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Publication numberUS2945800 A
Publication typeGrant
Publication dateJul 19, 1960
Filing dateJun 8, 1955
Priority dateJun 8, 1955
Publication numberUS 2945800 A, US 2945800A, US-A-2945800, US2945800 A, US2945800A
InventorsCiapetta Frank G, Coonradt Harry L, Garwood William E
Original AssigneeSocony Mobil Oil Co Inc
Export CitationBiBTeX, EndNote, RefMan
External Links: USPTO, USPTO Assignment, Espacenet
Multiple pass catalytic cracking
US 2945800 A
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Description  (OCR text may contain errors)

July` 19,

1960 F. G. clAPETTA ErAL MULTIPLE PASS CATALYTIC CRACKING Filed June 8, 1955 5 Sheets-Sheet 1 MaX- TW A oRNEY v July 19, 1960 F. G. CIAPETTA ETAL 2,945,800

MULTIPLE PAss CATALYTIC CRACKING Sheets-Sheet 2 Filed June 8, 1955 40 c'aAfz/fHs/a/v [100 #Wx/af), Vm.

60 40 Z0 caA//E/Ps/o/v /A/ro FUEL 0a, VUL.

INVENTORS BY n r l Z ff A RNEY July 19, 1960 F. G. CIAPETT ETAL MULTIPLE PASS CATALYTIC CRACKING Filed June 8, 1955 5 Sheets-Sheet 3 ATTORNEY July 1.9, 1960 l F. cal- CIAPETTA ETAL 2,945,800

n l i MULTIPLE PASS CATALYTIC cRAcxING File-d June 8, 1955 I 5 lSheets--Sheet 5 ATTO RN EY 2,945,800 MULTIPLE PASS CATALYTIC CRCKING Frank G. Ciapetta, Upper Darby,k Pa., and Harry L.

Coonradt, Woodbury, and William E. Garwood, Haddonleld, NJ., assignors to Socony Mobil Gil Conlpany, Inc., a corporation of New York Filed .lune 8, 1955, Ser. No. 513,972 4 Claims.V (Cl. Nsu-'59) This invention relates to the catalytic cracking of petroleum hydrocarbon stocks. lt is more particularly concerned with a processwherein relatively high-boiling hydrocarbon fractions' are converted, in the presence of platinumor palladium-containing catalysts, into valuable products;

As is well known to those familiar with the art, the major products of a cracking operation are dry gas, butanes, C54- light naphtha, heavy naphtha, and. cycle stock (boiling at temperatures higher than about 390 R). The light naphtha fraction usually has a relatively high octane number (about 90-92, F-1|-3 cc. TEL). On the other hand, the heavy naphtha fraction, particularly `that which is obtained by cracking in the presence of hydrogen, has a relatively low octane number (about 70-80). Accordingly, in order to produce a finished gasoline having a relatively high octane number, it has been the practice to blend the heavy naphtha fraction with the light naphtha fraction, and with butanes in amounts limited by the maximum permissible vapor pres-V sure. There is, however, a steadily increasing demand for higher octane gasolines (about 95 and higher).

As those skilled in the art will readily appreciate, such. octane requirements cannot be met by the aforedescribed conventional blending operations. Accordingly, the relatively low octane heavy naphtha. fraction Vhas been sube jected to reforming operations. As the increasing demand for higher octane gasolines must be satisfied largely by reforming, instead o f by blending, there is a greater demand for heavy naphtha fractions that can be reformed and a correspondingly lesser `demand for lightv naphtha fractions that can be used for blending purposes. The light naphtha fraction usually is not subjected to a reforming operation because it produces excessive amounts of dry gas, coke, etc. The yield of gasoline obtained by reforming. light naphtha, therefore, is prohibitively small. Accordingly, itwill be appreciated that a cracking' operation that will produce greater amounts ofthe heavynaphtha fraction and lesser amountsV ofli'ght naphtha fractions is highly desirable.

As is well known to those skilled in the art, the material that is obtained in a cracking operation that boilsat temperatures higher than about 390 F. is a fuel oil suitable for use in diesel engines and in domestic heating units. The burning efficiency of this fuel oil is dependent upon its diesel index-the higher the diesel index, the better the burning elliciency of the fuel. It is highly desirable, therefore, to obtain a fuel oil having the highest possible diesel index.

In copending application Serial Number 418,166, filed on March 23, 1954, which lis a continuation-in-part of application Serial Number 351,151, tiled on April 27, 1953, there is disclosed a once-through process for crack.- ing high-boiling hydrocarbon fractions in the presence of hydrogen and of catalysts comprising metals of the platinum and palladium series supported upon synthetic mixed oxide carriers. This process can be operated to produce substantially gasoline alone or fuel oil alone, or both. At intermediate conversion levels (levels at which both naphtha and fuel oil are made), a good product distribution is obtained. The amount of dry gas4 is rela.- tively small and the amounts of butanes and. of pentanes produced` are not in excess of those required to produce 2 10-pound R.V.P. (Reid Vapor Pressure) gasoline. Substantial yields of light and heavy naphtha are obtained and the diesel index of thefuel oil is relatively high. In View of the foregoing, however, the process would be more eliicient and its performance would be more desirable, if it could be operated to produce even less dry gas, less light naphtha, more heavy naphtha, and a fuel oil having a still higher diesel index.

It has now been found that the cracking process described in the aforementioned applications can be operated in a manner that will produce the following results:l lower dryA gas yield, lower yield of C5+ light naphtha, greater heavy naphtha production, and a fuel oilk having a higher diesel index; and, when combined with a reforming operation in the presence of hydrogen, a greater overall yield of high-octane gasoline. It has been discovered" that these results can be obtained by an operation in which a portion of the desired ultimate conversion into products boiling below about 390 E. (100- recycle) is effected in a rst reaction zone, and the remaining conversion is effected by cracking the effluent that boils above 390 F. in one or more separate reaction zones.

Accordingly, itis an objectV ofv this invention to provide a method forobtaining more eliicient product distribution in a cracking operation. Another object is to provide a process for cracking relatively high-boiling hydrocarbon fractions that will produce, as compared with a once-through operation, the following results: lower dry gas' yield, lower yield of light naphtha, greater heavy naphtha production, and a fuel oil having a higher diesel index. A further object is to provide a process for producing high octane gasoline in increased amounts. A specic object is to provide a process for cracking hydrocarbon charge stocks in the presence of hydrogen and of catalysts comprising metals of vthe platinum and palladium series deposited upon a synthetic composite of two or' more refractory oxides which has a relatively high cracking activity, that will produce the aforementioned results. Another specific object is to provide a process for cracking and reforming a hydrocarbon charge stock in the presence of hydrogen and of catalysts containing platinum and palladium series metals, that will produce increased amounts of high octane gasoline and a fuel oil having a high diesel index.

Otherl objects and advantages of. this invention will become apparent to those skilled in the art from the following detailed description considered in conjunction with the drawings, in which:

Figure l presentsa series of curves showing graphically the relationship' between the volume percent conversion into products boiling at temperatures lower than about 390 F. (10G-recycle) and the weight percent yield of dry gas obtained by 'cracking a typical virgin gas oil and a` typical Coker gas oil in the presence of hydrogen and of a platinum-containing catalyst, in a once-through operation;`

Fig. 2 presents aseries of curves showingV graphically the relationship between the volume percent conversion into products boiling at temperatures lower than about 390 (1D0-recycle) and the volume percent yield of butanes,l and between the volume percent conversion into products boiling at temperatures lower thanY about 390 F. (1GO-recycle) and the volume percent yield of C5+ light naphtha obtained by cracking a typical virgin gas oil and a typical coker gas oil in the presence of hydrogen and of a platinum-containing catalyst, in a once-through operation;

Fig. 3 presents a series of curves showing graphicallyv thev relationship between. thel volume percent conversion into products boiling at temperatures lower than about` 390 F. (1GO-recycle) and the volume percent yield of,l

- A A a heavy naphtha obtained by cracking a typical virgin gas oil and a typical coker gas oil in the presence of hydrogen and of a platinum-containing catalyst, in a oncethrough operation;

Fig. 4 presents a series of curves showingY graphically the relationship between the volume percent conversion into fuel oil, the material boiling at temperatures higher than about 390 F., and the diesel index of the fuel oil obtained by cracking a typical virgin gas oil and a typical coker gas oil in the presence of hydrogen and of a platinum-containing catalyst, in a once-through operation;

Fig. 5 presents a schematic arrangement of an embodiment of the process of this invention;

Fig. 6 presents a schematic arrangement of another embodiment of the process of this invention; and

Fig. 7 presents a schematic arrangement of still another embodiment of the process of this invention adapted to the production of high octane gasoline.

Stated broadly, the present invention provides a processboiling point of at least about 400 F., a 50 percent-pointv of at least about 500 F., and an end-boiling point of at least about 600 F., and boiling substantially continuously between said initial boiling point and said endboiliug point, with a catalyst comprising between about 0.05 percent and about 20 percent, by weight of the.

catalyst, of at least'one metal of the platinum and palladium series deposited upon a synthetic composite of at least two refractory oxides, said composite having an activity index of at least about 25, in the presence of hydrogen in amounts, expressed in molar ratio of hydrogen to hydrocarbon charge, varying between about 2 and about 80, at pressures varying between about 100 pounds per square inch gauge and about 2500 pounds per square inch gauge, at a liquid hourly-space velocity varying between about 0.1 and about 10, and at temperatures varying between about 400 F. and about 825 F., to effect a portion of the desired ultimate conversion into products boiling at temperatures lower than about 390 F., and to produce an efuent material comprising products boiling at temperatures lower than about 390 F., a product boiling at temperatures higher than about 390 F., and a gaseous fraction rich in hydrogen; contacting at least the portion of the efuent that boils at temperatures higher than about 390 F., in at least one' separate reaction zone with a catalyst of the type used in the first reaction zone, in the presence of hydrogen in amounts, expressed in molar ratio of hydrogen to hydrocarbon charge, varying between about 2 and about 80, at pressures varying between about 100 pounds per'V square inch gauge and about 2500 pounds per square inch gauge, at a liquid hourly space velocity varying between about 0.1 and about 10, and at temperatures l varying between about 400 F. and about 825 F., to

effect the remainder of the desired conversion into products boiling at temperatures lower than about 390 F.

Throughout the specification and the claims, the term conversion is intended to be a generic term for the amount of products boiling at temperatures lower than about 390 F. (100-recycle), of gasoline, or of fuel oil obtained in the process. It is expressed in terms of the volume percent of the initial charge which is transformed in the process. The amount of product boiling at temperatures lower than about 390 F. is obtained by subtracting the volume percent of cycle stock (fuel oil) that boils at temperatures higher than about 390 F. from 100 percent, i.e., from the initial volume of the charge. The expression (100-recycle) is an abbreviation for 100 percent minus the volume percent recycle.

' As the cycle stock (i.e., the efuent boiling at temperatures higher than about 390 F.) is an excellent fuel oil, conversion into fuel oil is the volume percent of product which boils at temperatures higher than about 390 F. The volume percent of conversion into products boiling at temperatures lower than about 390 F. (100-recycle) and the volume percent of conversion into fuel oil totals to 100 volume percent, based upon the initial charge. Dry gas refers to the methane, ethane, propane, and ethylene and propylene produced in a cracking process, expressed in terms of weight percent of the initial charge. C5+ light naphtha is the product that boils between about F. and about 170 F. The heavy naphtha" is the product that boils between about 170 F. and about 390 F. The diesel index of the fuel oil is a function of the A.P.I. gravity and of the aniline number, as defined by Becker et al. in the S.A.E. Journal (Transactions), vol. 35, No. 4, p. 377. The cracking activity of a carrier is expressed in terms of the percent, by volume, of a standard hydrocarbon charge which is cracked, under specific conditions, in the Cat. A test. This test is described by Alexander and Shimp in National Petroleum News, 36, page R-537 (August 2, 1944). The unit for rating the cracking activity of a material is called the Activity Index (A.I.).

The catalysts utilizable herein are those described in copending application Serial No. 351,151, filed on April 27, 1953, now abandoned; and in the continuation-inpart thereof, Serial No. 418,166, tiled on March 23, 1954, now abandoned. Briefly, these catalysts comprise between about 0.05%, by Weight, and about 20%, by Weight, of the nal catalyst, preferably between about 0.1% and about 5%, by weight, of the metals of the platinum and palladium series, i.e., those having atomic numbers of 44-46, inclusive, 76-7 8, inclusive, supported upon synthetic composites of two or more refractory oxides. The carrier is a synthetic composite of two or more oxides of the metals of groups IIA, HIB and IVA and B of the Periodic Arrangement of Elements [1. Chem. ed., 16, 409 (1939)]. vrI'hese synthetic composites of refractory oxides must have an activity index of at least about 25. They can also contain halogens and other materials which are known in the art as promoters for cracking catalysts, or small amounts of alkali metals that are added for thc purpose of controlling the activity index of the carrier. Non-limiting examples of the composites contemplated herein include silica-alumina, silica-Zirconia, silica-alumina-zirconia, alumina-boria, silica-alumina-uorine, and the like. The preferred support is a synthetic composite of silica and alumina containing between about 1%, by weight, and about by weight, of alumina. These synthetic composites of two or more refractory oxides can be made by any of the usual methods known to those skilled in the art of catalyst manufacture. Examples of methods of preparing them are set forth in co-pending applications Serial Nos. 351,151 and 418,166 referred to hereinbefore.

The following example illustrates a method of preparing a platinum-containing catalyst utilizable in the process of this invention:

EXAMPLE l A synthetic silica-alumina carrier or support containing 10%, by weight, alumina was prepared by mixing an aqueous solution of sodium silicate (containing 158 g. per liter of silica) with an equal amount of an aqueous acid solution of aluminum sulfate containing 39.4 g. A12(SO.-,)3 and 28.6 g. concentrated H2804 per liter. The mixture was dropped through a column of oil wherein gclation of the hydrogel was effected in bead form. The bead hydrogel was soaked in hot water (about F.) for about 3 hours.. The sodium in the hydrogel was then removed by exchangingthe gel with an aqueous solution ot' aluminum sulfate [1.5% Al2(SO.,)3 by weight] containing a small amount (0.2 percent by weight) of ammonium sulfate. The thus-exchanged hydrogel bead was waterwaslldf' Bleu, it was dried in superheated steam (aboutA tained on a 25-mesh screen (U.S. Standardy Screen Series). u

was used for catalyst preparation.

Portions of the crushed, calcined carrier: were then barely covered with aqueous solutions of chloroplatinic acid, of concentrations sufficient to produce the desired amount of'metal in the finished catalyst. The Vexcess solution Was ,removedV by centrifuging.V The thus-impregnated carrier was then dried at 230 F. for 24 hours. The catalyst was treated with hydrogen for. four hours at 400 F. Then, it wasV activated inv hydrogen for 16 hours before it was used. The catalyst thus-prepared contained 0.47% platinum, by weight ofthe catalyst, and the silica-alumina carrier had an activity index of 4'6.

The charge stocks utilizable herein are hydrocarbon fractions having an initial boiling point ofV at least: about 400 F., a 50l percent-point of at least about 500 F'. and an end-boiling point of at least about 600 F. and boiling substantially continuously between said initial boiling point and said end-boiling point. Such charge stocks include gas oils, residual. stocks, refractory cycle Stocks from` conventional cracking, whole topped crudesgf and heavy hydrocarbon fractions derived by the destructive hydrogenation of coal, tars, pitches, asphalt@ etc., such as, for example, middle oil.

As is'well known to those skilled in the art, the distillation of higher-boiling petroleum-fractions (those boiling at temperatures higher than about 750 Fg) must be carriedV out under vacuum, in order to avoid thermal cracking. Throughout the specilication and in the claims, however, the boiling temperatures are expressed in terms of the boiling point at atmospheric pressure. vIn other words, in all instances, the boiling points of fractions distilled under vacuum have been corrected to the corresponding boiling points at atmospheric pressure.

As is well known to those familiar vviththe` art, the term gas oil is a broad, lgeneral term that covers a variety of stocks. Throughout the specification and .in the claims, the term, unless further modiiied, includes any fraction distilled from petroleum which has an initial boiling point of. atleast ahout400 F., a 50'percent-point of at least about500" F., and an end-boiling point of at least about 600 "c F.,.and boiling substantially continuously betweenthe initial boiling point andl the end boiling point. The portion which is not distilled is considered residual stock. The exact boiling range of a gas oil, therefore, will be: determined" by the initial distillation temperature (initial. boiling point) .the 50 percent-point, and by the temperature at which distillation is 'cut-orf (end-boiling point). Y

In practice, petroleum distillations have been made under vacuum up to temperatures as high as 1100-1200" F. (corrected to atmospheric pressure). Accordingly, in the broad sense, a gas oil is a petroleum fraction which boilsy substantially continuously between two temperatures that establish a range falling within from about 400 F; to about l1001200 F., the 50 percent-point being at least about 500 F. Thus, a gas oil could boil over the entire range 400-1200 F. or it could boil over a narrower range, e.g., SOO-900 F.

The gas oils can be further roughly subdivided by overlapping boiling ranges. Thus a light gasoil boils between about 400 F. and about 60G-650 F. A medium gas oil distills between about 60G-650 F. and about 700-750 F. A heavy gas oil will boil. between about 60G-650 F. and about 80G-900 F. A gas oil boiling between about 800-850 F. and about 1100-1200" F. is sometimes designated as a vacuum gas oil. It must be understood, however, that a gas oil can overlap the foregoing ranges. lt can even span several ranges, i.e. include, for example, light and medium gas oils.

fraction which is not distilled. Therefore, any fraction, regardless of its initial boiling point, which includes all the heavy bottoms, such as tars, asphalts, etc., is a residual fraction. Accordingly, a residual stock can be the portion of the crude remainingl undistilled at ll00-l200 F., or it can be made upfof a gas oil fraction plus the portion undistilled atl l100-,1200 F. `A Whole vtopped crude, as the name implies, is the entire portion of the crude remaining after the light ends (the portion boiling up to about 400 F.) have been removed by distillation. Therefore, such a fraction includes the entire gas oil fraction (400 F. to ll00-l200 F.) and the undistilled portion of the crude petroleum boiling above 1100-l200 F. If it is desired, the residual fractions and the whole topped crude can be deasphalted by any means known tothe art. Such treatment, however, is not necessary for charge stocks intended for use in the process of this invention.

The refractory cycle stocks are cuts of conventionally cracked stocks whichV boil above the gasoline boiling range, usually, between about 40 F. and about 850 F. The refractory cycle stocks can be charged to the process of this invention in conjunction with a fresh petroleum charge stock, or they can be charged alone to the process. The process of this invention is particularly adaptable to the cracking of sulfur-containing charge stocks. The catalysts utilizable in the process of this invention, quite unexpectedly, are not deactivated by sulfur compounds, under the conditions of the process.`V

The presence of even relatively small amounts of nitro gen compounds in the charge stock interferes with the process of this invention. For relatively short terms of operation, the presence of nitrogen in amounts of as much as about 0.12 percent, by weight, and higher can be tolerated in the charge. When operating with such charge stocks, however, it is necessary to resort to intermittent operation.

Charge stocks .that contain about 0.1 percent nitrogen,

Vor less, can be cracked in a continuous operation, over long periods of time, without a loss in catalyst activity. Accordingly, the cracking charge stocks should contain less than about 0.1 percent nitrogen, by weight, when continuous operation is desired. Preferably, the nitrogen content should be less than about 0.08 percent, by weight. The reduction in nitrogen content can be effected by any of the methods well known in the art, such as, for ex-V ample, acid treatment, propane deasphalting, and hydrogenolysis under very high pressure, in contact with catalysts such as molybdenum or tungsten oxide, nickel sulfide, tungsten suliide, cobalt molybdate, cobalt tungstate, etc. As indicated hereinbefore, somewhat higher nitrogen contents can be tolerated, if the operation is intermittent or of relatively short duration. A higher nitrogen content can be tolerated in the charge, under more severe operating conditions, such as, at higher temperatures.

Once-through operation In order to compare the multiple-pass operation of this invention with a once-through operation, a series of runs was made at various conversion levels using a oncethrough operation.

EXAMPLE 2 The charge stocks used in this example were two gas oils derived from the Kuwait crude. They had the following properties:

As, mentioned hereinbefore, a residual stock is any iASTM distillati-nv ing in the presence of hydrogen and of the catalyst described in Example 1, after the latter had reached equilibrium, i.e., had been in continuous operation for more than ve days. The hydrogen pressure used was 1000 p.s.i.g. The liquid hourly space velocity was 0.5 and the molar ratio of hydrogen to oil was 40. Each portion of charge stock was contacted with the catalyst under these conditions at a different temperature. Pertinent data and the product distribution are set forth in Table I.

EXAMPLE 3 The charge stock used in this example was alight gas oil produced by coking a Mid-continent residual. This Mid-continent coker gas oil had the following properties:

API gravity 33.3 Distillation, ASTM 1BP F 420 50% .P 535 EP F 664 Sulfur weight percent 0.48 Diesel index 46.3

Portions of the gas oil were subjected to cracking m the presence of hydrogen and of the platinum-containing catalyst that was used in the runs described in Example 2. A different temperature was used in each run. In all runs, the pressure used was 1000 p.s.i.g., the liquid hourly space velocity was 0.5 and the hydrogen to oil molar ratio was 40. Pertinent data and product distribution for these runs are set forth in Table I.

TABLE I Kuwait Gas Oil Charge Stock Coker Gas Oil Temperature. F 620 G70 695 730 750 770 79 Conversion, Vol. percent. 21. 7 45. 1 63. 3 56. 3 25. 8 43. 2 66. 5 Dry Gas, Wt. percent 1.3 1.4 3. 1 2. 1 1.5 2. 5 4.2 Butanes, Vol. percent 3.2 4.8 v7. 7. 6 2.6 6.5 11.8 Cs-I- Light Naphtha, V01. percent .s 3.4 11.4 15.7 13.3 4.1 9.6 19.8 Heavy Naphtha, Vol.

percent 19. 7 35.6 51. 5 46. 6 24. 8 35. 5 46. 6 Fuel Oil, Vol. percent.... 78. 3 64. 9 36.7 43. 7 74.2 56. 8 33. 5 Diesel Index of Fuel Oill 73. 7 62.1 61.8 59. 5

The curves set forth in Figs. 1 through 4 are based upon the data in Table I. In Fig. 1, curve 1 shows the relationship between the volume percent conversion into products boiling at temperature lower than about 390 F. (G-recycle) and the weight percent of dry gas produced when the Kuwait gas oils were cracked in the presence of the platinum-containing catalyst in a once-through operation. Curve 2 shows a similar relationship obtained in the case in which the Mid-continent coker gas oil was cracked.

In Fig. 2, curve 3 shows the relationship between the volume percent conversion into products boiling at temperatures lower than about 390 W. (1D0-recycle) and the volume percent of butanes produced when the Kuwait gas oils were cracked, in a once-through operation, in the presence of the platinum-containing catalyst. Curve 4 shows a similar relationship obtained in the case in which the Mid-continent coker gas oil was cracked in a oncethrough operation.

Also in Fig. 2, curve 5 shows the relationship between the volume percent conversion into products boiling at temperature lower than about 390 F. (10U-recycle) and the volume percent of Cyllight naphtha produced when once-through operation.

In Fig. 3, curve 7 shows the relationship between the volume percent yield ofV heavy naphtha and the volume percent conversion into products boiling below about 390 F. (10U-recycle) in the case in which the Kuwait gas oils were subjected to cracking in a once-through operation. Curve 8 shows a similar relationship obtained in thecase in which the Mid-Continent coker gas oil was cracked in a once-through operation.

In Fig. 4, curve 9 shows the relationship between the diesel index of the fuel oil and the volume percent conversion into fuel oil boiling at temperatures higher than about 390 F. obtained by cracking the Kuwait gas oil (B) in a once-through operation. Curve 9a presents a similar relationship obtained in the case in which the Mid-Continent coker gas oil was cracked in a once-through operation.

Multiple pass operation A schematic arrangementY for carrying out one embodiment of the process of the present invention is shown in Fig. 5. The processing system comprises two reactors 10 and 11, arranged in series with an intermediate fractionator 12. In operation, a suitable hydrocarbon charge stock is introduced via a pipe 13 and pumped by a pump 14 through a pipe 15 into a furnace 16, or other suitable heating device, wherein it is pre-heated to the desired temperature. The heated charge stock is passed through pipes 17 and 18a into the reaction zone 10. Hydrogen gas, introduced via a pipe 19, is compressed to the desired reaction pressure by means of a compressor 20. The compressed hydrogen is passed through pipe 18, heated in a heater 51 and then passed into the reactor 10 through pipe 18a, in admxture with the hydrocarbon charge stock.

In the reactor 10, the mixture of hydrogen and hydrocarbon charge is brought into contact with a platform or palladium series metal catalyst of the type described hereinbefore, in order to effect partial conversion of the charge stock into products boiling at temperatures below about 390 F. The total effluent from the reactor 10, which includes products boiling at temperatures below about 390 F., products boiling at temperatures higher than 390 F. and hydrogen, is passed through a pipe 21 through a condenser 22, wherein it is cooled to temperatures at which hydrogen gas can be separated. The thus-cooled effluent is then passed through a pipe 23 into a high pressure gas separator 24.

In the gas separator 24, the effluent is separated into a liquid phase and a gas phase that comprises substantially pure hydrogen. This gas is removed from the gas separator via a pipe 25 and recycled through pipe 18 to supply a portion of the hydrogen required in the process. The liquid product from the gas separator 24 is passed through a pipe 26 into a depressuring zone 53, and thence through a pipe 26a into a fractionator 12.

In the fractionator 12, the products are separated into fractions comprising dry gas, butanes, C5+ light naphtha, heavy naphtha, and products boiling at temperatures higher than about 390 F. (cycle stock). The dry gas is removed through a pipe 27 and can be sent to the gas plant or other gas processing units. Butanes are removed through a pipe 28 for use in the blending of finished gasoline. 'I'he C54- light naphtha is removed through a pipe 29 and is also utilized for gasoline blending. The heavy naphtha fraction is removed through a pipe 30. This fraction has a relatively low octane number and, therefore, it is ordinarily subjected to a reforming process.

The product boiling at temperatures higher than 390 F. is removed from the fractionator 12 through a pipe 31 and. pumped by means of a pump 32 through a pipe 33 into a heater 34. In the heater 34, the product boiling above 390 F. is heated to the desired temperature. The thus-heated material is'passed through pipes 35'aud 36a into the reactor 11. Hydrogen gas, introduced through a pipe 37, is compressed to reaction' pressure by means of a. compressor 38 and passed through pipe 36 into a heater 2. The hydrogen is heated to suitable temperature in heater 52 and then passed through a pipe 36a where it is admiredY with the hydrocarbon charge. The mixture of hydrogen and hydrocarbon charge is passed intof the reactor 11.

' In the reactor 11, the mixture of hydrogen and hydrocarbon charge is contacted with a platinum or palladium seriesmetal catalyst of the type described hereinbefore, to effect Vfurther cracking of the charge. The catalyst contained in reactor 11 can be ofv the same type as that contained in reactor 10, or it' can be a different type of platinum or palladium 'series metal catalyst. The total efuent from the reactor 11 is passed through a pipe 39 and through a condenser 40, wherein it is cooled to temperatures at which hydrogen can be separated. The thuscooled' eiuent is passed through a pipe 41 into a high pressure gas separator 42.

In the gas separator 42, the eiiiuent is separated into a gaseous'phase and a liquid phase. The gaseous phase comprises substantially pure hydrogenV gasand it is recycled: to thel process through a pipe 43 and thence througha pipe 36', to supply a portion of the hydrogen required in the process. The liquid: product obtained in the gas separator 42` is passed through a pipe 44 into a depressuring zone 54, and thence through a pipe 44a into ay fractionator 45.

In the fractionator 45, the products are separatedv into suitable fractions.` Dry gas is removed through a pipe 46 andV passed to the gas plant. The butanes are' removed through. a: pipe 47- and transferred to the gasoline blend'- ing operations. The C5-llight naphtha is removed through apipe 48 and, likewise, is` sent to gasoline blending operations. The heavy naphtha fraction is removed through a pipe 49. This fraction will ordinarily be subjected to a reforming process; The remaining material lin the fractionator 45 is materialv boilingv at temperatures above about-4 390 F. and it is substantially No. 2 fuel oil. Thisfraction. is removed through a pipe 50 and can' be used as a heating-foil or as a diesel fuel, or the like.

EXAMPLE 4Y The charge stock` in this run was a light gas o'i derived from. coking a residualV obtainedl from a Mid-Continent crude, that was described in Example 3. This light Coker gas oil was subjected to a two-step cracking operation, such as is illustrated in Fig. 5*. The catalyst in the first stage reactor was the catalyst described in` Example l. The coker gasA oil was pumped into the reactor at a liquid hourly space velocityv of 0.5. Hydrogen was introduced in amounts, measured in terms of the molar ratio of hydrogen to hydrocarbon'charge, of 40 and at a pressure of about 1000 psig. The mixture of hydrogen and coker gas oil was contacted with the catalyst bed at temperatures of about 770 F., thereby eiecting 43.2

- volume percent conversion of the charge into products boiling at temperatures lower than about 390 F. The effluent was cooled to about 70 F. and passed into a gas separator wherein substantially pure hydrogen was separated from the liquid product. The liquid product was subjected to fractionation. The distribution of products obtained from this reaction is set forth in Table Il, column A.

, The product boiling at temperatures greater than about 390 F. amounted to about 56.8 percent of the original charge. This material was passed into a second reaction zone at a temperature of about 700 F. and at a liquid hourly space velocity of about 0.5. Hy-

drogen under pressures of about 1000 p.s.i.g. and in amounts, measured in terms of the molar ratio of hydrogen to hydrocarbon, of about 40was also passed into the second stage reactor in an admixture with the material boiling at temperatures above about 390 F. This mixture of hydrogen and lhydrocarbon was contacted with another portion of the catalyst described in Example 1. The total eliiuent `from the second reactor was cooled to about 70 F. and passed into a gas separator; wherein substantially pure hydrogenv gas was removed. The liquid product from the gas separator was subjected to fractionation. The distribution of products derived from this operation, based upon the amount of charge to the second stage reactor, is set forth in Table Il, column B.

TABLE II First Second Second Total On'cc- Pass Pass Pass Conver- Through A B l C 2 sion D Conversion E Conversion, Vol. per-` v cent 43. 2 51.6 29; 3 72. 5 72. 5 Dry Gas, Wt. percent 2. 5 2.3, 1.3 3.8 4.5 Butanes, Vol. percent". G. 5 8. 9 5.1 11. 6 v 12. 7 05+ Light N aphtha,

Vol. percent 9.6 Y 12.8 7.2 16. 8- 22.7 Heavy Naphtha, Vol. I

percent 35.6' 38.6- 2119 57. 4 48. 0 Diesel Index of Fuel Oil. 78.8 78. 8 78.8 59. 0

yto the process, it is necessary, therefore, to apply a facto-r of 56.8; the actual amount of material charged to the second reactor based upon the initial charge to the process. Accordingly, the figures set forth in column C of Table il represent the product distribution of the effluent from the second reactor based upon. the amount of the original charge to the process. The total product distribution and the total amount of conversion into products boilingA at temperatures below about 390 F. achieved in the process is the sum of the corresponding figures in column A and column C. These surns are set forth in column D of 'Fable' Il. Thus, for example, the total amount` of' conversion into products boiling below about 390 F. in the two-step process described in Example 4 4is about 72.5 volumeV percent, expressed in terms of the volume percent of initial charge.

Referring no-w to Figs. 1 through 4, the product distribution, in the case in which the conversion of 72.5 volume percent is achieved in a one-pass` operation can be obtained from curves 2, 4,. 6,V 8,v and 9a. This isdone by noting, in each instance, the point at which the abscissa corresponding to 72.5 volume percent conversion into products boiling at temperatures lower than about 390 F. (or to 27.5 volume percent yield of fuel oil) intersects the curve, and reading on thej ordinate the value for dry gas yield, butane yield, etc. at that point. The product distribution thus determined is set forth in column E of Table II for purposes of comparison.

By comparing columns D and E of Table Il, it will be noted that a much more desirable product distribution was achieved in the two-step operation of this invention. Thus, at the same conversion level, as compared to the yields obtained in a once-through operation, the amounts of dry gas, of butanes, and (D5-{- light naphtha were substantially less.V On the other hand, a substantially greater amountof heavy napht-ha was obtained. As stated hereinbefore, this nap'thav is`r readily reformed and the production of larger amounts of it is highly 'desirable for the overall process of producing high' octane gasoline.

of 550 F. material is all utilizable as No. 2 fuel oil, diesel fuel,

lor the like. From the data set forth in columns D and E of Table II, it is to be noted, furthermore, that the fuel oil derived from the multiple-pass operation of this invention has a diesel index of 78.8 as compared with a diesel index of only 59 when the same amount of fuel oil is produced in a once-through operation. This, of course, represents a greater eiciency in burner and in diesel engine operation.

EXAMPLE The charge stock used in this run was a gas oil (B) derived from a Kuwait crude and was described in Example 2. This gas` oil was subjected to a two-step cracking operation such as is illustrated in Fig. 5. The gas oil was pumped through a first-pass reactor, at a liquid hourly space velocity of 0.5 in contact with the catalyst used in Example 2 and in the presence of hydrogen. Hydrogen was pumped into the reactor in a molar ratio of hydrogen to hydrocarbon charge of 40 and under a pressure of about 1000 p.s.i.g. This mixture of hydrogen and gas oil was contacted with the catalyst bed at temperatures of 730 F. The effluent was subjected to gas separation and the liquid productA was then fractionated. The distribution of product obtained from this run is set forth in column A of Table III.

The amount of product boiling at temperatures above about 390 F. amounted to 43.7 volume percent of the original charge. This material was passed in a secondpass reaction wherein it was subjected to cracking in the presence of hydrogen and of the catalyst used in Example 2. This second-pass reaction was conducted at a temperature of 650 F., at a liquid hourly space velocity of about 0.5 and under a hydrogen pressure of about 1000 p.s.i.g., using a molar ratio of hydrogen to hydrocarbon oil of about 40. The distribution of products obtained from this run is set forth in column B of Table III.

TABLE III First Second Second Total Once- Pass Pass Pass Conver- Through A B 1 C i sion D Conversion E Conversion, Vol. percent 56. 3 26. 6 11. 6 67. 9 67. 9 Dry Gas, Wt. percent.-. 2.1 1.3 0.6 2.7 3.1 Bntanes, Vol. percent... 7. 6 3. 5 l. 5 9. 1 9. 7 05+ Light aphtha,

Vol. percent 13. 3 5. 2 2. 3 15. 6 18. 0 Heavy Naphtha, Vol.

percent 46. 6 21.2 11.6 55. 9 53. 5 Diesel Index of Fuel oil- 83. 2 83. 2 72 1 Based upon volume of charge to second reactor.

i Based upon volume of initial charge.

As in discussion of the data of Table II, the product distribution in the second pass is based upon the amount of material charged to the second reactor. In order to calculate the amount of product obtained from the second pass in terms o'f the amount of original charge of the process, it was necessary to apply a factor of 43.7. Accordingly, the figures set forth in column C of Table III represent the distribution of the products obtained from the second reaction pass based upon the amount of the original charge. The total product distribution and the total amount of conversion into products boiling below about 390 F. that was achieved yin the two-pass operation is set forth in column D of Table III.

By referring to Figs. l through 4, curves 1, 3, 5, 7, and 9, the product distribution when the conversion of 67.9 volume percent is achieved in a one-pass operation was obtained, in the manner described in connection with curves 2, 4, 6, 8, and 9a. This latter product distribution 12 is set forth in column E of Table III, for purposes of comparison.

Upon comparing columns D and E of Table III, it will be noted that a more favorable pro'duct distribution was achieved when the Kuwait gas oil was cracked in the twostep operation of this invention. At the same conversion level, as compared to a once-through operation, the amounts of dry gas, butanes, and of C54- light naphtha were substantially less. On the other hand, a substantially greater amount of heavy naphtha was obtained, for

use in reforming operations to produce high octane gasoline.'

The product boiling at temperatures above about 390 F. had an end-boiling po'int in the ASTM distillation of 586 F. Accordingly, all the material is utilizable as No. 2 fuel oil or as diesel fuel. Referring again to columns D and E of Table III, it will be noted that the fuel oil obtained by cracking a Kuwait gas oil in the multiplepass operation of this invention had a diesel index of 83.2; as compared with a diesel index o'f only 72 when the gas oil was cracked in a once-through operation. It will be appreciated, therefore, that the fuel oil produced in the process of this invention has superior qualities.

A flow-sheet for another embodiment of the multiplepass o'peration of this invention is shown in Fig. 6. The hydrocarbon charge stock, a gas oil or the like, is introduced through a pipe 60 into a pump 61. The charge is pumped through a pipe 62 into a heater 63 wherein it is heated to reaction temperature. The heated charge then passes through pipes 64 and 65 into a reactor 66 that contains a platinum or palladium series metal catalyst of the type described hereinbefore. Hydrogen gas is introduced through a pipe 67 into a compressor 68 and is compressed to the desired reaction pressure. The compressed hydrogen gas then passes through a pipe 69 and a heater 70 and thence through pipe 65 into the reactor 66, wherein the hydrogen in admixture with the hydrocarbon charge is contacted with the catalyst. The total eflluent fro'm the reactor, including hydrogen, gaseous products and liquid products, are withdrawn from the reactor 66 through a pipe 71. If the temperature of the effluent is sufliciently high for the second-pass, the ellluent can be passed through pipes 72 and 73 into a second reactor 74. On the other hand, if it is desired to adjust the temperature of this material to either a higher or lower temperature, this can be do'ne by passing the eifluent through a pipe 75 into a heat exchanger 76 and thence through pipes 77 and 73 into the reactor 74, by suitable manipulation of valves 78, 79 and 80. Make-up hydrogen, if required, can be introduced through a pipe 81 into a compressor 82. The compressed hydrogen is then passed through a pipe 83, a heater 84 and thence through pipes 85 and 86 into pipe 73 wherein it is admixed with the effluent fro'm the first reactor 66. The mixture of hydrogen and etiuent from reactor 66 is contacted with a platinum or palladium series metal catalyst in reactor 74. The catalyst in reactor 74 may be the same type or a different type of catalyst than that contained in reactor 66. The conditions in reactor 74 are adjusted in order to provide the desired degree of conversion in the second-pass operation. Then, the eluent from reactor 74 is withdrawn through a pipe 87 and passed through condenser 88 wherein it is cooled to temperatures at which hydrogen gas can be separated. The thus-cooled efliuent is passed via a pipe 89 into' a high pressure gas separator 90, and separated into a liquid fraction and gaseous fraction. The gaseous material, comprising substantially pure hydrogen, can be recycled to either pass of the process through pipes 91 and 92, and/or through pipes 91 and 93, by suitable manipulation of valves 94, 95 and 96. The liquid pro'duct obtained in the gas separator is withdrawn through a pipe 97 and passed into a depressuring zone 105, and thence through a pipe 97a into a fractionator 98. In the fractionator 98 suitable fractionation is effected. The

dry gas is removed through a pipe 99. Butanes. are removed through a pipe 100 and can be used fo'r gasoline blending. The C54- light naphtha, which is also used for gasoline blending, is withdrawn through a pipe 101. The heavy naphtha fraction is withdrawn through a pipe 102. This fraction will `ordinarily be subjected to a y reforming operation. A No. 2 fuel oil or diesel fuel is Withdrawn through a pipe 103.

It will be noted that the process described in Fig. 6 differs from the process described in Fig. 5,in that the two-pass operation is effected without intermediate separation of the light products. The many advantages of such an operation will be at once apparent. Thus, for example, the cost of intermediate gas separation and fractionation equipment is eliminated. indeed, in the case where the' effluentfrom the first-pass reactor is at a temperature that is sufficiently high for the seconrhpass cracking, even the intermediate heater will be eliminated.v It has Vbeen found that this type of operation can be carried out without substantial. degradation ofl the light hydrocarbon products from theY first-pass when they are passed through the second stageV of the reaction, as the second stage reactor requires a lower temperature than that used in the first stage for the same vamount of conversion. In this embodiment, as in the other the results are more advantageous than those that can be. obtained by cracking in a once-through operation.

The aforedescribed embodiments of the process of this invention have been illustrated by the use of two reactors. Itl must be understood, however, that a series of. two or more reactors, with or without intermediate fractionation steps, can be used in place of the single, second stage reaction zone shown in Figs. and 6. In other words, the process can be effectively carried out in two three,Y four or more reaction stages. Additionally,

Y the process can be carried out by using a plurality of separate catalyst beds encased ,within a single reactor. Although the product separation in the process has been illustrated with a combination involving a high pressure gasY separator, a depressuring zione, and a fractionator, the invention is not limited thereto. There can be employed any of the various means wellY known to those skilledA in the art for recovering reactor efliuents and separating them into fractions.

Any type of reaction zones adapted to carrying out exothermal cracking reactions in the presence of hydro'- gen under superatmospheric pressures can be used in the process of this invention. Thus, for example, the reaction zones can be equipped with coils to remove excess catalyst heat by means of heatV exchange uids or by means ofindirect heat transfer with reactor inlet streams. P`r'eferably however, the process is carried out in adiabatic reaction zones. The cracking process can be operated using a iixed bed of catalyst, or a moving bed of catalyst wherein therhydrocarbon flow can be concurrent or countercurrent to the catalyst ow. A iiuid type variesV between about 100 pounds per square inch gauge and about 2500 pounds per square inch gauge, preferably, however, between about 350 and about 2000 pounds per square inch gauge. The liquid hourly space velocity,

i.e., the liquid volume of hydrocarbon per hour per vol- `urne of catalyst varies between about 0.1 and 10, preferably, between about 0.1and 4. Generally, the molar ratiol of hydrogen to hydrocarbon charge varies between about 2 Vand about80, preferably, betweenA about 5 and about 50, as there is a net consumption of hydrogen in the cracking process. In general, the cracking conditions in the subsequent cracking stages of the process of this invention will" be milder than those inthe iirst cracking stage. Thus, for example, to effect about the same. amount of conversion in each stage, the cracking temperature employed in the second stage reactor will be less than that used in the first stage reactor.

l After the cracking catalyst has been in service for a substantial period of time, reactivation of the catalyst may be'necessary. This is accomplished readily by contacting the catalyst with air, or other oxygen-containing gases, at elevated temperatures in order to burn carbonaceous deposits from the catalyst. Generally, regeneration is effected at temperatures of about. 700 F. to 950 F., commencing with a gas of low oxygen con.- tent and gradually increasing the oxygen concentration throughout the regeneration period, which may last. from about 6 hours to about 24 hours.' It is important to maintain the regeneration temperature below 1100 F., as higher temperatures tend to impair the catalyst activity. The regenerated catalyst is then treated with hydrogen at temperatures of about 900 F. toabout 1000 F. for 2 to 10 hours to complete reactivation.

Combination process As has been mentioned hereinbefore, the heavy naphtha obtained in the aforedes'cribed multiple-pass operation is readily reformed into high octane gasoline. Indeed, it has been found that when the multiple-pass cracking process is combined with a hydroforming process, greater overall yields of high octane gasoline arev obtained than can be achieved by a combination of cracking andreformingy that involves only a once-through cracking operation. A suitable owlsheet for such a process is set forth in Fig. 7.

The crackingV stages in the schematic arrangement set forth -in Fig. 7 are substantially the same as those described hereinbefore in conjunction with Figs. 5 and 6. In order to effect a high yield of high octane gasoline in the process, however, it is essential that the amount of cracking effected in each' cracking stage rnust be less than 50 percent conversion into products boiling at temperatureslower than about 390 F. (100-recycle). The hydrocarbon charge stock is introduced into the process through a pipe and pumped into a heater 111 through a pipe 112 by means of a pump 113. The heated charge is then passed into a reactor 116 through pipes 114v and 115. Hydrogen is introduced into the system through a pipe 117 and compressed by a compressor 118 into a pipe 119. The hydrogen is heated in a heater 120 and then passed into reactor 116 through pipe 116. h1' the reactor 116 partial (less than 50v percent) conversion into products boiling at temperatures below about 390 F. is effected in the presence of a platinum or palladium type catalyst. The effluent from the reactor is removed through a pipe 12.1.

Substantially as described in conjunction with Fig. 5, the eifluent can be passed through a condenser 122 and thence through a pipe 123 into a high pressure gas VVseparator 124, by suitable manipulation of valves 125 and 126. In the gasv separator 124, substantially pure hydrogen gas is removed and recycled to the process through a pipe 127. The liquid product from the separator 124 is passed through a pipe 129 into a depressuring zone 182, and thence through pipe 129e into a fractionator 130, wherein separation of dry gas, butanes, C5-jlight naphtha, heavy naphtha, and products boiling higher thanabout 390 F. is eifected. The heavy naphtha is removed from the fractionator 130 through a pipe 131 and passed through a pipe 132. into a reforming unit 133. This latter operation is described hereinafter. The product boiling at temperatures higher than about 390 F. is removed from the fractionator through a pipe 1340 and 15 t passed to the second stage operation. This material is pumped by a pump 135 through a pipe 136 into a heater 137. The heated material is then passed through pipes 138, 139 and 140 into a second stage reactor 141.

Hydrogen is introduced into the second stage reaction system through a pipe 142 and compressed by means of a compressor 143. The compressed hydrogen is passed through a pipe 144 and through a heater 145 and thence through pipes 139 and 140 into the reactor 141. In the reactor 141, additional (less than 50 percent) conversion of the hydrocarbon into products boiling at temperatures below about 390 F. is effected in the presence of a catalyst of the platinum or palladium series metal similar or dissimilar from that used in the reactor 116. The effluent from the reactor 141 is removed through a pipe 146 and cooled by means of a condenser 147. The cooled effluent is passed through a pipe 14S into a high pressure gas separator 149. The hydrogen gas removed in the gas separator 149 is passed through pipe 150 into a pipe 144 for recycle gas. The liquid product from the gas separator 1149 is transferred through the pipe 152 into a depressurizing zone 183, and thence through a pipe 152a into a fractionator 153. In the fractionator separation of dry gas, butanes, C-llight naphtha, and heavy naphtha is effected. The bottoms from the fractionator constitute No. 2 fuel oil of high quality. This is removed through a pipe 154. The heavy naphtha from the fractionator 153 is removed through a pipe 155 and passed into the reforming unit 133 through the pipe 132. In combination with the heavy naphtha from the fractionator 130 that is introduced through the pipe 131, the total heavy naphtha is reformed in the reforming unit 133.

The reforming operation is carried out in the presence of hydrogen and of a suitabe reforming catalyst. Such catalysts include the metals and the compounds, such as, the oxides or sulfides, of the metals in the left hand column of group VI and in group VIII of the Periodic Classification of the Elements. The catalysts compounds can be used alone, or upon a suitable support. Catalysts that comprise platinum or palladium metal deposited upon supports, such as silica-alumina, silica-zirconia, alumina-boria, alumina, halogen-activated alumina, and the like are particularly suitable. Especially preferredcatalysts comprise platinumV or palladium supported upon silica that has composited therewith alumina in minor amounts that are correlated in accordance with the disclosure of copending application Serial Number 420,092, filed March 3l, 1954, a continuation-in-part of copending application Serial Number 373,516, filed August 1l', 1953.

The reforming operation isgenerally carried out at temperatures of between about 700 F. and about 1000 F., preferably between about 725 F. and about 950 F. The liquid hourly space velocity will vary between about 0.1 and about l0, preferably between about 0.5 and 4. The hydrogen pressure will vary between about 100 pounds per square inch gauge and about 1000 pounds per square inch gauge, preferably between about 350 and about 750 pounds per square inch gauge. The molar ratio of hydrogen to hydrocarbon charge varies between about l and about 20, preferably, between about 4 and about 12.

In the reforming process, there is a net production of hydrogen. The reformate from the reforming unit is removed through a pipe 156 whence it can be transferred to blending operation for making finished gasoline. The net production of hydrogen from the reformer 133 is used to supply all or part of the hydrogen required in the cracking operation. This hydrogen is removed through a pipe 157 and then transferred to pipes 158 and 159 into the hydrogen streams of the cracking stages. Usually the amount of hydrogen produced in the reforming operation is insufficient to supply the requirements of hydrogen consumed in the cracking operation. The

hydrogen available from the reforming operation can be distributed so that all the hydrogen is used in one cracking stage or it can be apportioned between each of the stages. This Vcan be effected by suitable manipulation of valves 160, 161,162 and 163.

This type of combined operation can also be effected using the cracking operation shown in Fig. 6, wherein there is no intermediate fractionation of the products from the first stage. This is effected by passing the total effluent removed from the reactor 116 through the pipe 121 and through the pipes 165 and 140 directly in-4 to the second stage reactor 141. This is effected by suitable manipulation of the valves 125, 126 and 166. If adjustment of the temperature of this effluent is desired, the effluent can be diverted from the pipe through a pipe 167 in a heat exchanger 168 and thence through a pipe 169 back into the pipe 165, by suitable adjust'- ment of valves 170, 171, and 172. The heat exchanger 168 will be a heater or cooling device, dependent upon whether the temperature of the effluent must be adjusted upwards or downwards. When this type of operation is` used, of course, there will be no hydrogen recycle stream through the pipe 127 and this pipe will be closed off by means of a valve 173. If desired, of course, hydrogen can still be supplied to the first stage reaction from the reforming unit through pipe 158. When the system is in operation, the amount of hydrogen obtained in the reforming operation may be sufficient to maintain the cracking operation without the addition of external hydrogen. In that case there will be no additional introduction of hydrogen through the pipes 117 and 142. In the case where the amount of hydrogen obtained from the reforming operation is not sufficient, which is usually the case, enough hydrogen to meet the requirements of the overall process can be introduced through the pipes 117 and 142.

Regardless of whether the cracking operation is effected with or without intermediate fractionation, the butanes and the C54- light naphtha fractions are used in blending finished gasoline. The butanes are removed from the fractionator 130 (if it is employed) through the pipe and from the fractionator 155 through the pipe 177 and passed into a gasoline blending zone 179. The C5| light naphtha fractions are removed from the fractionator 130 (if used) through the pipe 176 and from fractionator 153 through the pipe 178, and transferred to the blending zone 179. The reformate from the reforming zone 133 is transferred to the gasoline blending zone 179 through the pipe 156.

In the gasoline blending zone 179, the high octane reformate from the zone 133 is blended with butanes and with C5+ light naphtha, in suitable proportions to produce a finished gasoline of the desired Reid vapor pressure and octane number. In the operation just described, greater overall eiciency is obtained than can be obtained when the cracking stage is carried out in a once-through operation. The amounts of butanes and of C54- light naphtha are sufiicient for blending purposes, but there will be no substantial excess amounts of these materials. The heavy naphtha is obtained in greater amounts in the present process. In addition, as the process in each stage is controlled to effect less than 50 percent conversion into products boiling below about 390 F., the heavy naphtha is less paraffinic than that obtained in a once-through operation. As is well known to those familiar with the art, this means that the heavy naphtha from the process of this invention is more readily reformed and gives higher yields of high octane reformate. Accordingly, as compared with operations using a oncethrough cracking operation, the overall process of this invention results in decreased losses from dry gas and excess butanes and C5+ light naphtha, and increased yields of high octane reformate as a result of a higher 11F-VY` naphtha production that is reformed to give-a larger percent yield of high octane reformate. It will 'conversion per pass in the cracking operation must be less than 50 percent conversion into products boiling at temperatures lower than about 390 F. The temperature and the liquid hourly space velocity lare the main factors alecting the degree of conversion into products boiling at temperatures lower than about 390 F. The primary elect of varying the hydrogen to hydrocarbon molar ratio and :the hydrogen pressure is upon the amount of coke produced. Even at lower pressures and/ or at lower hydrogen-to-hydrocarbon molar ratios, both Ias specified herein, which conditions favor increased coke formation, there is very little coke produced. Therefore, the variables that must be correlated to control the amount of cracking that is effected per pass are the temperature and the liquid hourly space Velocity. Such correlation is discussed fully in copending application Serial Number 418,166, filed March 23, 1954. In View of that discussion, such correlation can be readily established by those skilled in the art.

Although the present invention has been described with preferred embodiments, it is to be understood that modiications and variations vmay be resorted to Vwithout departing from the spirit and scope of this invention, as those skilled in the art will readily understand. Such Variations and modiications'areconsidered to be withinVv the purview and scope of the appended claims.

What is claimed is: y Y

f1. A process for hydrocrackinghydrocarbon fractions in a manner which produces lower dry gas yield, lower C-llight naphtha yield, higher heavy naphtha yield and a fuel oil with a higher diesel index as compared to once-through operations, vwhich comprises: contacting, 'in a rst reaction zone, a hydrocarbon fraction having Van initial boiling point of at least about 400 F., a 50 percent point of at least 500 F. and an end boiling point of at least about A600 F., with a catalyst comprising between about 0.05 percent and about 20 percent, by weight of the catalyst, of at least one metal of the platinum and palladium series deposited upon a synthetic composite of oxides of lat least two metals of groups IIA,

18 IIIB and IV of the Periodic Arrangement of the Elements having an activity index of at least 25 in the presence of hydrogen and under hydrocracking conditions which effect a portion of the desired ultimate conversion of vsaid hydrocarbon fraction into products boiling below 390 F. with a net consumption of hydrogen by said fraction; contacting at least the portionl of the eiuent of said rst reaction zone which boils` above 390 F. with a catalyst of the same type used in the first reaction zone in a second reaction zone under hy drocracking conditions milder thanthose of the first reaction zone, which conditions include a temperature subst'antially below the temperature used in the rst reaction zone, to effect the remainder of the desired conversion to products boiling below 390 F. with a net consumption of hydrogen in said second reaction zone and the reaction conditions in both of said zones including temperatures within the range about 400 to about 825 F., hydrogen pressures within the range about 100 to 2500 pounds per square inch gauge, a molar ratio of hydrogen'to hydrocarbon charge within the range about 2 to about I80 and a liquid hourly space velocity within the range about 0.1 to about 10.

2. The process Vof claim 1 further limited to the separation of material boiling below 390 F. from the ellluent of the first reaction zone and passing the remainder of said eluent boiling above 390 F. to said second reaction zone.

`3. The process of claim 1 further limited to separating a heavy naphtha from the products of the hydrocracleing boiling below 390 F. and catalytically reforming said naphtha to produce a' high octane gasoline.

4. The process of claim 1 further limited to the degree vof conversion to products boiling below 390 F. in both of said reaction zones being less than percent.

References Cited in the le of this patent UNITED STATES PATENTS Pier .Aug. 8, 1933 vUTE-TED STATES PATENT OFFICE CERTIFICATE OF CORRECTION Patent No.,v 2,945,800 'Y July 19.1960

Frank G. Ciapet'ta e1; al..` E

It is herebir certified that error appears in th`e-prnted Specification of the above ynumbered patent requiring correction and 'that the, saidvLe'tbers Patent should read as'correcbed below.

Column 8,k line 4L, for platform read e# platinum we Signed and sealedv this 10th day of January"l9,6l

(SEAL) Attest: .l KARL H0 AXLINE ROBERT C. WATSON Attesting Officer Commissioner of Patents

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Referenced by
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US3118830 *Mar 8, 1961Jan 21, 1964Texaco IncHydroconversion of hydrocarbons
US3125502 *Feb 1, 1961Mar 17, 1964California Research Corposcott
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Classifications
U.S. Classification208/59
International ClassificationC10G65/00, C10G47/14, C10G65/10, C10G47/00, C10G63/00, C10G63/04
Cooperative ClassificationC10G63/04, C10G47/14, C10G65/10
European ClassificationC10G63/04, C10G65/10, C10G47/14