US3671420A

US3671420A - Conversion of heavy petroleum oils

Info

Publication number: US3671420A
Application number: US101444A
Authority: US
Inventors: Raymond F Wilson; Reese A Peck; Frank E Guptill Jr
Original assignee: Texaco Inc
Current assignee: Texaco Inc
Priority date: 1970-12-24
Filing date: 1970-12-24
Publication date: 1972-06-20
Anticipated expiration: 1989-06-20
Also published as: SE372035B; ZA717670B; ES397401A1; FR2118919A1; BE776823A; NL7117718A; DE2149370A1; DE2149370C3; AU3634671A; CA960982A; BR7108434D0; DE2149370B2; FR2118919B1; GB1346336A; IT944220B

Abstract

Residue-containing petroleum oils are converted into lighter products by a combination of catalytic hydrocracking and catalytic cracking.

Description

United States Patent Wilson et al. 1 June 20, 1972 [54] CONVERSION OF HEAVY PETROLEUM [56] Refenences Cited I S 0 L UNITED STATES PATENTS [72] 2 2 g? g mfgy 3,098,029 7/1963 Snyder ..208/6l 1 3,186,935 6/1965 Vaell ..208/59 [73] Assignee: Texaco Inc., New York, N.Y. 3,211,641 1 Halik ----2 8/59 3,448,037 6/1969 Bunn et al ..208/l64 [22] Filed: Dec. 24, 1970 Primary Examiner-Delbert E. Gantz [21] 101444 AssistantExaminer-G. E. Schmitkons Related Application Data Attorney-Thomas H. Whaley, Carl G. Ries and Robert Knox,

Jr. [63] Continuation-impart of Ser. No. 811,604, March 28,

1969, Pat. No. 3,607,723. [57] ABSTRACT Residue-containing petroleum oils are converted into lighter [52] U.S.Cl ..208/61,208/ 164 products y a combination of catalytic hydrocracking and [51 Int. Cl. ..Cl0g 13/00, ClOg 37/02 catalytic cracking [58] Field of Search ..208/59, 61, 164

9 Claims, No Drawings CONVERSION OF HEAVY PETROLEUM OILS This application is a continuation-in-part of our copending application, Ser. No. 811,604 filed Mar. 28, I969, now US. Pat. No. 3,607,723.

This invention relates to the catalytic treatment of heavy hydrocarbon materials and more particularly to a process which produces substantially complete conversion of said heavy hydrocarbon materials to lower boiling hydrocarbons and selectivity in such conversion to lower boiling hydrocarbons which boil within a particularly preferred boiling range.

Motor fuel, diesel fuel and jet fuel are, for the most part, the most valuable products obtained from petroleum. Consequently the petroleum industry is geared to produce maximum amounts of these products. To this end, crude petroleum is distilled to obtain these desired fractions and that portion of the distillate boiling above the desired range is subjected to hydrocracking or to cracking to convert it to lower boiling material.

However, the still residue is a heavy hydrocarbon oil rich in tar and asphalt and having a relatively high concentration of metals. Attempts to convert still residues such as a vacuum residuum into lighter materials by means of catalytic processes have not been particularly successful as the tar and asphalt deposit on the catalyst producing a coke layer on the catalyst preventing contact of the catalyst and oil. In addition in many cases the metals will deposit on the catalyst causing its deactivation. As a result, the most popular method for converting residua to lighter materials is coking, in which process the oil is heated and retained at elevated temperature until a substantial portion thereof is converted to coke and the balance to a lighter liquid. However, disposal of the coke so formed can present a problem. In fact, much of the still residue produced in petroleum refineries is sold as residual fuel" but even this is no longer a suitable use because of its high sulfur content.

It is therefore an object of this invention to provide a process for the catalytic conversion of residue containing hydrocarbon oils into valuable lighter products. Accordingly, our invention provides a process for the conversion of a residuecontaining petroleum fraction into lighter products which comprises maintaining in a hydrocracking zone a first catalytic zone below and a second catalytic zone above a point of entry into said hydrocracking zone, introducing the residuecontaining petroleum fraction through said point of entry into said hydrocracking zone at a temperature between about 600 and 850 F. and a pressure between about 500 and 5,000 psig, introducing hydrogen into said first catalytic zone to flow upwardly countercurrent to a portion of said residue-containing petroleum fraction at a rate of at least 3,000 SCF per barrel of residue-containing petroleum fraction sufiicient to maintain liquid hydrocarbon in said second catalytic zone, separately recovering product from said first and second catalytic zones, combining at least a portion of the product from said second catalytic zone with a gas oil fraction and contacting the mixture at a temperature between 850 and l,l F. with a fluidized cracking catalyst and combining the product from said first catalytic zone with catalytically cracked cycle gas oil and contacting the mixture with a cracking catalyst at a temperature between about 800 and l,000 F.

The process of our invention may be used for the treatment of residue-containing fractions such as atmospheric residua,

vacuum residua, visbreaker bottoms, whole crude such as San Ardo Crude, shale oil, tar sand oil and the like.

It has now been found that substantial conversion of heavy hydrocarbon materials to lower boiling hydrocarbons can be accomplished in a split flow hydrocracking process which comprises introducing a heavy hydrocarbon charge stock in downward flow into a hydrocracking catalyst zone, said catalyst zone comprising a first catalyst zone below and a second catalyst zone above the point of entry of the heavy hydrocarbon charge stock, introducing hydrogen into said first catalyst zone in countercurrent relationship to the flow of said heavy hydrocarbon charge stock, maintaining a lower boiling liquid in the second catalyst zone, recovering a high boiling effluent from the first catalyst zone and recovering lower boiling hydrocarbons from the second catalyst zone.

In carrying out the'process of this invention, the heavy hydrocarbon charge stock is introduced into a hydrocracking catalyst zone herein defined to include a first catalyst zone in downflow relationship to the downward flow of the heavy hydrocarbon charge stock, a second catalyst zone above the point of entry of the heavy hydrocarbon charge stock and in upfiow relationship to the lower boiling hydrocarbons which proceed from the first catalyst zone into the second catalyst zone. By the use of the tenn downward flow" is meant that the heavy hydrocarbon charge stock proceeds in downflow relationship to the first catalyst zone. By the use of the term above" in reference to the second catalyst zone is meant only that the second catalyst zone is in upfiow relationship to the flow of the hydrogen containing gas and in upfiow relationship to the volatile hydrocarbon and entrained liquid hydrocarbon which proceed from the first catalyst zone into a second catalyst zone. The word above" is used to define a flow relationship with the first catalyst zone, which relationship provides for the flow of hydrogen, volatile hydrocarbons and entrained lower boiling liquid hydrocarbons from the first catalyst zone in countercurrent relationship with the downward flow of the heavy hydrocarbon charge stock into a second catalyst zone. Thus the second catalyst zone can be located directly in a space dimension above the first catalyst zone such as when the first and second catalyst zone are present in a vertical reactor with an intermediate point of entry for the heavy hydrocarbon charge stock. However this invention contemplates that the second catalyst zone can be present as a separate reactor which is connected to the first reactor by conduit means, although it is preferred in carrying out the process of this invention to use a vertical reactor wherein the first catalyst zone and second catalyst zone are present in the same reactor. Within the first and second catalyst zone is a catalyst which has hydrocracking activity under process conditions of temperature, pressure and space velocity which are utilized during the process. In addition, the catalyst in the first catalyst zone can be either. the same or different than the catalyst-present in the second catalyst zone.

The heavy hydrocarbon charge stock upon entry to the catalyst zone proceeds downwardly in downflow relationship tothe first catalyst zone. Hydrogen is introduced into the first catalyst zone at or near the lower extremity and/or at intermediate points in said first catalyst zone in countercurrent relationship to the hydrocarbon flow through the first catalyst zone and in upfiow relationship to the second catalyst zone, the volatile hydrocarbons and the lower boiling liquid hydrocarbons hereinafler referred to as liquid proceed into the second catalyst zone. The volatile hydrocarbons and the liquids which are present in the second catalyst zone proceed from the second catalyst zone and are recovered by conventional means such as by cooling of the hydrocarbon vapors and liquid. The hydrogen which proceeds from the second catalyst zone can then be recycled together with fresh hydrogen in the first catalyst zone. In addition, hydrogen optionally can be blended with the heavy hydrocarbon charge stock and introduced at ambient temperature or higher such as temperatures up to hydrocracking temperatures into the catalyst zone.

As stated above, liquid is maintained in the second catalyst zone. In general, a liquid is maintained in the second catalyst zone by the rate of introduction of hydrogen into the first catalyst zone by any of the means set forth above for the introduction of hydrogen. In order to maintain liquid in the second catalyst zone utilizing hydrogen, it has been found that hydrogen gas rates of at least 3,000 SCF per barrel of charge preferably from 3,000 SCF per barrel up to about 25,000 SCF per barrel are required in the first catalyst zone. The hydrogen need not be pure and gases containing more than about 65 volume percent hydrogen may be used. In this connection, the term hydrogen" is also intended to include dilute hydrogen, reformer by-product hydrogen, hydrogen produced by the partial oxidation of hydrocarbon materials followed by shift conversion and electrolytic hydrogen.

The hold-up of the liquid hydrocarbon charge stock in the first catalyst zone can be varied somewhat by varying the upward flow of hydrogen. In general it is preferred to have high liquid hold-up, that is a hold-up of hydrocarbon charge stock which provides for maximum catalytic effectiveness for the conversion of the charge stock to lower boiling hydrocarbons.

The lower boiling liquid which is maintained in the second catalyst zone in general is derived from the heavy hydrocarbon material, and in general is a lower boiling hydrocarbon which is present initially in the heavy hydrocarbon charge stock and/or which is formed during the process. In general the liquid material has a boiling point below 850 F. It is preferred that the liquid which is present in the second catalyst zone have at least 90 percent by weight of the liquid boiling below 850 F. more preferably at least about 97 percent by weight boiling below 850 F.

The first stage of the process of this invention is utilized for the hydrocracking of heavy hydrocarbon charge stocks which term hydrocracking is herein defined to mean destructive hydrogenation in which a substantial portion of the product boils at a temperature below the initial boiling point of the charge heavy hydrocarbon material. In general percent conversions by weight per single pass of the 850 F.+ material of the charge stock varies from about to 80 percent more preferably from about to 60 percent. The hydrocracking conditions as to pressure, temperature and space velocity can be varied over a wide range, the conditions utilized being those which in combination produce substantial conversion of the heavy hydrocarbon charge stock to lower boiling hydrocarbons.

The first and second catalyst zone conditions that are utilized in the split flow process of this invention are in general temperatures of from about 600 F. to about 850 F., preferably 725 to 840 F.; pressure of from about 500 to about 5,000 psig, preferably 1,500 to 2,000 psig and liquid hourly space velocities of from about 0.05 to about 10, preferably 0.25 to 2.5, volumes of feed per volume of catalyst per hour.

In general it is preferred to have approximately the same conditions in the first and second catalyst zones although the gas rates in the first and second catalyst zones will differ depending upon the amount of hydrogen which is blended together with the heavy hydrocarbon charge stock prior to the introduction into the catalyst zone and/or hydrogen consumed in the process. Thus hydrogen gas rates in the second catalyst zone may be different than the hydrogen rates in the first catalyst zone. In general the liquid hourly space velocity in the second catalyst zone will be greater than that in the first catalyst zone. In addition as in the case where the catalyst zones are not present in the same reactor, temperature and pressure can be different.

The hydrocracking catalyst utilized for the conversion of the aforementioned hydrocarbon charge stocks can be crystalline metallic alumino-silicate zeolite, having a platinum group metal (e.g. platinum or palladium) or an iron group metal alone or in conjunction with a Group VI metal, their compounds and mixtures thereof eg cobalt oxide and molyb- 1 denum oxide or nickel sulfide and tungsten sulfide deposited thereon or composited therewith. These crystalline zeolites are characterized by their highly ordered crystalline structure and uniformly dimensioned pores, and have an alumino-silicate anionic cage structure wherein alumina and silica tetrahedra are intimately connected to each other so as to provide a large number of active sites, with the uniform pore openings facilitating entry of certain molecular structures. It has been found that crystalline alumino-silicate zeolites, having effective pore diameter of about 6 to 15, preferably 8 to 15 angstrom units, when composited with the platinum group metal, and particularly after base exchange to reduce the alkali metal oxide (e.g. Nat-,0) content of thezeolite to less than about 10 wt. preferably less than 2.0%, are effective hydrocracking catalysts. Advantageously, the support will also contain at least one amorphous inorganic oxide such as silica,

alumina, magnesia or zirconia or a mixture thereof such as silica-alumina. Such composite supports preferably contain about 15-45 percent zeolite.

Altemately, the catalyst support may be totally amorphous inorganic oxide. Suitable such carriers or supports include acidic supports such as: silica-alumina, silica-magnesia, and other well-known cracking catalyst bases; the acidic clays; fluorided alumina; and mixtures of inorganic oxides, such as alumina, silica, zirconia, and titania, having sufficient acidic properties providing high cracking activity.

The same or different catalysts may be used in each catalyst zone. Suitably the catalyst in the first zone contains an amorphous support and the catalyst in the second zone contains a crystalline zeolite of low alkali metal content in the support.

Hydrogen is separated from the effluent from the second catalytic zone and if desired may be recycled'to the first catalytic zone with or without purification for the removal of compounds such as hydrogen sulfide and/or ammonia. Lower boiling hydrocarbons eg those boiling up to about 525-550 F. are also removed from the second catalytic zone effluent and the balance is subjected to catalytic cracking as is the effluent from the first catalytic stage. Since the effluent from the second catalytic zone or overhead is high in saturates and the effluent from the first catalytic zone or bottoms is high in aromatics, they are subjected to different conditions of catalytic cracking. This is done effectively with a zeolite cracking catalyst in a fluid catalytic cracking unit comprising a reactor, a regenerator and at least two elongated reaction zones or risers where the reactor contains a dense phase and a dilute phase of the catalyst.

In the simplest embodiment of this invention, an FCCU with two risers is employed with the operating conditions in the risers including a temperature of 800-l,l50 F., conversion of 30-80 volume percent and space velocities in the overhead riser and the bottoms riser being 10-100 w/hr/w and 50-200 w/hr/w, respectively.

Both riser cracking and fluidized dense phase cracking may be employed.

In one embodiment the cracking of the overhead and the bottoms is restricted to the risers by discharging the efiluent from both risers into the dilute phase of catalyst in the reactor vessel. The reactor vessel in this case is utilized as a disengaging space with substantially no cracking taking place therein.

In another embodiment, the overhead is subjected to both riser and dense phase cracking while the cracking of the bottoms is limited to its riser. The effluent from the bottoms riser is discharged into the dilute phase of catalyst, the effluent from the overhead riser is discharged into the dense phase of catalyst and the vaporous reaction mixture from the overhead riser is passed through the dense phase of catalyst under catalytic cracking conditions effecting an additional conversion of 5-30 volume percent with the total per pass conversion of the overhead not exceeding volume percent. By adjusting the operating conditions, the conversion in the overhead riser may be lower, equal to or higher than that in the bottoms riser.

In a further embodiment, the overhead is subjected only to riser cracking while the bottoms is cracked in both the riser and the dense phase of catalyst. The effluent from the overhead riser is discharged directly into the dilute phase of catalyst in the reactor vessel, while the effluent from the bottoms riser is discharged into the dense phase of catalyst and passed through this dense phase under catalytic cracking conditions effecting an additional conversion of 5-30 volume percent. The per pass conversion of the bottoms does not exceed 80 volume percent.

In another embodiment, both the overhead and the bottoms are subjected to both riser cracking and dense phase bed cracking by discharging the effluent from both risers into the dense phase of catalyst and passing them therethrough under catalytic cracking conditions to effect an additional conversion of 5-30 percent. In this embodiment the total conversion of all oils passing through the catalytic cracking unit does not exceed 80 volume percent.

In a preferred embodiment, the overhead is subjected only to riser cracking and the bottoms is subjected to both riser and dense phase cracking.

It is also contemplated that other materials may be fed to the'FCCU. For example, a virgin gas oil may be introduced into the riser with the overhead and the unconverted oil which ordinarily is recycled to the cracking unit is introduced into a separate riser with the bottoms. In one of the more specific embodiments of our invention a crude oil is fractionated at atmospheric pressure to produce naphtha, kerosene, atmospheric gas oils and an atmospheric residuum, the atmospheric residuum is subjected to split flow hydrocracking, the atmospheric gas oil is subjected to catalytic cracking, the overhead from the split flow hydrocracking is combined with the atmospheric gas oil as fresh feed to the catalytic cracking zone and the unconverted feed is combined with the bottoms from the split flow hydrocracking and catalytically cracked under conditions such that the conversion of the lighter material is at least as great as that of the heavier material and may be as much as 30 percent more.

The catalyst employed in the instant invention comprises a large pore crystalline aluminosilicate customarily referred to as a zeolite and an active metal oxide, as exemplified by silicaalumina gel or clay. The zeolites employed as cracking catalysts herein possess ordered rigid three-dimensional structures having uniform pore diameters within the range of from about 5 to about 15 A. The crystalline zeolitic catalysts employed herein comprise about 1 to 25 wt. zeolite, about to 50 wt. alumina and the remainder silica. Among the preferred zeolites are those known as zeolite X and zeolite Y wherein at least a substantial portion of the original alkali metal ions have been replaced with such cations as hydrogen and/or metal or combination of metals such as barium, calcium, magnesium, manganese or such rare earth metals, for example. cerium, lanthanum, neodymium, praseodymium, samarium and yttrium.

As contemplated herein the overhead and bottoms are introduced into elongated reaction zones which are operated to effect a lower conversion of the bottoms stream. In its simplest form, a two riser FCCU is employed. The operating conditions for both the overhead riser and the bottoms riser include an operating temperature of 800-l,150 F., preferably 840l, 000 F. and a conversion per pass of 30-80 percent, preferably 40-75 percent. Other operating conditions within the risers include a residence time of 2-20 seconds, preferably 3-10 seconds and a vapor velocity of -50 ft/sec, preferably -40 ft/sec. The space velocity in the overhead riser is 10-100 w/hr/w, preferably 40-90 w/hr/w and the space velocity in the bottoms riser is 50-200 w/hr/w, preferably 75-150 w/hr/w. The conversion per pass in the bottoms riser is 0-30 percent lower than the conversion in the overhead riser with the overall conversion in the overhead riser not exceeding 80 volume percent.

Where the embodiment employed includes additional cracking of either effluent in the dense phase of catalyst in the reactor the operating conditions within the dense phase include a temperature of 800-l,150 F., a vapor velocity of 0.5-4 ft/sec. preferably 1.3-2.2 ft/sec and a space velocity of l-40 w/hr/w, preferably 3-25 w/hr/w. The vaporout reaction products from a riser which passes through the dense phase of catalyst obtains a further conversion of 5-30 volume percent.

In operation of our process it has been found desirable to operate the hydrocracking stage under such conditions that the end point of the overhead or second catalytic zone effluent is about 850 F. and correspondingly the initial boiling point of the bottoms or first catalytic zone effluent is about 850 F. Of course it will be appreciated that because of the solubility of hydrocarbons in heavier hydrocarbons there is no sharp demarcation and there will be an overlap between the initial and end boiling points.

One of the features of our process is that in the hydrocracking stage under similar reaction conditions our split flow operation will produce six times as much 850 F. and lighter material as will a conventional process in which both the charge and hydrogen pass downwardly through the catalyst bed. Another feature is that the overhead from the hydrocracking stage is high in saturates whereas the bottoms product is high in aromatics. In conventional downflow hydrocracking there is only one product. Furthermore, even if the product from conventional downflow hydrocracking is fractionated, there is no aromatic-saturate separation equivalent to that obtained by the split-flow hydrocracking stage of our process.

The fact that the overhead and bottoms of the split flow hydrocracker have different characteristics means that each can be cracked separately under conditions most suitable for that fraction. It has also been found that the overhead from the split flow hydrocracking of atmospheric residuum behaves like virgin gas oil when subjected to catalytic cracking whereas the split flow hydrocracker bottoms is more aromatic and requires moresevere cracking conditions. In this respect the bottoms product resembles unconverted gas oil which is recycled to the catalytic cracking unit.

Another feature of our invention is that by introducing the bottoms product into the catalytic cracking unit, considerably more carbon than usual is introduced into the catalyst bed thereby permitting greater deposition of carbon on the catalyst which in turn permits operation with a regenerated catalyst having a carbon level up to about 4.0 weight percent. By operating in this manner it is possible to burn off the deposited carbon from the catalyst under carefully controlled conditions to provide a catalyst containing from 0.2 to 2.5 weight percent retained coke and have the catalyst at a temperature between about 1,100 and 1,250 F. for reintroduction into the catalytic cracking zone. The presence of the retained carbon on the regenerated catalyst serves to prolong the life of the catalyst as the metallic components of the bottoms product become incorporated in the carbon layer and do not affect the activity and selectivity of the catalyst.

Another feature of our process is that the more easily cracked feed to the cracking stage can be introduced and reacted separately from the more difficultly cracked material under less severe conditions thereby avoiding overcracking with the undesirable production ofgases obtained in conventional processes where the feed is a single mixture of several streams or where the only difference between two or more feeds lies in the boiling range and not in the type of hydrocarbons in the feeds.

The following examples are for illustrative purposes only and should not be considered as limiting the invention in any manner.

EXAMPLE I In this example, a South Louisiana reduced crude having an API Gravity of 2 l 1, a sulfur content of 0.48 wt. and a Conradson Carbon Residue of4. 14 wt.% is hydrocracked by being passed downwardly with hydrogen through a bed of pelleted catalyst containing 3.4 wt. N10 and 15.9 wt. Mo 0;, supported on alumina. Reaction conditions and yield data for three runs are tabulated below:

EXAMPLE 11 In this example the countercurrent split flow technique of the first stage of the process of our invention is employed using the same charge and catalyst as in Example 1. Reaction conditions and yield data for three runs are tabulated below:

TABLE 2 Split Flow Hydrocracking of So. Louisiana Reduced Crude Run No. 4 5 6 Temp. F. 750 775 775 Pressure. psig 1700 1700 1700 LHSV. Vo/Hr/Vc 1.10 1.07 0.54 H. Rate. SCFB(with feed) 725 750 880 M Rate. SCFB(counter- 7500 7500 12700 current) Product Balance. Wt. Feed Overhead 13.48 19.59 22.05 Bottoms 86.29 79.35 77.10 Total Recovery 99.91 99.22 99.18 Hydrogen Consumption,

SCFB 234 314 552 Overhead Product Gravity. AP1 34.0 35.2 36.3 Sulfur. Wt. 0.040 0.004 0.003 Overhead Product Dist. F. Vol.

lBP-lO 356-473 326-424 310-406 30 589 548 508 50 647 620 586 70 693 676 654 90 735 726 End Point 742 748 668 7r Recovery 93 96 79 Bottoms Product Gravity. AP1 22.1 20.5 22.0 Sulfur. Wt. 7r 0.26 0.20 0.15 Carbon Residue. Wt. 7r 4.00 4.00 3 40 DP] Distillation F.. Wt. Z

lBP-400 0 0 0 400-650 0 0 0 650-850 31.6 30.2 33.0 850-up EP 68.4 69.8 67.0

A comparison of Example 11 with Example 1 shows the superiority of counter-current split flow hydrocracking over conventional downfiow hydrocracking.

EXAMPLE III In this example the products from Runs 3 and 6 are catalytically cracked under substantially identical conditions using a cracking catalyst containing 2.0 wt. cerium, 0.93 wt. lanthanum. 18.0 wt. 7: decationized zeolite Y, 0.94 wt. sodium. 34.1 wt. alumina and the balance silica and having a surface area of 329 m /g and a pore volume of 0.72 cc/g. ln Runs 7. 8 and 9 the feeds are the product from Run 3, the overhead from Run 6. and the bottoms from Run 6 respectively. The extent of cracking is shown by the gas chromatographic analysis of the charges and the products appearing in Table 3.

It will be noted that the residue or bottoms frotn th e current split flow hydrocracking gives substantially the same cracking yield as the entire product from conventional downflow hydrocracking and that the overhead product from the countercurrent split flow hydrocracking gives a high liquid yield boiling below C (400 F.) thereby giving a greater overall cracking yield.

EXAMPLE IV In this example the overhead and bottoms from Run 6 are introduced separately into a dual riser fluid catalytic cracking unit in which the fresh feed is a gas oil blend and in which 430 F.+ product is recycled. The cracking catalyst is the same as that used in Example 111. The proportion of feed to the risers is given below as percent of total feed to the unit.

Riser No. 1 is operated at a temperature of 860 F. and Riser No. 2 at a temperature of 940 F. with substantially all of the cracking taking place in the risers. The conversion of 430 F material into 430 F. material in Riser l is 67.0 volume percent and in Riser 2 49.0 percent giving an overall conversion of of 58.0 percent. The total yield of debutanized naphtha is 49.2 volume percent basis feed having a Research Octane No. (with 3 cc TEL/gal) of96.5.

Various other modifications of the invention as hereinbefore set forth may be made without departing from the spirit and scope thereof, and therefore, only such limitations should be imposed as are indicated in the appended claims.

We claim: I

1. A process for the conversion of a residue-containing petroleum fraction into lighter products which comprises maintaining in a hydrocracking zone a first catalytic zone below and a second catalytic zone above a point of entry into said hydrocracking zone, introducing the residue-containing petroleum fraction through said point of entry into said hydrocracking zone at a temperature between about 600 and 850 F. and a pressure between about 500 and 5,000 psig, introducing hydrogen into said first catalytic zone to flow upwardly countercurrent to a portion of said residue-containing petroleum fraction at a rate of at least 3,000 SCF per barrel of residue-containing petroleum fraction sufficient to maintain liquid hydrocarbon in said second catalytic zone, separately recovering product from said first and second catalytic zones, catalytically cracking the overhead product from said second catalytic zone and separately catalytically cracking the bottoms from said first catalytic zone.

2. The process of claim 1 in which the overhead is catalytically cracked in one riser of a multiple riser catalytic cracking zone and the bottoms product is cracked in another riser of said multiple riser catalytic cracking zone.

3. The process of claim 2 in which a virgin gas oil is also introduced into said one riser and catalytically cracked with said overhead.

4. The process of claim 3 in which the product of the catalytic cracking of said overhead and said virgin gas oil is separated into a naphtha fraction and a fraction boiling above the naphtha range and the latter is catalytically cracked with said bottoms product.

at a temperature between about 800 and 1,000 F. and said other riser is operated at a temperature between about 850 and 1,l00 F.

8. The process of claim 2 in which the catalyst in each of said risers contains between about 0.5 and 4.0 weight percent carbon.

9. The process of claim 1 in which the residue-containing petroleum fraction has an initial boiling point of about 850 F.

Claims

2. The process of claim 1 in which the overhead is catalytically cracked in one riser of a multiple riser catalytic cracking zone and the bottoms product is cracked in another riser of said multiple riser catalytic cracking zone.
3. The process of claim 2 in which a virgin gas oil is also introduced into said one riser and catalytically cracked with said overhead.
4. The process of claim 3 in which the product of the catalytic cracking of said overhead and said virgin gas oil is separated into a naphtha fraction and a fraction boiling above the naphtha range and the latter is catalytically cracked with said bottoms product.
5. The process of claim 3 in which naphtha and kerosene fractions are recovered from the catalytic cracking product of said virgin gas oil and said overhead and that portion of the product boiling above the kerosene range is catalytically cracked with said bottoms product.
6. The process of claim 2 in which the bottoms product after being at least partially catalytically cracked in said riser, is withdrawn therefrom and subjected to additional cracking in the dense phase of a fluidized cracking catalyst bed.
7. The process of claim 2 in which said one riser is operated at a temperature between about 800* and 1,000* F. and said other riser is operated at a temperature between about 850* and 1,100* F.
8. The process of claim 2 in which the catalyst in each of said risers contains between about 0.5 and 4.0 weight percent carbon.
9. The process of claim 1 in which the residue-containing petroleum fraction has an initial boiling point of about 850* F.