively less sorbed raffinate having a high octane number; and Withdrawing and recovering a selectively sorbed sorbate having a high concentration of normal aliphatic hydrocarbons as a separate effluent from said adsorbing unit.

Since the advent of the internal combustion engine, there has been a trend in the refinery industry to increase the yield of gasoline from crude oil. Various cracking processes have been discovered to accomplish this objective as is evidenced by the history of the petroleum process developments. Starting with thermal cracking, improvements were made to improve the selectivity of the cracking reactions which lead to the development of the wellknown cat crackers. Recently further improvements in the yield and selectivity have been made by the hydrocracking process. Unlike the prior processes, hydrocracking means cracking hydrocarbon molecules in the presence of excess hydrogen such that there occurs a rapid hydrogenation of the broken carbon to carbon bond resulting in enhanced yields and selectivity. One of the problems of this new hydrocracking process is the excessive production of butane. A certain quantity of butane is desirable in finished gasoline for vapor pressure purposes. Hydrocracked gasolines generally have too low an octane number to be useful directly in todays internal combustion engines and usually these hydrocracked gasolines are reformed to upgrade the octane number. The reforming step also produced butane and the combined butane from the hydrocracking and reforming steps generally 3,364,132 Patented Jan. 16, 1968 exceeds the required butane for vapor pressure purposes in the finished gasoline. This excess butane is only useful as a direct burning fuel such as in L.P.G. and accordingly has a lower dollar value. Other light hydrocarbons such as methane, ethane and propane are also too light and volatile to be used in gasoline and accordingly have limited use as direct burning fuels.

It is known that normal paraflins in the gasoline boiling range have low octane numbers. If the normal paraffin components of gasoline are extracted from the gasoline, there would be a substantial increase in octane number. For example, normal hexane has an F-l clear octane number of 24.8, normal heptane has an F-l clear octane number of 0.0 and normal octane even lower. Todays engines need gasolines having at least an octane number of and as much as and any appreciable concentration of normal paraflins in the gasoline results in a substantial depression of the octane number.

Reforming of gasoline helps to reduce the concentration of normal paraffins in most gasolines by reactions such as isomerization, hydrocracking and dehydrocyclization. In the isomerization reaction, if there is an excess concentration over equilibrium of normal parafiins, the normals will be isomerized to their more highly branched isomers with a resulting increase in octane number. For example, converting normal hexane to 2-methyl pentane increases the F-1 clear octane number to {73.4. In the hydrocracking reaction the higher molecular weight molecules are cracked into lighter components with an increase in octane number although there is a decrease in yield. For example, if normal heptane is converted into normal pentane and ethane, the resulting normal pentane has an F-l clear octane number of 61.8 while the ethane would be removed from the gasoline since it is too volatile a component. This reforming hydrocracking reaction has the undesirable characteristics of causing a yield loss and consuming hydrogen. Generally, as a reforming catalyst reactivates, the temperatures of the reforming reactor are increased to maintain a constant octane number which has the effect of increasing the amount of hydrocracking that occurs. In the dehydrocyclization reaction a normal parafiin would be converted to an aromatic molecule with an evolution of four moles of hydrogen and a substantial increase in octane number. For example, normal heptane might be converted into toluene having an F-l clear octane number of 119.7. Unfortunately, this latter reaction is slow and occurs only to a limited ex tent, if at all.

Thus, looking at the overall effect of the reforming reactor, the hydrocracking reaction results in the production of light hydrocarbons such as ethane, propane and butane while increasing the octane number, said light hydrocarbons being an undesirable product. It would be preferable to minimize the reforming hydrocracking reaction consistent with achieving the desired octane number.

Recently separation processes selective for normal aliphatic hydrocarbons have been developed employing molecular sieve zeolites having pore sizes of about 5 Angstrom units. Such a separation process can be com bined with a hydrocracking process and a reforming process to selectively extract the low octane number normal parafiins leaving a gasoline of enhanced octane number with the hydrocracker removing material such as sulfur and nitrogen to maintain sieve stability and reforming catalyst stability. This will permit the reforming reactor to be operated at milder conditions, i.e. lower temperatures, to maintain the same octane number While minimizing the undesirable reforming hydrocracking reaction.

It is accordingly an object of this invention to convert heavy hydrocarbons to high octane gasoline while minimizing the formation of light hydrocarbons selected from the group consisting of methane, ethane, propane and butane.

It is another object of this invention to obviate the above-mentioned difficulties in producing high octane gasoline from heavy hydrocarbons.

It is a further object of this invention to disclose a combination hydrocracking-reforming-separation process to efiiciently produce high octane gasoline.

It is a more specific object of this invention to minimize the production of butane when converting heavy hydrocarbons boiling above the gasoline boiling range into high octane gasoline.

It is another more specific object of this invention to disclose a combination hydrocracking-reforming-separation process to produce high octane gasoline While maintaining butane balancethat is, only to produce sufficient butane to satisfy the vapor pressure requirement of the gasoline.

It is still another more specific object of this invention to produce high octane gasoline from a once through reformer maintained at mild reforming conditions.

It is a further more specific object of this invention to produce a stream of high purity normal paraflins in the gasoline boiling range suitable for use as a high quality fuel, in the production of polymers and plasticizers and as an intermediate in the production of chemicals.

The hydrocracking step of the process is carried out by passing the heavy hydrocarbon charge stock over a fixed bed of hydrocracking catalyst in the presence of hydrogen at elevated temperatures and pressure. A preferable hydrocracking catalyst is a refractory oxide support comprising silica and alumina and having incorporated thereon a metal having appreciable hydrogenation activity. Suitable supports comprise amorphous silica-alumina, crystalline aluminosilicates such as faujasite and mordenite (preferably in the hydrogen form), etc. Said metal comprises platinum, nickel, cobalt, molybdenum, iron, tungsten, palladium etc. The amorphous refractory oxide support 'may be prepared by oil dropping techniques to form spherical particles or may be prepared in pellet form using pilling machines. The crystalline aluminosilicate may be prepared by crystallization from suitable basic solutions containing sodium cations and anions selected from the group consisting of silicate, aluminate and hydroxyl. The metal may then be impregnated or ion-exchanged upon the refractory oxide as a metal salt and activated by oxidation. The finished catalyst is loaded into a fixed bed reactor and the heavy hydrocarbon is passed through said fixed bed. The reactor is maintained at temperatures of from about 550 F. to about 900 F., depending upon the charge stock, the reactor pressure, the space velocity'and the desired degree of conversion. The pressure should be maintained at least as high as 500 p.s.i.g. and preferably from 1,500 p.s.i.g. to 2,000 p.s.i.g. It is necessary to maintain excess hydrogen in the reactor such that when a carbon to carbon bond is broken, hydrogen is available to react with the broken bond. This may be achieved by recycling separator gas which is rich in hydrogen and supplying additional makeup hydrogen to satisfy the hydrogen consumption. Gas rates of at least 2,500 standard cubic feet per barrel of heavy charge (s.c.f./bbl.) and preferably 5,000 to 50,000 s.c.f./bbl. should be supplied to the reactor. Liquid hourly space velocities of from about to about 0.5 may be employed. The reactor effluent is separated such that a fraction suitable for reforming is produced. Such fractions preferentially have as the lightest component at least six carbon atom molecules and a maximum Engler distillation end point of about 400 F. It is preferable to employ a two stage hydrocracking process, the first stage substantially converting organic nitrogen into ammonia and organic sulfur into hydrogen sulfide and the second stage substantially hydrocracking the hydrocarbons to produce a gasoline boiling range material. Preferable first stage catalysts comprise a Group VI and a Group VIII metal on amorphous silica-alumina supports such as nickel-molybdenum or nickel-tungsten on silica-alumina. Preferable second stage catalysts comprise a Group VIII metal on crystalline alumina-silicate such as nickel, palladium or platinum on faujasite or mordenite.

Suitable charge stocks for the hydrocracking step comprise heavy hydrocarbon streams A preferable stock would be gas oil having an initial boiling point of at least 400 F. and having an end point as high as about 900 F. to about 1000" F. Cycle oils from catalytic crackers and cokers would be another preferable charge stock. Middle distillates and even kerosenes can also be used as charge stocks. Of course, mixtures of the above may also be used.

The reforming step of the process is carried out by passing the gasoline boiling range fraction of the hydrocracking reactor effluent over a fixed bed of reforming catalyst in the presence of hydrogen at elevated temperature and pressure. Preferred reforming catalysts comprise an alumina support having platinum and combined halogen, especially combined fluorine and/ or combined chlorine deposited thereon. The alumina-combined halogen particles may be prepared by oil dropping techniques to form spherical particles and thereupon drying and oxidizing the particles. The platinum is thereafter impregnated upon the particles as a metal salt and activated by oxidation. The finished catalyst is loaded into a fixed bed re actor and the gasoline boiling range feed is passed through said fixed bed. Reformers are conventionally maintained at temperatures of from about 850 F. to about 1050 F depending upon the charge stock, the reactor pressure, the space velocity and the desired octane number. Typical reforming pressures vary from 500 p.s.i.g. to 200 p.s.i.g., depending primarily upon the charge stock. It is necessary to maintain excess hydrogen in the reactor to avoid the formation of undesirable side products such as carbon aceous deposits. Said excess hydrogen may be supplied by recycling separator gas which is rich in hydrogen. Gas rates sufiicient to give a hydrogen to charge ratio of at least 2 and preferably at least 3 up to about 15 should be supplied to the reactor. Liquid hourly space velocities generally run from 5.0 to 0.5, depending primarily upon the charge stock and the desired octane number. The normally liquid reforming reactor eflluent, called the reformate, is separated from the total effluent and is charged to the separation process step.

Part of the essence of this invention involves the opera tion of the reforming reactor at mild conditions to produce a reformate of moderate octane number and then to extract the normal paraflins out of the reformate thereby increasing the octane number of the relatively less sorbed raffinate stream hereinafter described. By mild conditions I mean primarily reducing the temperature of the reactor to range of from about 800 F. to about 980 F. In certain cases it might also be possible to increase the space velocity in the reforming reactor to a minimum of 1.0 and as much as 5.0. In other cases lower pressures are contemplated as part of the mild conditions.

A preferable separation apparatus is shown in U.S. Letters Patent No. 2,985,589, issued on May 23, 1961. The separation step of this invention employs a fixed bed of solid sorbent selective for normal parafiins and preferably is operated in a simulated countercurrent flow of solid and liquid processing scheme. The simulated countercurrent flow is achieved by means of a rotary valve such as is described and claimed in U.S. Letters Patent No. 3,040,-' 777, issued on July 26, 1962. One of the essential parts of the separation step is the sorbent contacting chamber shown in the figure. Said chamber is capable of having introduced to it continuously the reformate feed and a desorbing fluid while simultaneously having Withdrawn a relatively less sorbed high octane raifinate and a sorbent rich in normal paraflins. This separation step may be carried out by introducing said reformate into a first zone of a fixed bed of solid sorbent, selective for normal aliphatic hydrocarbons, containing at least four serially interconnected zones having fluid flow connecting means between adjacent zones and between the outlet of one terminal zone and the inlet of the other terminal zone in the series to thereby provide cyclic fluid flow in said process, substantially simultaneously withdrawing relatively less sorbed raflinate having a high octane number from a second zone immediately downstream of said first zone, substantially simultaneously introducing a desorbing fluid into a third zone immediately downstream of said second zone, substantially simultaneously withdrawing resulting sorbate comprising selectively sorbed normal aliphatic hydrocarbons from a fourth zone immediately downstream of said third zone, continuously circulating a stream of fluid through said series of interconnected zones, and periodically advancing downstream the point in said fixed bed of introducing said reformate while simultaneously and equally advancing downstream the point of introducing desorbent and withdrawing raflimate and sorbate.

The sorbent contacting chamber is operated at conditions of temperature, pressure and under other process conditions which depend upon the particular feed stock involved and the required purity of product. Although this chamber may be operated either in the liquid phase or vapor phase, it is preferable to operate in the liquid phase. Typical liquid phase operation is, for example, temperatures of from 30 F. to 500 F. and more preferably from 100 F. to 380 F., and pressures of from slightly superatmospheric to 30 atmospheres or even higher. Generally higher pressures will be employed for lower molecular weight feed stocks or desorbents to maintain liquid phase in the contacting chamber. The maximum charge rate of feed stock through the fixed bed of solid sorbent is limited by the tolerable pressure drop through said fixed bed and the rates of sorption. The minimum charge rate of feed stock through said fixed bed is limited to a rate suflicient to avoid back mixing (i.e., to maintain substantially plug flow through said beds). It is expected that liquid hourly space velocities (on solid sorbent) of from about 0.01 to about 20 will be employed depending upon the operating conditions of pressure and temperature, the feed stock and the equipment limitations. In order to maintain sieve stability it is preferable that polar compounds such as sulfur, nitrogen and oxygen compounds be removed from the feed to the contacting cham ber. The hydrocracking step of this invention is satisfactory to accomplish this result.

Suitable sorbents would be any substance which can be produced in discrete particles within the size range of from about to about 200 mesh and which have an appreciable degree of selectivity for normal parafiins. One suitable sorbent is dehydrated crystalline metallic aluminosilicate, commonly called molecular sieves. These molecular sieves are made up of a crystalline structure having many small cavities connected by still smaller pores of uniform size. Although molecular sieves may be made in pore sizes of from 3 Angstrom units up to 12 or or even more Angstrom units, the pore size should be about 5 Angstrom units in order to selectively sorb normal parafiins. One preferable molecular sieve is the so-called Type A (calcium form) sieve described in US. Patent 2,882,243.

Suitable desorbents comprise normally liquid hydrocarbons having a boiling point outside of the boiling point range of the feed to the contacting chamber and an appreciable concentration of normal aliphatic hydrocarbons (at least The desorbent may then be easily removed from the ratfinate and the sorbate by ordinary simple fractionation and is generally recycled back to the sorbent contacting chamber.

One preferable embodiment is shown in the accompanying figure. It should be recognized that there are numerous variations in the basic flow scheme shown and it is not intended to limit the scope of this invention to the shown arrangement in the figure. For example, reformate fractionator 23 could precede depentanizer 12 or depentanizer 12 could be debutanizer or a dehexanizer; or the hydrocracking reactor may be operated at suflicient severity to not produce any heavy product, thus obviating the need for fractionator 23, or a two stage hydrocracker as described hereinbefore can be employed. Also for simplification of the drawing and of the description, conventional equipment such as control instruments for observing and controlling temperatures, pressures, flow rates and liquid levels and the like, pumps, valves, and various heat exchangers are not indicated or described specifically. Further, methods of economically and efliciently using the heated and cooled streams are not shown. This conventional equipment is within the skill of an engineer and is omit-ted to avoid unduly describing incidental details to the main invention.

The heavy hydrocarbon feed stock is introduced into flow conduit 1 where it joins flow conduit 4 carrying recycled separator gas and makeup hydrogen. The mixture flows through flow conduit 5 and enters hydrocracking reactor 6 containing hydrocracking catalyst. The reactor eflluent leaves through flow conduit 7 and enters high pressure separator 8. The normally gaseous effiuent is drawn through flow conduit 9 by means of recycle compressor 10, through flow conduit 3 where it joins the makeup hydrogen flowing in flow conduit 2 and the mixture moves into flow conduit 4. The normally liquid efflucnt leaves separator 8 where it flows through flow conduit 11 and into depentanizer 12. It is normally preferable to depentanize prior to reforming since the reforming reactions can do very little to improve octane number of the isomers of pentane and cyclization of pentane seldom occurs. The overhead fraction having pentane as its heaviest component is removed through flow conduit 13 where it enters overhead receiver 14. This overhead stream may be subject to partial condensation and the light components such as hydrogen, methane, ethane, etc., and contaminants such as hydrogen sulfide, ammonia, etc., are removed through flow conduit 17. A portion of the condensed overhead is returned to fractionator 12 as reflux by means of flow conduit 15. The net liquid overhead comprising pentane and some butane is removed through flow conduit 16. The bottoms fraction having hexane as its lightest component is removed through flow conduit 18 where a portion of the bottoms passes through flow conduit 19, heater 20 and returns to fractionator 12 by means of flow conduit 21. The net bottoms is withdrawn through flow conduit 22 where it flows into fractionator 23.

The overhead fraction from fractionator 23 is removed through flow conduit 24 where it flows into overhead receiver 25. This overhead stream comprises a material within the gasoline boiling range having hexane as its lightest component in any significant concentration and having an Engler distillation end point no higher than about 400 F. A portion of the overhead is returned to fractionator 23 by means of flow conduit 26 as reflux, while the net overhead stream is removed through flow conduit 27 where it is sent to reforming reactor 35. The bottoms fraction of fractionator 23 is removed through flow conduit 28 where a potion flows through flow conduit 29, heater 30 and returns to fractionator 23 by means of flow conduit 31. The net bottoms is removed through flow conduit 32 and comprises a hydrocarbon fraction suitable for use as a kerosene or middle distillate. If desired, this stream can be recycled to hydrocracking reactor 6 so as to increase the gasoline yield or it may be used directly as a product. This heavy product has many desirable burning characteristics such as a high smoke point and comprises a superior premium heavy fuel. The ratio of the heavy product to the gasoline range product is determined by the hydrocracking operating conditions of pressure, temperature, liquid hourly space velocity,

7 hydrogen to oil ratio and the charge stock. Generally any ratio may be produced depending upon such economic factors as time of season, marketing demands and type of climate.

The gasoline boiling range fraction flowing in flow conduit 27 joins recycled separator gas flowing in flow conduit 33 and the mixture flows into reforming reactor 35 by means of flow conduit 34. The oil and gas contact the reforming catalyst in reactor 35 and the effluent is withdrawn through flow conduit 36 and into high pressure separator 37. The normally gaseous eflluent is withdrawn from conduit 38 where a portion is recycled through flow conduit 39, recycle compressor 81 and into flow conduit 33 where it ultimately returns back to reactor 35. Since the reactions present in a reformer produce hydrogen and other light hydrocarbon gases, the net gas is removed through flow conduit 40. The normally liquid eflluent is withdrawn through conduit 41 whereupon it enters stabilizer 42.

An overhead fraction is removed from stabilizer 42 through flow conduit 43 where it enters overhead receiver 44. A portion of the overhead is returned to stabilizer 42 as reflux by means of flow conduit 45 while the remaining portion is withdrawn through flow conduit 46. The stabilizer may be operated to remove butane and remaining lighter components or it may also be utilized to remove pentanes in addition to the butane and lighter components. The decision would depend primarily upon whether normal pentane is desired in the finished gasoline. Normally, the stabilizer is operated as a debutanizer but since the separation step will extract the normal pentane out of the gasoline fraction, it is preferable to remove all the pentanes overhead in the stabilizer and recombine the pentanes with the finished high octane gasoline. The bottoms fraction comprising the depentanized reformate is removed through flow conduit 47 where a portion flows through fiow conduit 48, heater 49 and returns to stabilizer 42 by means of flow conduit 50. The net bottoms fraction is withdrawn through flow conduit 51 where it continuously passes through rotary valve 79 and enters sorbent contacting chamber 80 at the upstream point of zone I.

Desorbent, from a source hereinafter described, continually flows through flow conduit 78 through rotary valve 79 and enters sorbent contacting chamber 80 at the upstream point of zone III. Relatively less sorbed raffinate is continuously withdrawn from the downstream point of zone I (also the upstream point of zone 1V) through flow conduit 56, through rotary valve 79 and finally enters rafiinate fractionator 57. An overhead fraction is removed from fractionator 57 through flow conduit 58 Where it enters overhead receiver 59. A portion of the overhead fraction is returned to fractionator 57 as reflux by means of flow conduit 60. The remaining overhead portion is withdrawn through flow conduit 61 where it subsequently enters flow conduit 78 and comprises a portion of the desorbent. The bottoms fraction is removed from fractionator 57 through flow conduit 62 where a portion flows through flow conduit 63 and heater 64 and returns to fractionator 57 by means of flow conduit 65. The net bottoms stream is withdrawn through flow conduit 66 and consists of high octane gasoline.

Selectively sorbed sorbate is continuously withdrawn from sorbent contacting chamber 80 at the downstream point of zone III (also the upstream point of zone 11) where it enters flow conduit 67, passes through rotary valve 79 and finally enters sorbate fractionator 68. An overhead fraction is removed from fractionator 68 through flow conduit 69 where it enters overhead receiver 70. A portion of the overhead fraction is returned to fractionator 68 as reflux by means of flow conduit 71. The remaining portion is withdrawn through flow conduit 72 where it subsequently enters flow conduit 78 and comprises the remaining portion of the desorbent. The bottoms fraction is removed from fractionator 68 through flow conduit 73 where a portion flows through flow conduit 74, heater 75 and returns to fractionator 68 by means of flow conduit 76. The net bottoms stream is withdrawn through flow conduit 77 and comprises normal paraflins in the gasoline boiling range. This particular flow scheme employs a desorbent of lower boiling range than the gasoline, It should be observed that a desorbent heavier than gasoline could also be employed in which case the recovered products would come overhead from these fractionators.

A continuous stream of circulating fluid is maintained in sorbent contacting chamber 80 by means of flow conduits 52, 53 and and pump 54. The points of introduction of feed and desorbent and the withdrawal of raffinate and sorbate are periodically shifted downstream by means of rotary valve 79. It should be observed that the position of the zones will shift as the points of introduction and withdrawal are shifted since the zones are defined from the point of introduction and withdrawal of the various streams. It is the continuous periodic shifting of the inlet and outlet streams to different points in the sorbent contacting chamber which results in the simulated countercurrent flow of solid sorbent and liquid.

The recovered normal paraflins could be recycled to reforming reactor 35, thereby isomerizing, hydrocracking and dehydrocyclizing the normal parafiins to high octane material giving enhanced gasoline yield, or the normal paraflins may be recovered as a useful product. Since one of the principal objects of this invention is to avoid excess production of butane, preferably the normal paraflins are not recycled unless additional butane is required to satisfy the vapor pressure requirement of the finished gasoline. Normal paraffins have excellent buming qualities and are also useful as an intermediate in the production of chemicals, polymers and plasticizers.

Example I the reactor variable conditions were operated as follows:

catalyst peak temperature, 825 F.; pressure, 2000 p.s.i.g.;

I liquid hourly space velocity, 0.5; and hydrogen circulation rate, 50,900 s.c.f./bbl. A material balance was run around this step of the process and resulted in the follow ing yield breakdown: 53.5 s.c.f. of methane/bbl. of charge, 44.5 s.c.f. ethane/bbl., 78 s.c.f. propane/bbl., 7.8 liquid volume percent butane of feed, 5.6 liquid volume percent pentane, 9.8 liquid volume percent hexane and 95.3 liquid volume percent 0 fraction. It is estimated that the Engler distillation end point of the 0 fraction was about 580 F. The normally liquid efiluent was withdrawn through flow conduit 11 where it was fractionated in fractionators 12 and 23 and produced a gasoline boiling range material flowing in flow conduit 27 having the following properties: API gravity, 55.3; Engler initial boiling point, 190 F.; Engler end point, 400 F.; paraflin content, 45 volume percent; naphthene content, 49 volume percent; and aromatic content, 6 volume per cent.

Reforming reactor 35 was loaded with cc. of reforming catalyst Weighing 51.3 grams. The above described gasoline boiling range material passing through flow conduit 27 was introduced into reforming reactor 35 at a rate of 198 cc. per hour (measured at 60 F.) while the reactor viable conditions were operated as follows: average catalyst temperature, 883 F.; pressure, 400 p.s.i.g.; liquid hourly space velocity, 1.98; and hydrogen to oil mole ratio, 10.7. A material balance was around this step of the process and resulted in the following yield breakdown; 930 s.c.f. of hydrogen per bbl. of reformer feed, 80 s.c.f. methane/bbl., 45 s.c.f. ethane/bbL, 60 s.c.f. propane/bbl., 4.1 liquid volume percent butane of reformer feed, 3.7 liquid volume percent pentane and 82.5 liquid volume percent C fraction. It is estimated that the octane number of the C reformate is 88.0 F-l clear and the C fraction is 88.3 F-l clear. The normally liquid reactor efiiuent is introduced into stabilizer 42 which is operated as a depentanizer. The C fraction is withdrawn through flow conduit 51 whereupon it is introduced into sorbent contacting chamber 80.

The desorbent flowing in flow conduit 78 comprises normal butane. The high octane gasoline is withdrawn in flow conduit 66. It is estimated that the F-1 clear octane number has increased from 88.3 to 96.3 while the C gasoline yield has decreased from 82.5 to 75.6 liquid volume percent as a result of processing through the sorbent contacting chamber.

It is further estimated that if the reformer severity had been increased to directly achieve the F-1 clear 96.3 octane number on the 0 fraction, typical yields would have been as follows: 78 s.c.f. ethane/bbL, 84 s.c.f. propane/bbl., 6.0 liquid volume percent butane and 75.2 liquid volume percent (3 fraction. It should be noted that the production of light hydrocarbons such as ethane, propane and butane has increased While the overall yield of gasoline has remained relatively constant. This means in the overall results that instead of producing light hydrocarbons having a low dollar value, a stream of high purity normal parafiins has been obtained having a much higher value by operating the reforming reactor at mild operating conditions.

I claim:

1. A combination process for the production of high octane gasoline from a heavy hydrocarbon feed stock accompanied by minimum butane production which comprises:

(a) introducing said heavy hydrocarbon feed stock into contact with a hydrocracking catalyst in the presence of hydrogen in a hydrocracking reaction zone;

(b) withdrawing a hydrocracking reactor efiiuent from said reaction zone and separating said efiluent by fractionation to produce a gasoline fraction suitable for reforming;

(c) introducing said fraction into contact with a reforming catalyst in the presence of hydrogen, and in the absence of recycled normal aliphatic hydro- 10 carbons, in a reforming reaction zone while maintaining said reforming zone at mild reforming conditions including a temperature of 800-980 F.;

(d) withdrawing from said reforming reaction zone a normally liquid reformate of moderate octane number;

(e) simultaneously introducing said reformate and a desorbent into a sorption unit containing a fixed bed of molecular sieve adsorbent having uniform pore openings of about 5 Angstroms while maintaining said sorption unit at a temperature of from F. to about 380 F. and under a pressure sufficient to maintain liquid phase conditions therein;

(f) withdrawing from said sorption unit a first stream consisting of relatively less sorbed rafiinate having a high octane number in admixture with said desorbent;

(g) simultaneously withdrawing from said sorption unit a second stream consisting of selectively sorbed sorbate having a high concentration of normal aliphatic hydrocarbons in admixture with said desorbent; and

(h) separately subjecting said first and second streams to fractionation to recover a high octane gasoline stream as one product and a normal aliphatic hydrocarbon-rich stream as another product.

2. The process of claim 1 further characterized in that said reforming catalyst comprises alumina, platinum and a halogen selected from the group consisting of chlorine and fluorine combined therewith, and said sorbent is Type A zeolite.

3. The process of claim 2 further characterized in that the hydrocracking reaction zone contains two zones, the first zone containing a Group VI and a Group VIII metal on a silica-alumina catalyst and the second zone containing a Group VIII metal on a crystalline aluminosilicate catalyst.

4. The process of claim 1 further characterized in that the desorbent is a normaHy liquid hydrocarbon having an appreciable amount of normal aliphatic hydrocarbons and a boiling range outside of the gasoline boiling range.

References Cited UNITED STATES PATENTS 3,008,895 11/1961 Hansford et al. 208-68 3,081,255 3/1963 Hess et a1. 208--88 3,159,564 12/1964 Kelley et a1. 20859 ABRAHAM RIMENS, Primary Examiner.

Claims (1)

1. A COMBINATION PROCESS FOR THE PRODUCTION OF HIGH OCTANE GASOLINE FROM A HEAVY HYDROCARBON FEED STOCK ACCOMPANIED BY MINIMUM BUTANE PRODUCTION WHICH COMPRISES: (A) INTRODUCING SAID HEAVY HYDROCARBON FEED STOCK INTO CONTACT WITH A HYDROCRACKING CATALYST IN THE PRESENCE OF HYDROGEN IN A HYDROCRACKING REACTION ZONE; (B) WITHDRAWING A HYDROCRACKING REACTOR EFFLUENT FROM SAID REACTION ZONE AND SEPARATING SAID EFFLUENT BY FRACTIONATION TO PRODUCE A GASOLINE FRACTION SUITABLE FOR REFORMING; (C) INTRODUCING SAID FRACTION INTO CONTACT WITH A REFORMING CATALYST IN THE PRESENCE OF HYDROGEN, AND IN THE ABSENCE OF RECYLED NORMAL ALIPHATIC HYDROCARBONS, IN A REFORMING REACTION ZONE WHILE MAINTAINING SAID REFORMING ZONE AT MILD REFORMING CONDITIONS INCLUDING A TEMPERATURE OF 800-980*F.; (D) WITHDRAWING FROM SAID REFORMING REACTION ZONE A NORMALLY LIQUID REFORMATE OF MODERATE OCTANE NUMBER;

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