US2917450A

US2917450A - Petroleum refining

Info

Publication number: US2917450A
Application number: US699781A
Authority: US
Inventors: Robert J Hengstebeck
Original assignee: Standard Oil Co
Current assignee: Standard Oil Co
Priority date: 1957-11-29
Filing date: 1957-11-29
Publication date: 1959-12-15
Anticipated expiration: 1976-12-15

Description

Dec- 15, 1959 R. J. HENGsTEBEcK 2,917,450

PETROLEUM REFINING Filed Nov. 29. 1957 INVNTOR (ngsfebeck ,m

ATTORNEY Robert J. Hel"y PETROLEUM REFINING Robert J. Hengsteheck, Valparaiso, Ind., assignor to ntndard Oil Company, Chicago, Ill., a corporation of lana Application November 29, 1957, Serial No. 699,781

Claims. (Cl. 208-66) This invention relates to a process for convening gas oil to naphtha of high octane, and it particularly concerns a series of integrated processing steps to achieve this purpose.

The widespread use of naphtha hydroforming using platinum-type catalysts has provided petroleum reners withy substantial amounts of hydrogen. These amounts of hydrogen are generally insulicient to satisfy the reliners need for hydrogen in more recently proposed processes for the hydrogen treating of various refinery oil streams.

An object of the present invention is to provide a refining process for converting gas oil to high octane naphtha in substantial yields and at the same time producing hydrogen for use within the process or for use within other processes. Another object is to provide an integrated sequence of processing steps by which gas oil is converted to naphtha of higher than ordinary octane number. A further object is to provide an integrated process for converting gas oil to high octane naphtha which minimizes the conversion of oil to coke and gas and thereby obtains high yields of naphtha per barrel of gas oil charge. Other objects and advantages of the present invention will be apparent from the detailed description thereof.

In accordance with the present invention a gas oil is dehydrogenated in the presence of a hydrogenation-dehydrogenation catalyst at temperatures in the range of 750 to 950 F. and in the presence of recycled hydrogen. Substantial amounts of hydrogen are produced. The gas oil is rendered more unsaturated and minor amounts of high octane naphtha are formed. This unsaturated oil, preferably after the removal of naphtha, is then subjected to mild catalytic cracking. The conditions employed in the cracking step are such that no more than about 30 to 40% of the dehydrogenated gas oil charge is converted into naphtha and other products which boil below 400 F. A catalytic cycle oil is recovered from the products of the mild cracking step and is subjected to solvent extraction with a solvent which preferentially extracts aromatic hydrocarbons therefrom. Thereby, an aromatics-rich hydrocarbon extract oil is separated from an aromatics-lean hydrocarbon raliinate oil. The raflinate oil is subjected to severe catalytic cracking so that at least 60% of the charged oil is conveited into naphtha and other hydrocarbons which boil below 400 F. The aromatics-rich extract oil is subjected to destructive hydrogenation by contacting it With a catalyst having hydrogenation and cracking properties in the presence of 2000 to 10,000 s.c.f. of hydrogen/barrel of extract oil While employing a temperature of 800 to l000 F. and a pressure of 1000 to 5000 p.s.i.g. Hydrogen produced in the dehydrogenation step is employed as a source of hydrogen in the destructive hydrogenation step. Naphtha of very high octane number is produced together with some incompletely converted but partially hydrogenated oil boiling above 400 F. This partially hydrogenated oil may be recycled to the destructive hydrogenation step and/or it may be passed as part of the charge oil to the solvent extraction Step. The catalytic cycle oil produced in the severe catalytic cracking step may be recycled thereto or preferably it is passed as part of the charge oil to the solvent extraction step.

At least a major portion of the gas oil initially charged to the process should be virgin gas oil which preferably has an end boiling point of 550 to 650 F. The higher boiling portion of the gas oil may be charged directly to the severe catalytic cracking Step. When the initial gas oil charge to the dehydrogenation step has a high end boiling point of 800 F. or higher or an ASTM 50% point of 600 F. or higher, then the partially hydrogenated oil boiling'above 400 F. which is produced in the destructive hydrogenation step may desirably be passed as charge oil to the severe catalytic cracking step.

In the dehydrogenation step, substantial amounts of hydrogen are produced which are necessary in the later destructive hydrogenation step. By restricting the charge gas oil to the dehydrogenation step to a virgin gas oil hav ing an end boiling point of about 550 to 650 F., coke formation upon the dehydrogenation catalyst and deactivation of the catalyst is reduced. By carrying out the mild catalytic cracking step under conditions such that no more than 30 to 40% conversion is obtained, coking upon the cracking catalyst and deactivation thereof are minimized. Because of the previous dehydrogenation step, the naphtha produced from mild catalytic cracking is of higher than ordinary octane number. This is partly due to its increased olefinic content and the more aromatic nature of the naphtha. Solvent extraction of the cycle oil from the mild catalytic cracking step segregates the cycle oil into fractions which are optimum for the subsequent processing steps. The rainate oil is catalytically cracked under severe conditions without substantial coke formation upon the catalyst, since substantial amounts of aromatics have been removed during the extraction step. The destructive hydrogenation step produces very high octane gasoline on the order of .F-l or higher from the aromatics-rich extract phase, whereas the octane number would be substantially lower if extraction had not been carried out. By recycling the catalytic cycle oil from the severe catalytic cracking step as charge to the solvent extraction step, and by recycling as charge to the extraction step the partially hydrogenated oil boiling above 400 F. which iS obtained from the products of the destructive hydrogenation step, these recycle streams are again segregated into optimum streams for catalytic cracking and destructive hydrogenation. When the initial charge to the dehydrogenation step of the process is a high boiling gas oil, e.g. having an ASTM 50% point of 600 F. or higher, the partially hydrogenated oil boiling above 400 F. which is recovered from the destructive hydrogenation step may be passed to the severe catalytic cracking step because higher octane gasoline, e.g. 95 F-l would be obtained therefrom than would be obtained by destructive hydrogenation thereof. Because it is partially hydrogenated it is improved as a charge to catalytic cracking. Also, if it were destmctively hydrogenated then the conversion would be much less than the conversion of a lower boiling fraction and the octane number produced from this high boiling fraction would only be about 88-90 F-l. It is is apparent that many benefits are provided by integration of the particular sequence of refining steps.

The gure shows in diagrammatic form an embodiment of the present invention whereby gas oil is converted to high octane naphtha fractions in high yield. Numerous pumps, heaters, coolers, and other incidental features which are apparent to those skilled in the art have been omitted from this figure to `are dehydrogenated in part.

v kieselguhr, fullers earth or the like. cobalt oxide-molybdenum oxide-alumina catalyst contain- 3 obtain improved clarity of the fundamental features of the invention.

In this embodiment a crude oil is charged from source l11 by way of line-12 into fractionation system represented herein by vessel 13 .Naphthaand lower boiling hydrocarbons are removed This virgin oil is fractionated.

from the system by way of line 14 and passed to refining means not shown herein. Residuum is removed by way of line 16. The gas oil is fractionated at a cut point of about 600 F. The higher boiling portion is removed from fractionating means 13 by way of line 17 and is processed by means which will be discussed later. The lower boiling fraction of the gas oil such as boils substantially within the range of 400 or 450 F. to about 600 or 650 F. is removed from fractionating means 13 and passed by way of line 18 together with recycled hydrogen through furnace 19 wherein the mixture is heated to the temperatures employed in dehydrogenation.

. of groups 6a and/or 8 of the periodic table carried upon a support such as alumina, bauxite, or clays such as For example, a

ing about 3% cobalt oxide and 9% molybdenum oxide may suitably be used. 'Ihe dehydrogenation reaction is carried out in the presence of hydrogen, usually in the amount of from `1000 to 4000 s.c.f./barrel of oil charged in order to suppress the formation of coke upon the catalyst. Conditions of temperature and pressure which are conducive to dehydrogenation are employed. Such temperatures may be from 750 to 950 F. and pressures of from 200 to 800 p.s.i.g. Space velocities of from about Y 0.2 to volumes of oil/hour/volume of catalyst may be employed.

In the embodiment described herein a series of three dehydrogenation reactors designated as

reactors

22, 23,

and` 24 are employed. The particular arrangement enables the maintenance of a more nearly constant temperature within each dehydrogenation reactor.

Because the reaction is endothermic, the arrangement shown in this embodiment prevents Wide variation in reaction temperatures Within the reactors and enables operation at more nearly the optimum temperature. The heated light gas -oil and hydrogen are passed by way of line 21 from furnace 19 into dehydrogenation reactor 22. Additional hydrogen may be recycled from the hydrogen separator and passed into line 27. Separate means for preheating hydrogen to the reactor temperatures may be employed in the process if desired. The reaction products from reactor 22 are removed therefrom and passed by way of line 27 through furnace 28 where the mixture is heated. The heated mixture of partially dehydrogenated oil and hydrogen is removed from furnace 28 by way of line 29. The mixture is passed by way of line 31 into dehydrogenation reactor 23. The total reaction products from reactor 23 are withdrawn by way of line 33. Recycled hydrogen may be introduced into line 33. The mixture passes by way of line 33 into furnace 34 wherein it is heated to the reaction temperature. The heated mixture is removed from the furnace and passed by way of line 36 into dehydrogenation reactor 24. The total reaction products are removed from reactor 24 and passed by way of line 37 into gas separator 38. A hydrogen stream is stream is recycled by way of line 41 and

valved lines

42 and 43 to the various dehydrogenation reactors.

a' -i iatevzftzso-l A liquid stream is removed from the bottom of gas separator 38 and passed by way of line 44 into fractionator 46. A small quantity of a naphtha fraction is recovered therefrom and passed by way of line 47 to further retining or to blending as shown herein. Lighter gases are removed from fractionator 46 by way of line 48. The dehydrogenated gas oil boiling range material, i.e. the oil boiling above 400 F. is removed as a bottom stream from fractionator 46 and passed by way of line 49 through the heating tubes of furnace 51. The dehydrogenated gas oil, from which a substantial portion of the sulfur and nitrogen compounds have been removed by virtue of the dehydrogenation step, is heated in furnace 51 to the temperatures employed in the mild catalytic cracking step. The heated oil is passed from the furnace by Way of line 52 into reactor 53 wherein the mild catalytic cracking step is carried out.

The dehydrogenated gas oil is catalytically cracked -under conditions to achieve a conversion of from 30 to 40% thereof into lower boiling hydrocarbon products. These cracking conditions may comprise a temperature of from 850 to 1050 F., a pressure of 5 to 50 p.s.i.g., a catalyst to oil ratio in the range of about 2:1 to 20:1 on a Weight basis, and a Weight space velocity in the range of about 1.0 to 30 pounds of oil/hour/pound of catalyst. The particular space velocity, temperature, etc. may vary somewhat depending upon the activity of the cracking catalyst used. Nevertheless, the conversion should be limited to 30 to 40%, for at higher levels of conversion excessive formation of carbonaceous deposits on the cracking catalyst will occur. Either a lluidized, moving, or fixed bed catalytic cracking reactor may be used. Siliceous cracking catalysts such as natural clay, activated natural clay, synthetic catalysts such as silicaalumina, silica-magnesia, silica-alumina-zirconia, etc. may be employed. In the embodiment shown herein the dchydrogenated light gas oil is catalytically cracked using a silica-alumina catalyst, a temperature of about 950 F., a pressure of about 25 p.s.i.g., a catalyst to oil weight ratio of about 5-l0, and a space velocity of about 10 pounds of oil/hour/pound of catalyst in the lluidized reactor. A conversion of about 35% of the gas oil charge into lower boiling products, principally high octane oletinic naphtha, is obtained.

The total reaction products are recovered from the mild catalytic cracking step and passed by way of line 54 into gas separator 55. Light gases are removed from the system by way of line 56. The liquid is passed from the separator by way of line 57 into fractionator 58. The high octane olenic naphtha is separated and passed by Way of line 59 to line 47 wherein it is blended with the naphtha from the dehydrogenation step. The catalytic cycle oil, i.e. oil which boils above 400 F. and which is recovered from the products of the mild catalytic cracking step, is removed from the bottom of fractionator 58 and passed by way of line 61 as charge oil to the solvent extraction step.

, vessel 62 employing liquid solvents which are selective l removed therefrom by way of line 39. A portion of this This catalytic cycle oil has a rather high aromatics content, e.g. 50 to 60% because of the earlier dehydrogenation step. This high aromatics content makes it an undesirable catalytic cracking charge stock since it would cause deposition of large amounts of coke upon the cracking catalyst and lead to rapid deactivation of the catalyst. The catalytic cycle oil from the mild catalytic cracking step, together with catalytic cycle oil from the severe catalytic cracking step and together with partially hydrogenated oil boiling above 400 F. from the destructive hydrogenation products, is passed by way of line '61 into the bottom of solvent extraction vessel 62. A typical solvent extraction process is carried out in `are liquid SO?, phenol, cresol, chlorex, furfural, etc.

They may be .used inamounts of from 25 to 200 volume percent based upon oil in a process which employs one or many extraction stages. In the embodiment shown herein liquid SO2 is employed as the selective solvent at an extraction temperature of about l to 25 C., eg. about 15 C. and under suiicient pressure to maintain the SO2 in the liquid phase. The extraction is carried out in three stages, employing about 50 volume percent of liquid SO2 based upon oil in each stage. ln the schematic diagram shown in the iigure liquid SO2 :from source 63 is passed by way of line 6d into the top of extraction vessel 62. The descending stream of liquid SO2 passes downwardly through the ascending stream of lighter oil and extracts aromatic hydrocarbons, as well as any residual amounts of sulfur compounds, from the catalytic cycle oil.

The raffinate phase which consists primarily of oil which is now lean in aromatics, together with some occluded SO2, is removed from the top of extraction vessel 62 and passed by way of line 6o into iiash drum 67. SO2 is vented from the oil in flash drum o7 and the SO2 is passed by way of line 63 into line ed for return to the extraction step. The hydrocarbon rafnate oil is removed from the bottom of flash drum 57 by way of line 69, freed of residual SO2 by equipment not shown herein, and passed as the charge stock to the severe catalytic cracking step. An extract phase consisting of liquid SO2 containing dissolved oil which is enriched in aromatic hydrocarbons, is removed from the bottom of extraction vessel 62 and passed by way of line 71 into ilash drum 72. SO2, which is iiashed from the extract phase in flash drum 72, is taken overhead and passed by way of line 73 into line 63 by which it is recycled to extraction vessel 62. The aromatics-rich hydrocarbon extract phase, which now has an aromatics content of about 80% more or less, is removed from the bottom of flash drum 72, freed ot residual SO2 by equipment not shown herein, and passed to the destructive hydrogenation step.

In the destructive hydrogenation step the aromaticsrich hydrocarbon extract is converted to naphtha having an F-l octane number of 95 to 100. Because substantially all of the sulfur compounds have been removed during the dehydrogenation step long catalyst life and unusually high octane naphtha are obtained during destructive hydrogenation. The catalyst employed in destructive hydrogenation is a dual-functioning catalyst which combines hydrogenation properties and cracking proper ties so as to cause hydrogenation of the extract oil and thereafter cracking of the oil. The hydrogenation cornponents of such a catalyst may be the oxides and/or sulfides of the metals of group 6 and/ or 8 of the periodic table (or the metals themselves). These are supported on a carrier having cracking properties such as natural and activated clays, synthetic catalytic cracking catalysts such as silica-alumina, silica-magnesia, silica-alumina* zirconia, or cracking bases such as HF promoted alumina. The catalyst may contain between l to 10%, preferably `about or thereabouts by weicht, of the hydrogenation component 'supported in extended form upon the cracking component. The catalyst may be prepared by any of the conventional techniques such as by impregnation of the support with an aqueous solution of the hydrogenation component, by precipitation of the hydrogenation component upon the cracking support, or by coprecipitation of the hydrogenation component with the cracking component. For example a silica-alumina cracking catalyst containing between 5 and 20% alumina with the remainder being silica, may be impregnated with a solution of ammonium molybdate, the impreg- -nated catalyst dried and then calcined to convert the ammonium molybdate to molybdenum oxide; thereby producing al catalyst containing about 5% MoOS. Other catalysts such as nickel on 4silica-alumina, iron on ..-ilica-a1umina,platinum on silica-alumina, rplatinum onv tluorided alumina, cobalt molybdate on fluorided alumina, molybdenum oxide on fluorided terrana earth,.and similar dual-functioning catalysts may be employed. This dual-functioning catalyst converts the polycyclic aromatics in the extract oil to naphtha by hydrogenating `one ring of the polycyclic and thereafter by reason of the cracking component of the catalyst this hydrogenated ring is cracked whereupon the naphtha boiling range monocyclic aromatic is produced. The destructive hydrogenaton in the presence of the dual-functioning catalyst is carried out at a temperature between about 800 to about l000 and at a pressure of about 1000 to 5000 p.s.i.g. while employing hydrogen in the amount of about 2000 to 10,000 scf/barrel of charge oil. A space velocity of from l to 20 volumes of oil/hour/volume of catalyst may be used. Conversions to lower boiling products on the order of or higher are obtained, most of it being naphtha having an octane of to 100 F--l or higher.

The extract oil is removed from flash drum 72 by way of line 74. A portion is passed by way of line 76 through the tubes of furnace '77 wherein the oil is heated to the temperature needed for the destructive hydrogenation reaction. The heated oil is then passed by way of line 7S into destructive hydrogenation reactor 79. The major portion of the hydrogen stream separated from the products of the dehydrogenation step is diverted from line 39 and passed by way of line 81 through the tubes of furnace 82 wherein it is heated and then passed by way of line 83 into line 78 for introduction with the charge oil to the destructive hydrogenation step. A portion of the extract oil is diverted from line 74, and thus bypasses furnace '77, and is passed by way of line 84 through a manifolding system and into destructive hydrogenation reactor 79. This cool oil stream which is introduced at different heights within the reactor assists in minimizing the temperature increase caused by the exothermic hydrogenation reaction. In this embodiment the destructive hydrogenation is carried out at a temperature of about 950 F., a pressure of about 3000 p.s.i.g., a hydrogen recycle rateiof about 5000 s.c.f./barrel of charge oil, a space velocity of about 5 volumes of oil/hour/volume of catalyst, while employing a molybdena on silica-alumina containing about 5% M003 and about 20% A1203. The large quantity of hydrogen generated during the dehydrogenation step provides hydrogen for the 2000 to 3000 s.c.f. of hydrogen consumed per barrel of oil charged to the destructive hydrogenation step.

The products from the destructive hydrogenation step are removed from reactor 79 and passed by way of line 86 into gas separation means 87. A hydrogen gas stream is taken overhead and passed by way of line 88 .into line S1 by which it is recycled to destructive hydrogenation reactor 79. A liquid bottom stream is removed from separator 87 and passed by way of line 89 into fractionator 91. The high octane naphtha is removed overhead and is passed by way of line 92 into line 47 wherein it is blended with the other high octane naphtha. components. A bottom stream is removed from fractionator 91 by way of line 93. This stream consists of an oil which has been partially hydrogenated during the destructive hydrogenation step, and this oil boils at a temperature above 400 F. It may be recycled in whole or in part by way of valved line 94 to destructive hydrogenation reactor 79. It may be fractionated and that portion thereof which boils above 600 F. may be passed by way of valved line 96 as a portion of the charge to the severe catalytic cracking step to be discussed later. This latter operation may suitably be carried out because the partially hydrogenated nature of the oil has reduced its coking tendency upon catalytic cracking and because it would produce a higher octane naphtha-during catalytic cracking than it would during destructive hydrogenation. This latter mode of operation mayalso .be employed when the gas oil charged to the dehydro- `a very substantial or major portion of oil boiling above 600 F. and this total oil in line 93 may suitably be passed to the severe catalytic cracking step. In the embodiment discussed herein, it is preferred to pass the partially hydrogenated oil boiling above 400 F. by way of line 93 through valved line 97 into line 61 by which it is recycled to the extraction vessel 62. In this manner the aromatic components are eventually returned to the destructive hydrogenation step and the paralnic components pass out with the railinate oil and are passed to the severe catalytic cracking step. This technique enables maximum yields of highest octane naphtha from the stream contained in line 93.

The raiinate oil which is removed from flash drum 67 by way of line 69 is passed by way of line 9S into the furnace tubes of furnace 99. The heavy gas oil, which boils at a temperature above 600 F., which has been segregated from the light virgin gas oil in fractionator 13 is passed by way of line 17 into line 98. The mixture of heavy virgin gas oil and raffinate oil which has been depleted in aromatics is heated in furnace 99 to the temperatures necessary for subsequent catalytic cracking. Heated oil is removed from the furnace and passed by way of line 101 to severe catalytic cracking reactor 102'. The oil is cracked under a severity to provide a conversion of at least 60% thereof into hydrocarbons boiling below 400 F. The conditions under which this is carried out comprise a temperature of 850 to 1050 F., a pressure of 5 to 50 p.s.i,g., a catalyst to oil ratio in the range of about 2:1 to 20:1 on a weight basis, and a weight space velocity between about 0.1

and 5 pounds of oil/hour/pound of catalyst. Siliceous cracking catalysts such as silica-alumina and the other types previously described in connection with the mild catalytic cracking step are used. While the precise conditions employed in the severe catalytic cracking step will vary to some extent depending upon the activity of the catalyst, particular conditions which may be used to achieve the conversion of at least 60% of the.

oil to lower boiling products are easily determined by those skilled in the art. In the embodiment described herein, the oil is catalytically cracked to achieve the minimum 60% conversion by employing a silica-alumina catalyst, a temperature of about 950 F., a pressure of about 25 p.s.i.g., a catalyst to oil weight ratio of about 10, and a space velocity of about 0.3 to 0.7 pound of oil/hour/pound of catalyst. Any of the various types of commercial catalytic cracking processes such as the uidized bed technique employed in this embodiment, moving bed, or fixed bed, etc. may be used.

The hydrocarbon products from the catalytic cracking step are moved from catalytic cracking reactor by way of line 103 and passed to gas separator 104. Light gases are taken overhead and removed from the process by way of line 106. The liquid is passed from the gase separator by way of line 107 into fractionating column 108. High octane naphtha is removed overhead from the fractionator and passed by way of line 109 into line 47 wherein it is blended with the other high octane naphtha fractions and the desired amount of light hydrocarbons to form high octane gasoline. A cycle oil stream is removed from the bottom of fractionator 108 by way of line 111. If desired this stream may be recycled by way of valved line 112. into line 17 and thereafter returned to the catalytic cracking step. In the embodiment shown herein it is'preferred to pass a portion or all of this cycle stock from severe catalytic cracking by way of valved line 113 into line 61 by which it is introduced as a part of the charge oil to the solvent extraction step. This enables the segregation of the aromatic components from the paranic com'- ponents of the cycle oil, the aromatic components being passed subsequently to destructive hydrogenation and the paraffnic components being returned to the severe catalytic cracking step.

-It is apparent from the foregoing description that the present invention provides an integrated system for producing maximum yields of high octane naphtha with minimum coke and gas formation by process which is self-sulcient with respect to hydrogen requirements.

Thus having described the invention what is claimed is:

1. A process for the conversion of gas oil to naphtha which comprises dehydrogenating a gas oil in the presence of hydrogen and a hydrogenation-dehydrogenation catalyst at a temperature between about 750 and 950 F. and a pressure in the range of about 200 to 800 p.s.i.g. and thereby producing hydrogen and a partially dehydrogenated oil, separating the products from the dehydrogenation step into a hydrogen stream, a naphtha stream, and a dehydrogenated gas oil stream, separating the hydrogen stream into rst and second portions, recycling the rst portion of the hydrogen stream to the dehydrogenation step, subjecting said dehydrogenated gas oil stream to mild catalytic cracking under conditions such that no more than 40% of the dehydrogenated gas oil stream is converted into hydrocarbons boiling below'400 F., recovering a catalytic cycle oil from the products of the mild catalytic cracking step and solvent extracting said cycle oil to separate an aromatics-rich hydrocarbon extract oil from an aromatics-lean hydrocarbon ratlnate oil, subjecting said rainate oil to severe catalytic cracking under conditions to convert at least 60% of said raffinate oil into hydrocarbons boiling below 400 F., and subjecting the aromatics-,rich extract oil to destructive hydrogenation by contacting it together with the second portion of the hydrogen stream produced during the dehydrogenation step with a catalyst having hydrogenation and cracking properties in the presence of 2000 to 10,000 s.c.f. of hydrogen/barrel of extract oil at a temperature of about 800 to 1000 F. and a pressure between about 1000 and 5000 p.s.i.g. and thereby producing high octane naphtha and a partially hydrogenated oil boiling above 400 F.

2. The process of claim 1 wherein the initial gas oil to be converted is fractionated at a cut point no higher than about 600 F. to produce a lower boiling gas oil and a higher boiling gas oil, and wherein the lower boiling gas oil is charged to said dehydrogenation step and the higher boiling gas oil is chargedto the severe catalytic cracking step.

3. The process of claim 1 wherein the products from said severe catalytic cracking step are fractionated to separate a naphtha fraction from a second catalytic cycle oil, and said second catalytic cycle oil and said partially hydrogenated oil boiling above 400 F. which is produced during the destructive hydrogenation step are passed as part of the charge oil to the solvent extraction step.

4. The process of claim 1 wherein the gas oil charged to the dehydrogenation step has an ASTM 50% point of at least 600 F. and wherein the partially hydrogenated oil boiling above 400 F. which is produced during the destructive hydrogenation step is passed to the severe catalytic cracking step.

5. A process for the conversion of gas oil to naphtha which comprises fractionating a virgin gas oil at a cut point no higher than about 600 F. and thereby producing a lower boiling gas oil fraction and a higher boiling gas oil fraction, contacting said lower boiling gas oil fraction in the presence of hydrogen and a hydrogenation-dehydrogenation catalyst at a temperature between about 750 and 950 F. and a pressure in the range of about 200 to 800 p.s.i.g. and thereby producing hydrogen and a partially dehydrogenated oil, separating the products from the dehydrogenation step into a hydrogen stream, a naphtha stream, and a dehydrogenated gas oil stream, separating the hydrogen stream into [first and second portions, recycling the iirst portion of the hydrogen stream to the dehydrogenation step, subjecting said dehydrogenated gas oil stream to mild catalytic cracking under conditions such that no more than 40% of the dehydrogenated gas oil stream is converted into hydrocarbons boiling below 400 F., recovering a rst catalytic cycle oil from the products of the mild catalytic cracking step and solvent extracting said cycle oil to separate an aroniatics-rich hydrocarbon extract oil from an aromatics-lean hydrocarbon rainate oil, subjecting said ranate oil and said higher boiling virgin gas oil fraction to severe catalytic cracking under conditions to convert at least 60% of the oil into hydrocarbons boiling below 400 rF., fractionating the products from said severe catalytic cracking step to separate a naphtha fraction from a second catalytic cycle oil, passing said second catalytic cycle oil as charge oil to the solvent extraction step, subjecting the aromatics-rich extract oil to destructive hydrogenation by contacting it together with the second portion of the hydrogen stream produced during the dehydrogenation step with a catalyst having hydrogenation and cracking properties in the presence of 2000 to 10,000 s.c.f. of hydrogen/ barrel of extract oil at a temperature of about 800 to 1000 F. and a pressure between about 1000 and 5000 p.s.i.g. and thereby producing high octane naphtha and a partially hydrogenated oil boiling above 400 F., and passing said partially hydrogenated oil boiling above 400 F. as ycharge oil tothe solvent extraction step.

References Cited in the le of this patent UNITED STATES PATENTS 2,242,504 Benedict et al May 20, 1941 2,249,584 Thomas July 15, 1941 2,748,055 Payne May 29, 1956

Claims

1. A PROCESS FOR THE CONVERSION OF GAS OIL TO NAPHTHA WHICH COMPRISES DEHYDROGENATING A GAS OIL IN THE PRESENCE OF HYDROGEN AND A HYDROGENATION-DEHYDROGENATION CATALYST AT A TEMPERATURE BETWEEN ABOUT 750* AND 950*F. AND A PRESSURE IN THE RANGE OF ABOUT 200 TO 800 P.S.I.G. AND THEREBY PRODUCING HYDROGEN AND A PARTIALLY DEHYDROGENATED OIL, SEPARATING THE PRODUCTS FROM THE DEHYDROGENATION STEP INTO A HYDROGEN STREAM, A NAPHTHA STREAM, AND A DEHYDROGENATED GAS OIL STREAM SEPARTING THE HYDROGEN STREAM INTO FIRST AND SECOND PORTIONS, RECYCLING THE FIRST PORTION OF THE HYDROGEN STREAM TO THE DEHYDROGENATION STEP, SUBJECTING SAID DEHYDROGENATED GAS OIL STREAM TO MILD CATALYTIC CRACKING UNDER CONDITIONS SUCH THAT NO MORE THAN 40% OF THE DEHYDROGENATED GAS OIL STREAM IS CONVERTED INTO HYDROCARBON BOILING BELOW 400* F., RECOVERING A CATALYTIC CYCLE OIL FROM THE PRODUCTS OF THE MILD CATALYTIC CRACKING STEP AND SOLVENT EXTRACTING SAID CYCLE OIL TO SEPARATE AN AROMATICS-RICH HYDROCARBON EXTRACT OIL FROM AN AROMATICS-LEAN HYDROCARBON RAFFINATE OIL, SUBJECTING SAID RAFFINATE OIL TO SEVERE CATALYTIC CRACKING UNDER CONDITIONS TO CONVERT AT LEAST 60% OF SAID RAFFINATE OIL INTO HYDROCARBON BOILING BELOW 400*F.,