IMMOBILISATION OF VANADIUM DEPOSITED ON CATALYTIC MATERIALS DURING THE CONVERSION OF OIL THAT CONTAINS HEAVY METALS AND COKE PRECURSORS

Abstract: A process is disclosed for catalytic cracking a hydrocar­ bon oil feed having a significant vanadium content to produce lighter products. The catalyst, from the cracking step, coated with coke and vanadium in an oxidation state less than +5, is regenerated in the presence of an oxygen-­ containing gas at a temperature high enough to burn off a portion of the coke under conditions keeping the vanadium in an oxidation state less than +5.

Description: IMMOBILIZATION OF VANADIA DEPOSITED ON CATALYTIC MATERIALS DURING CARBO-METALLIC OIL CONVERSION Description Technical Field This invention relates to processes for converting heavy hydrocarbon oils into lighter fractions, and especially to processes for converting heavy hydrocarbons containing high concentrations of coke precursors and heavy metals into gasoline and other hydrocarbon fuels. Background Art The introduction of catalytic cracking to the petroleum industry in the 1930's constituted a major advance over previous techniques with the object of increasing the yield of gaseoline and its quality. Early fixed bed, moving bed, and fluid bed catalytic cracking FCC processes employed vacuum gas oils (VGO) from crude sources that were considered sweet and light. The terminology of sweet refers to low sulfur content and light refers to the amount of material boiling below approximately 1000-102S F. The catalysts employed in early homogenous fluid dense beds were of an amorphous siliceous material, prepared synthetically or from naturally occurring materials activated by acid leaching. Tremendous strides were made in the 1950's in FCC technology in the areas of metallurgy, processing equipment, regeneration and new more-active and more stable amorphous catalysts. However, increasing demand with respect to quantity of gasoline and increased octane number requirements to satisfy the new high horsepower-high corn- pression engines being promoted by the auto industry, put extreme pressure on the petroleum industry to increase FCC capacity and severity of operation. A major breakthrough in FCC catalysts came in the early 1960's with the introduction of molecular sieves or zeo- lites. These materials were incorporated into the matrix of amorphous and/or amorphous/kaolin materials constituting the FCC catalysts of that time. These new zeolitic catalysts, containing a crystalline aluminosilicate zeolite in an amorphous or amorphous/kaolin matrix of silica, alumina, silica-alumina, kaolin, clay or the like were at least 100010,000 times more active for cracking hydrocarbons than the earlier amorphous or amorphous/kaolin containing silicaalumina catalysts. This introduction of zeolitic cracking catalysts revolutionized the fluid catalytic cracking process. Innovations were developed to handle these high activities, such as riser cracking, shortened contact times, new regeneration processes, new improved zeolitic catalyst developments, and the like. The new catalyst developments revolved around the development of various zeolites such as synthetic types X and Y and naturally occurring faujasites; increased thermalsteam (hydrothermal) stability of zeolites through the inclusion of rare earth ions or ammonium ions via ion-exchange techniques; and the development of more attrition resistant matrices for supporting the zeolites. These zeolitic catalyst developments gave the petroleum industry the capability of greatly increasing throughput of feedstock with increased conversion and selectivity while employing the same units without expansion and without requiring new unit construc ton. After the introduction of zeolite-containing catalysts the petroleum industry began to suffer from a lack of crude availability as to quantity and quality accompanied by increasing demand for gasoline with increasing octane values. The world crude supply picture changed dramatically in the late 1960's and early 1970's. From a surplus of light, sweet crudes the supply situation changed to a tighter supply with an ever-increasing amount of heavier crudes with higher sulfur contents. These heavier and higher sulfur crudes presented processing problems to the petroleum refiner in that these heavier crudes invariably also contained much higher metals and Conradson carbon values, with accom panying significantly increased asphaltic content. Fractionation of the total crude to yield cat cracker charge stocks also required much better control to ensure that metals and Conradson carbon values were not carried overhead to contaminate the FCC charge stock. The effects of heavy metals and Conradson carbon on a zeolite-containing FCC catalyst have been described in the literature as to their highly unfavorable effect in lowering catalyst activity and selectivity for gasoline production and their harmful effect on catalyst life. These heavier crude oils also contained more of the heavier fractions and yielded a lower volume of the high quality FCC charge stocks which normally boil below about 10250F and are usually processed so as to contain total metal levels below 1 ppm, preferably below 0.1 ppm, and Conradson carbon values substantially below 1.0. With the increasing supply of heavier crudes, which yield less gasoline, and the increasing demand for liquid transportation fuels, the petroleum industry began a search for processes to utilize these heavier crudes in producing gasoline. Many of these processes have been described in the literature and include Gulf's Gulfining and Union Oil's Unifining processes for treating residuum, UOP's Aurabon process, Hydrocarbon Research's H-Oil process, Exxon's Flexicoking process to produce thermal gasoline and coke, H-Oil's Dynacracking and Phillip's Heavy Oil Cracking (HOC) processes. These processes utilize thermal cracking or hydrotreating followed by FCC or hydrocracking operations to handle the higher content of metal contaminants (Ni-V-Fe-Cu-Na) and high Conradson carbon values of 5-15. Some of the drawbacks of these types of processing are as follows: coking yields thermally cracked gasoline which has a much lower octane value than cat cracked gasoline, and is unstable due to the production of gum from diolefins, and requires further hydrotreating and reforming to produce a high octane pro duct; gas oil quality is degraded due to thermal reactions which produce a product containing refractory polynuclear aromatics and high Conradson carbon levels which are highly unsuitable for catalytic cracking; and hydrotreating requires expensive high pressure hydrogen, multi-reactor systems made of special alloys, costly operations, and a separate costly facility for the production of tydrogen. To better understand the reasons why the industry has progressed along the processing schemes described, one must understand the known effects of contaminant metals (Ni-V-Fe Cu-Na) and Conradson carbon on the zeolite-containing cracking catalysts and the operating parameters of an FCC unit. Metal content and Conradson carbon are two very effective restraints on the operation of an FCC unit and may even impose undesirable restraints on a Reduced Crude Conversion (RCC) unit from the standpoint of obtaining maximum conversion, selectivity and catalyst life. Relatively low levels of these contaminants are highly detrimental to an FCC unit. As metals and Conradson carbon levels are increased still further, the operating capacity and efficiency of an RCC unit may be adversely affected or made uneconomical. These adverse effects occur even through there is enough hydrogen in the feed to produce an ideal gasoline consisting of only toluene and isomeric pentenes (assuming a catalyst with such ideal selectivity could be devised). The effect of increased Conradson carbon is to increase that portion of the feedstock converted to coke deposited on the catalyst. In typical VGO operations employing a zeolitecontaining catalyst in an FCC unit, the amount of coke deposited on the catalyst averages around about 4-5 wt% of the feed. This coke production has been atrributed to four different coking mechanisms, namely, contaminant coke from adverse reactions caused by metal deposits, catalytic coke caused by acid site cracking, entrained hydrocarbons resulting from pore structure adsorption and/or poor stripping, and Conradson carbon resulting from pyrolytic distillation of hydrocarbons in the conversion zone. There has been postulated two other sources of coke present in reduced crudes in addition to the four present in VGO. They are: (1) ad sorbed and absorbed high boiling hydrocarbons which do notvaporize and cannot be removed by normally efficient stripping, and (2) high molecular weight nitrogen-containing hydrocarbon compounds adsorbed on the catalyst's acid sites. Both of these two new types of coke producing phenomena add greatly to the complexity of resid processing. Therefore, in the processing of higher boiling fractions, e.g., reduced crudes, residual fractions, topped crude, and the like, the coke production based on feed is the summation of the four types present in VGO processing (the Conradson carbon value generally being much higher than for VGO), plus coke from the higher boiling unstrippable hydrocarbons and coke associated with the high boiling nitrogen-containing molecules which are adsorbed on the catalyst. Coke production on clean catalyst, when processing reduced crudes, may be estimated as approximately 4 wt% of the feed plus the Conradson carbon value of the heavy feedstock. The coked catalyst is brought back to equilibrium activity by burning off the deactivating coke in a regeneration zone in the presence of air, and the regenerated catalyst is recycled back to the reaction zone. The heat generated during regeneration is removed by the catalyst and carried to the reaction zone for vaporization of the feed and to provide heat for the endothermic cracking reaction. The temperature in the regenerator is normally limited because of metallurgical limitations and the hydrothermal stability of the catalyst. The hydrothermal stability of the zeolite-containing catalyst is determined by the temperature and steam partial pressure at which the zeolite begins to rapidly lose its crystalline structure to yield a low-activity amorphous material. The presence of steam is highly critical and is generated by the burning of adsorbed and absorbed (sorbed) carbonaceous material which has a significant hydrogen content (hydrogen to carbon atomic ratios generally greater than about 0.5). This carbonaceous material is pricipally the high-boiling sorbed hydrocarbons with boiling points as high as 1500-1700 F or above that have a modest hydrogen content and the high boiling nitrogen containing hydrocarbons, as well as related porphyrins and asphaltenes. The high molecular weight nitrogen compounds usually boil above 1025DF and may be either basic or acidic in nature. The basic nitrogen compounds may neutralize acid sites while those that are more acidic may be attracted to metal sites on the catalyst. The porphyrins and asphaltenes also generally boil above 10250F and may contain elements other than carbon and hydrogen. As used in this specification, the term "heavy hydrocarbons includes all carbon and hydrogen compounds that do not boil below about 10250F, regardless of the presence of other elements in the compound. The heavy metals in the feed are generally present as porphyrins and/or asphaltenes. However, certain of these metals, particularly iron and copper, may be present as the free metal or as inorganic compounds resulting from either corrosion of process equipment or contaminants from other refining processes. As the Conradson carbon value of the feedstock increases, coke production increases and this increased load will raise the regeneration temperature; thus the unit may be limited as to the amount of feed that can be processed because of its Conradson carbon content. Earlier VGO units operated with the regenerator at 1150-1250 F. A new development in reduced crude processing, namely, Ashland Oil's "Reduced Crude Conversion Process", as described in pending U.S. applications Ser. Nos. 94,091, 94,092, 94,216, 94,217 and 94,227, all filed on November 14, 1979, can operate at regenerator temperatures in the range of 1350-14000F. But even these higher regenerator temperatures place a limit on the Conradson carbon vlue of the feed at approximately 8, which represents about 12-13 wt% coke on the catalyst based on the weight of feed. This level is controlling unless considerable water is introduced to further control temperature, which addition is also practiced in Ashland's RCC processes. The metal-containing fractions of reduced crudes contain Ni-V-Fe-Cu in the form of porphyrins and asphaltenes. These metal-containing hydrocarbons are deposited cn the catalyst during processing and are cracked in the riser to deposit the metal or are carried over by the coked catalyst as the metallo-porphyrin or asphaltene and converted to the metal oxide during regeneration. The adverse effects of these metals as taught in the literature are to cause nonselective or degradative cracking and dehydrogenation to produce increased amounts of coke and light gases such as hydrogen, methane and ethane. These mechanisms adversely affect selectivity, resulting in poor yields and quality of gasoline and light cycle oil. The increased production of light gases, while impairing the yield and selectivity of the process, also puts an increased demand on gas compressor capacity. The increase in coke production, in addition to its negative impact on yield, also adversely affects catalyst activity-selectivity, greatly increases regenerator air demand and compressor capacity, and may result in uncontrollable and/or dangerous regenerator temperatures. These problems of the prior art have been greatly minimized by the development at Ashland Oil, Inc., of its Reduced Crude Conversion (RCC) Processes described in the copending applications reference above and incorporated herein by reference. The new processes can handle reduced crudes or crude oils containing high metals and Conradson carbon values previously not susceptible to direct processing. It has long been known that reduced crudes with high nickel levels present serious problems as to catalyst deactivation at high metal on catalyst contents, e.g., 500010,000 ppm and elevated regenerator temperatures. It has now been recognized that when reduced crudes with high vanadium levels are processed over zeolite containing catalysts, especially at high vanadium levels on the catalyst, rapid deactivation of the zeolite can occur. This deactivation manifests itself as a loss of zeolitic structure. This loss has been observed at vanadium levels of 1000 ppm by weight or less. This loss of zeolitic structure becomes more ra pid and severe with increasing levels of vanadium and at vanadium levels about 5000 ppm, particularly at levels approaching 10,000 ppm complete destruction of the zeolite may occur. Prior to the present invention, it was believed impossible to operate economically at vanadium levels higher than 10,000 ppm because of this phenomenon. Previously, deactivation of catalyst by vanadium at vanadium levels of less than 10,000 ppm has been retarded by lowering regenerator temperatures and increasing the addition rate of virgin catalyst. Lowering regenerator temperatures has the disadvantage of requiring higher catalyst to oil ratios which increase the amount of coke produced and adversely affect yields. Increasing catalyst addition rates is costly and can result in an uneconomical operation. It has been found that vanadium is especially detrimental to catalyst life. The vanadium deposited on the catalyst under the reducing conditions in the riser is in an oxidation state less than +5. At the elevated temperatures and oxidizing conditions encountered in the regenerator the vanadium on the catalyst is converted to vanadium oxides, in particular vanadium pentoxide. The vanadium pentoxide has a melting point lower than temperatures encountered in the regeneration zone, and it can become a mobile liquid, flowing across the catalyst surface and plugging pores. This vanadia may also enter the zeolite structure, neutralizing the acid sites and, more significantly, irreversibly destroying the crystalline aluminosilicate structure and forming a less active amorphous material. In addition, this molten vanadia can, at high vanadia levels, especially for catalyst materials having a low surface area, coat the catalyst microspheres and thereby coalesce particles which adversely affects their fluidization. Summarv of the Invention In accordance with this invention a process has been provided for converting a vanadium-containing hydrocarbon oil feed to lighter products comprising the steps of con tacting said oil feed under conversion conditions with a cracking catalyst to form lighter products and coke, whereby vanadium in an oxidation state less than +5 is deposited on said catalyst together with coke. The lighter products are separated from the spent catalyst and the catalyst is regenerated by contacting it with an oxygen-containing gas under conditions whereby said coke is burned forming CO and CO2 and said vanadium is maintained in an oxidation state less than +5. This invention, by retaining vanadium in an oxidation state wherein the vanadium has a high melting point, permits the recycle of catalyst to levels of vanadium as high as 10,000 ppm, or even 20,000 ppm or 50,000 ppm. The adverse effects, such as clumping of the catalyst and pore closings brought about by molten pentavalent vanadium, are thus avoided. Inasmuch as the catalyst can withstand a much higher vanadium loading than previously experienced the amount of make-up catalyst is reduced. Brief Description of the Drawings Figs. 1 and 2 are schematic designs of catalyst regeneration and associated cracking apparatus which may be used in carrying out this invention. Best and Various Other Modes For Carrying Out the Invention The invention may be carried out by controlling the regeneration of the spent, vanadium-containing catalyst using several methods, alone or in combination. The objective of these methods is to retain vanadium in a low oxidation state, either by not exposing the vanadium to oxidizing conditions, or by exposing vanadium to oxidizing conditions for too short a time to oxidize a significant amount of vanadium to the +5 state. The concentration of vanadium on the catalyst particles increases as the catalyst is recycled, and the vanadium on the catalyst introduced into the reactor becomes coated with coke formed in the reactor. In one methods of carrying out the invention, the generator conditions are selected to ensure that the concentration of coke is retained at at least a minimum level on the catalyst. This coke may serve either to ensure a reducing environment for the vanadium or to provide a barrier to the movement of oxidizing gas to underlying vanadium. The concentration of coke on the catalyst particles is at least about 0.05 percent and the preferred coke concentration is at least about 0.15 percent. In one method of carrying out this invention, which may be combined with the foregoing method of retaining at least about 0.05 percent coke on the catalyst or may be used to achieve lower concentrations of coke, the regeneration is carried out in an environment which is non-oxidizing for the vanadium in an oxidation state less than +5. This may be accomplished by adding reducing gases such as, for example, CO or ammonia to the regenerator, or by regenerating under oxygen-deficient conditions. Oxygen-deficient regeneration increases the ratio of CO to CO2 and in this method of providing a non-oxidizing atmosphere the CO/CO2 ratio is at least about 0.25, preferably is at least about 0.3, and most preferably is at least about 0.4. The CO/CO2 ratio may be controlled by controlling the extent of oxygen deficiency within the regenerator. The CO/CO2 ratio may be increased by providing chlorine in an oxidizing atmosphere within the regenerator, the concentration of chlorine preferably being from about 100 to 400 ppm. This method of increasing the CO/CO2 ratio is disclosed in copending applications Ser. No. filed March 23, 1981 for "Addition of MgCl2 to Catalyst and Ser. No. filed March 23, 1981 for "Addition of Chlorine to Regenerator", both in the name of George D. Myers. The use of a reducing atmosphere within the regenerator is especially useful in combusting coke in zones where the coke level approaches or is reduced below about 0.05 per cent, and it is preferred to have a CO/CO2 ratio of at least about 0,25 in zones where the coke loading is less than about 0.05 percent by weight. It is especially useful to keep the vanadium in a reduced state under conditions wherein the particles are in contact or in relatively frequent contact with each other. Consequently, it is especially contemplated, in carrying out this method, of maintaining a reducing atmosphere in zones within the regenerator wherein the catalyst particles are in a relatively dense bed, such as in a dense fluidized or settled bed. A reducing gas such as CO, methane, or ammonia may be added to a zone having a dense catalyst phase, such as for example a bed having a density of about 25 to about 50 pounds per cubic foot. In another method of carrying out this invention, a riser regenerator is used as one stage in a multi-stage regenerator to contact the catalyst with an oxidizing atmosphere for a short period of time, such as for example less than about two seconds and preferably less than about one second. The riser stage of the regenerator has the advantage in reducing the carbon concentration to a level less than about 0.15 percent orless than about 0.05 percent, that vanadium, which is no longer protected by a coating of carbon, may not be in an oxidizing atmosphere for a long enough time to form molten +5 vanadium. Further, the low density of the particles in the riser-regenerator, minimizes coalescence of those particles which may have liquid pentavalent vanadia on their surfaces. In the preferred method of using a riser regenerator, the particles are contacted with a reducing atmosphere, such as one containing CO or other reducing gas, after leaving the riser. The particles may then be accumulated, as for example, in a settled bed, before being recycled to contact additional fresh feed. The catalyst particles to be accumulated are contacted with a reducing atmosphere, preferably immediately after leaving the riser and before accumulating in a dense bed of regenerated particles, and in the pre ferred method of carrying out this process the particles are retained in a reducing atmosphere within such dense bed, and in the most preferred method a reducing atmosphere is provided for the particles until about the time they are contacted with fresh feed. The preferred riser regenerator is similar to the vented riser reactor as is disclosed in U.S. Patents 4,066,533 and 4,070,159 to Myers et al which achieves ballistic separation of gaseous products from catalyst. This apparatus has the advantages of achieving virtually instantaneous separation of the regenerated catalyst, now containing some vanadia to which any oxygen present would have access, from the oxidizing atmosphere. In the preferred method of reducing the coke concentration to a level less than about 0.15 and especially to less than 0.058 the catalyst is contacted with a reducing atmosphere, preferably immediately after its separation from the oxidizing atmosphere and most preferably also in collection zones for the regenerated catalyst. This invention may be used in processing any hydrocarbon feed containing a significant concentration of vanadium, and FCC as well as RCC processes are contemplated. It is, however, especially useful in processing reduced crudes having high metal and high Conradson carbon values, and the invention will be described in detail with respect to its use in processing an RCC feed. The carbo-metallic feed comprises or is composed of oil which boils above about 650"F. Such oil, or at least the 6500fit portion thereof, is characteri=ed by a heavy metal content of at least about 4, preferably more than about S, and most preferably at least about 5.5 ppm of Nickel Equivalents by weight and by a carbon residue on pyrolysis of at least about 1% and more preferably at least about 2 by weight. In accordance with the invention, the carbo-metallic fed, in the form of a pumpable liquid, is brought into contact with hot conver sion catalyst in a weight ratio of catalyst to feed in the range of about 3 to about 18 and preferably more than about 6. The feed in said mixture undergoes a conversion step which includes cracking while the mixture of feed and catalyst is flowing through a progressive flow type reactor. The feed, catalyst, and other materials may be introduced at one or more points. The reactor includes an elongated reaction chamber which is at least partly vertical or inclined and in which the feed material, resultant products and catalyst are maintained in contact with one another while flowing as a dilute phase or stream for a predetermined riser residence time in the range of about 0.5 to about 10 seconds. The reaction is conducted at a temperature of about 900C to about 1400 F, measured at the reaction chanber exit, under a total pressure of about 10 to about 50 psia (pounds per square inch absolute) under conditions sufficiently severe to provide a conversion per pass in the range of about 50% or more and to lay down coke on the catalyst in an amount in the range of about 0.3 to about 3% by weight and preferably at least about 0.5%. The overall rate of coke production, based on weight of fresh feed, is in the range of about 4 to about 142 by weight. At the end of the predetermined residence time, the catalyst is separated from the products, is stripped to reprove high boiling components and other entrained or adsorbed hydrocarbons and is then regenerated with oxygen-contan- ing combustion-supporting gas under conditions of time, temperature and atmosphere sufficient to reducethe carbon on the regenerated catalyst to about 0.25% or less. Depending on how the process of the invention is practiced. one or more of the following additional advantages may be realized. If desired, and preferably, the process may be operated without added hydrogen in the reaction chamber. If desired, and preferably, the process may be operated without prior hydrotreating of the feed and/or without other process of removal of asphaltenes of metals from the feed, and this is true even where the carbo-metallic oil as a whole contains more than about 4, or more than about 5 or even more than about 5.5 ppm Nickel Equivalents by weight of heavy metal and has a carbon residue on pyrolysis greater than about 1%, greater than about 1.4% or greater than about 2 by weight. Moreover, all of the converter feed, as above described, may be cracked in one and the same conversion chamber. The cracking reaction may be carried out with a catalyst which has previously been used (recycled, except for such replacement as required to compensate for normal losses and deactivation) to crack a carbo-metallic feed under the above described conditions. Heavy hydrocarbons not cracked to gasoline in a first pass may be recycled with or without hydrotreating for further cracking in contact with the same kind of feed in which they were first subjected to cracking conditions, and under the same kind of conditions, but operation in a substantially once-through or single pass mode (e.g. less than about 15% by volume of recycle based on volume of fresh feed) is preferred. According to one preferred e-,bodiment or aspect of the invention, at the end of the predetermined residence time referred to above, the catalyst is projected in a direction established by the elongated reaction chamber or an extension thereof, while the products, having lesser momentum, are caused to make an abrupt change of direction, resulting in an abrupt, substantially inst n- taneous ballistic separation of products from catalyst. The thus separated catalyst is then stripped, regenerated and recycled to the reactor as above described. According to another preferred embodiment or aspect of the invention, the converter feed contains 6500F+ mater al which has not been hydrotreated and is characterized in part by containing at least about 5.5 parts per million of nickel equivalents of heavy metals. The converter feed is brought together not only with the above mentioned cracking catalyst, but also with additional gaseous material including steam whereby the resultant suspension of catalyst and feed also includes gaseous material wherein the ratio of the partial pressure of the added gaseous material relative to that of the feed is in the range of about 0.25 to about 4.0. The vapor residence time is in the range of about 0.5 to about 3 seconds when practicing this embodiment or aspect of the invention. This preferred' embodiment or aspect and the one referred to in the preceeding paragraph may be used in combination with one another or separately. According to another preferred embodiment or aspect of the invention, the carbo-metallic feed is not only brought into contact with the catalyst, but also with one or more additional materials including particularly liquid water in a weight ratio relative to feed ranging from about 0.04 to about 0.25, more preferably about 0.04 to about 0.2 and still more preferably about 0.05 to about 0.15. Such additional materials, including the liquid water, may be brought into admixture with the feed prior to, during or after mixing the feed with the aforementioned catalyst, and either after or, preferably, before, vaporization of the feed. The feed, catalyst and water (e.g. in the form of liquid water or in the form of steam produced by vaporization of liquid water in contact with the feed) are introduced into the progressive flow type reactor, which may or may not be a reactor embodying the above described ballistic separation, at one or more points along the reactor. While the mixture of feed, catalyst and steam produced by vaporization of the liquid water flows through the reactor, the feed undergoes the above mentioned conversion step which includes cracking. The feed material, catalyst, steam and resultant products are maintained in contact with one another in the above mentioned elongated reaction chamber while flowing as a dilute phase or stream for the above mentioned predetermined riser residence time which is in the range of about 0.5 to about 10 seconds. The present-invention provides a process for the continuous catalytic conversion of a wide variety of carbometallic oils to lower molecular weight products, while maximizing production of highly valuable liquid products, and making it possible, if desired, to avoid vacuum distillation and other expensive treatments such as hydrotreating. The term "oils", includes not only those predominantly hydrocarbon compositions which are liquid at room temperature (i.e., 639F), but also those predominantly hydrocarbon con positions which are asphalts or tars at ambient tempera- ture but liquify when heated to temperatures in the range of up to about 8000F. The invention is applicable to carbo-metallic oils, whether of petroleum origin or not For example, provided they have the requisite boiling range, carbon residue on pyrolysis and heavy metals content, the invention may be applied to the processing of such widely diverse materials as heavy bottoms from crude oil, heavy bitumen crude oil, those crude oils known as "heavy crude" which approximate the properties of reduced crude, shale oil, tar sand extract, products from coal liquification and solvated coal, atmospheric and vacuum reduced crude, extracts and/or bottoms (raffinate) from solvent de-asphalting, aromatic extract from lube oil refining, tar bottoms, heavy cycle oil, slop oil, other refinery waste streams and mixtures of the foregoing. Such mixtures can for instance be prepared by mixing available hydrocarbon fractions, including oils, tars, pitches and the like. Also, powdered coal may be suspended in the carbo-metallic oil. Persons skilled in the art are aware of techniques for demetalation of carbometallic oils, and demetalated oils may be converted using the invention; but it is an advantage of the invention that it can employ as feedstock carbo-metallic oils that have had no prior demetalation treatment. Likewise, the invention can be applied to hydrotreated feedstocks; but it is an advantage of the invention that it can successfully convert carbo-metallic oils which have had substantially no prior hydrotreatment. However, the preferred application of the process is to reduced crude, i.e., that fraction of crude oil boiling at and above 650"F, alone or in admixture with virgin gas oils. While the use of material that has been subjected to prior vacuum distillation is not excluded, it is an advantage of the invention that it can satisfactorily process material which has had no prior vacuum distillation, thus saving on capital investment and operating costs as compared to conventional FCC processes that require a vacuum distillation unit. In accordance with the invention one provides a carbometallic oil feedstock, at least about 705, more preferably at least about 85% and still more preferably about 100% (by volume) of which boils at and above about 6500F. All boiling temperatures herein are based on standard at=ros- pheric pressure conditions. In carbo-metallic oil partly or wholly composed of material which boils at and above about 650"F, such material is referred to herein as 6500F+ material; and 650"F+ material which is part of or has been separated from an oil containing components boiling above and below 6500F may be referred to as a 6500F+ fraction. but the terms "boils above" and "650"F+" are not intended to imply that all of the material characterized by said terms will have the capability of boiling. The carbometallic oils contemplated by the invention may contain material which may not boil under any conditions; for example, certain asphalts and asphaltenes may crack thermally during distillation, apparently without boiling. Thus, for example, when it is said that the feed comprises at least about 70 by volume of material which boils above about 6500F, it should to understood that the 70t in question may include some material which will not boil or volatilize at any temperature. These -non-boilable - materials when present, may frequently or for the most part be concentrated in portions of the feed which do not boil below about 10000F, 10250F or higher. Thus, when it is said that at least about 10%, more preferably about 15% and still more preferably at least about 20% (by volume) of the 6500F+ fraction will not boil below about l0000F or 1025 F, it should be understood that all or any part of the material not boiling below about 10000 or 10250F, may or may not be volatile at and above the indicated temperatures. Preferably, the contemplated feeds, or at least the 6;0of+ material therein, have a carbon residue on pyrolysis of at least about 2 or greater. For example, the Conradson carbon content may be in the range of about 2 to about 12 and most frequently at least about 4. A particularly common range is about 4 to about 8. Preferably, the feed has an average composition characterized by an atomic hydrogen to carbon ratio in the range of about 1.2 to about 1.9, and preferably about 1.3 to about 1.8. The carbo-metallic feeds employed in accordance with the invention, or at least the 6500F+ material therein, may contain at least about 4 parts per million of Nickel Ecuivalents, as defined above, of which at least about 01. ppm is vanadium. Carbometallic oils within the above range can be prepared from mixtures of two or more oils, some of which do and some of which do not contain the quantities of Nickel Equivalents and vanadium set forth above. It should also be noted that the above values for Nickel Equivalents and nickel represent time-weighted averages for a substantial period of operation of the conversion unit, such as one month, for example. It should also be noted that the heavy metals have in certain circumstances exhibited some lessening of poisoning tendency after repeated oxidations and reductions on the catalyst, and the literature describes criteria for establishing "effective metal" values. For example, see the article by Cimbalo, et al, entitled "Deposited Metals Poison FCC Catalyst", Oil and Gas Journal, May 15, 1972, pp 112-122, the contents of which are incorporated herein by reference. If considered necessary or desirable, the contents of Nickel Equivalents and vanadium in the carbometallic oils processed according to the invention may be expressed in terms of "effective metal" values. Notwithstanding the gradual reduction in poisoning activity noted by Cimbalo, et al, the regeneration of catalyst under normal FCC regeneration conditions may not, and usually does not, severely impair the dehydrogenation, demethanation and aromatic condensation activity of heavy metals accumulated on cracking catalyst. It is known that about 0.2 to about 5 weight per cent of "sulfur" in the form of elemental sulfur and/or its corn- pounds (but reported as elemental sulfur based on the weight of feed) appears in FCC feeds and that the sulfur and modi fied forms of sulfur can find that way into the resultant gasoline product and, where lead is added, tend to reduce its susceptibility to octane enhancement. Sulfur in the product gasoline often requires sweetening when processing high sulfur containing crudes. To the extent that sulfur is present in the coke, it also represents a potential air pollutant since the regenerator burns it to SO2 and 503. However, we have found that in our process the sulfur in the feed is on the other hand able to inhibit heavy metal activity by maintaining metals such as Ni, V, Cu and Fe in the sulfide form in the reactor. These sulfides are much less active than the metals them selves in promoting dehydrogenation and coking reactions. Accordingly, it is acceptable to carry out the invention with a carbo-metallic oil having at least about 0.3%, acceptably more than about 0.88 and more acceptably at least about 1.58 by weight of sulfur in the 550 F+ fraction. The carbo-metallic oils useful in the invention may and usually do contain significant quantities of compounds containing nitrogen, a substantial portion of which may be basic nitrogen. For example, the total nitrogen content of the carbo-metallic oils may be at least about 0.05% by weight. Since cracking catalysts owe their cracking activity to acid sites on the catalyst surface or in its pores, basic nitrogen-containing compounds may temporarily neutralize these sites, poisoning the catalyst. However, the catalyst is not permanently damaged since the nitrogen can be burned off the catalyst during regeneration, as a result of which the acidity of the active sites is restored. The carbo-metallic oils may also include significant quantities of pentane insolubles, for example at least about 0.5% by weight, and more typically 2% or more or even about 48 or more. These may include for instance asphaltenes and other materials. Alkali and alkaline earth metals generally do not tend to vaporize in large quantities under the distillation conditions employed in distilling crude oil to prepare the vacuum gas oils normally used as FCC feedstocks. Rather, these metals remain for the most part in the "bottoms" fraction (the non-vaporized high boiling portion) which may for instance be used in the production of asphalt or other by-products. However, reduced crude and other carbometallic oils are in many cases bottoms products, and therefore may contain significant quantities of alkali and alkaline earth metals such as sodium These metals deposit upon the catalyst during cracking. Depending on the composition of the catalyst and magnitude of the regeneration temperatures to which it is exposed, these metals may undergo interactions and reactions with the catalyst (including the catalyst support) which are not normally experienced in processing VGO under conventional FCC processing conditions. If the catalyst characteristics and regeneration conditions so require, one will of course take the necessary precautions to limit the amounts of alkali and alkaline earth metal in the feed, which metals may enter the feed not only as brine associated with the crude oil in ts natural state, but also as components of water or steam which are supplied to the cracking unit. Thus, careful desalting of the crude used to prepare the carbo-metallic feed may be important when the catalyst is particularly susceptible to alkali and alkaline earth metals. In such circumstances, the content of such metals (hereinafter collectively referred to as "sodium") in the feed can be maintained at about 1 ppm or less, based on the weight of the feedstock. Alternatively, the sodium level of the feed may be keyed to that of the catalyst, so as to maintain the sodium level of the catalyst which is in use substantially the same as or less than that of the replacement catalyst which is charged to the unit. According to a particularly preferred embodiment of the invention, the carbo-metallic oil feedstock constitutes at least about 70% by volume of material which boils above about 650"F, and at least about 10 of the material which boils above about 6500F will not boil below about l0250F. The average composition of this 6500F+ material may be further characterized by: (a) an atomic hydrogen to carbon ratio in the range of about 1.3 to about 1.8; (b) a Conradson carbon value of at least about 2; (c) at least about four parts per million of Nickel Equivalents, as defined above, of which at least about two parts per million is nickel (as metal, by weight), at least about 0.1 part per million vanadium; and (d) at least one of the following: (i) at least about 0.3% by weight of sulfur, (ii), at least about 0.05% by weight of nitrogen, and (iii) at least about 0.5% by weight of pentane insolubles. Very commonly, the preferred feed will include all of (i), (ii) and tii4), the other components found in oils of petroleum and non-petroleum origin may also be present in varying quantities providing they do not prevent operation of the process. Although there is no intention of excluding the possibility of using a feedstock which has previously been subjected to some cracking, the present invention has the definite advantage that it can successfully product large conversions and very substantial yields of liquid hydrocarbon fuels from carbo-metallic oils which have not been subjected to any substantial amount of cracking. Thus, for example, and preferably, at least about 85%, more preferably at least about 90 and most preferably substantially all of of the carbo-metallic feed introduced into the present process is oil which has not previously been contacted with cracking catalyst under cracking conditions. Moreover, the process of the invention is suitable for operation in a substantially once-through or single pass mode. Thus, the volume of recycle, if any, based on the volume of fresh feed is preferably about 158 or less and more preferably about 10% or less. In general, the weight ratio of catalyst to fresh feed (feed which has not previously been exposed to cracking catalyst under carving conditions) used in the process is in the range of about 3 to about 19. Preferred and more preferred ratios are about 4 to about 12, more preferably about 5 to about 10 and still more preferably about 6 to about 10, a ratio of about 10 presently being considered most nearly optimum. Within the limitations of product quality reauirements, controlling the catalyst to oil ratio at relatively low levels within the aforesaid ranges tends to reduce the coke yield of the process, based on fresh feed. In conventional FCC processing of VGO, the ratio between the number of barrels per day of plant through-put and the total number of tons of catalyst undergoing circula- tion throughout all phases of the process can vary widely. For purposes of this disclosure, daily plant through-put is defined as the number of barrels of fresh feed boiling above about 6500F which that plant processes per average day of operation to liquid products boiling below about 4300F. For example, in one commercially successful type of FCC-VGO operation, about 8 to about 12 tons of catalyst are under circulation in the process per 1000 barrels per day of plant through-put. In another commercially successful process, this ratio is in the range of about 2 to 3. While the present invention may be practiced in the range of about 2 to about 30 and more typically about 2 to about 12 tons of catalyst inventory per 1000 barrels of daily plant through-put, it is preferred to carry out the process of the present invention with a very small ratio of catalyst weight to daily plant through-put. More specifically, it is preferred to carry out the process of the present invention with an inventory of catalyst that is sufficient to contact the feed for the desired residence time in the above indicated catalyst to oil ratio while minimi=ing the amount of catalyst inventory, relative to plant through-put, which is undergoing circulation or being held for treatment in other phases of the process such as, for example, stripping, regeneration and the like. Thus, more particularly, it is preferred to carry out the process of the present invention with about 2 to about 5 and more preferably about 2 tons of catalyst inventory or less per thousand barrels of daily plant through-put. In the practice of the invention, catalyst may be added 'continuously or periodically, such as, for example, to make up for normal losses of catalyst from the system. Moreover, catalyst addition may be conducted in conunc- tion with withdrawal of catalyst, such as, for e: ?le, to maintain or increase the average activity level of the catalyst in the unit. For example, the rate at which virgin catalyst is added to the unit may be in the range of about 0.1 to about 3, more preferably about 0.15 to about 2, and most preferably to about 0.2 to about 1.5 pounds per barrel of feed. If on the other hand equilibrium catalyst from FCC operation is to be utilized, replacement rates as high as about 5 pounds per barrel can be practiced. Where circumstances are such that the catalyst employed in the Unit is below average in resistance to deactivation and/or conditions prevailing in the unit are such as to promote more rapid deactivation, one may employ rates of addition greater than those stated above; but in the opposite cir cumstances, lower rates of addition may be employed. By way of illustration, if a unit were operated with a met al(s) loading of 5000 ppm Ni + V in parts by weight on equilibrium catalyst, one might for example emplov a replacement rate of about 2.7 pounds of catalyst introduced for each barrel (42 gallons) of feed processed. However, operation at a higher level such as 10,000 ppm Ni + V on catalyst would enable one to substantially reduce the replacement rate, such as for example to about 1.3 pounds of catalyst per barrel of feed. Thus, the levels of metal(s) on catalyst and catalyst replacement rates may in general be respectively increased and decreased to any value consistent with the catalvst activity which is available and desired for conducting the process. Without wishing to be bound by any theory, it appears that a number of features of the process to be described in greater detail below, such as, for instance, the residence time and optional mixing of steam with the feedstock, tend to restrict the extent to which cracking conditions produce metals in the reduced state on the catalyst from heavy metal sulfide(s), sulfate(s) or oxide(s) deposited on the catalyst particles by prior exposures to carbo-metallic feedstocks and regeneration conditions. Thus, the process appears to afford significant control over the poisoning effect of heavy metals on the catalyst even when the accu mulations of such metals are quite substantial. Accordingly, the process may be practised with catalyst bearing high accumulations of heavy metal(s) in the form of elemental metal(s), oxide(s), sulfide(s) or other compounds. Thus, operation of the process with catalyst bearing heavy metals accumulations in the range of about 3800 or more ppm Nickel Equivalents, on the average, is contemplated. The concentration of Nickel Equivalents of metals on catalyst can range up to about 50,000 ppm or higher. More specifically, the accumulation may be in the range of about 300d to about 30,000 ppm, preferably in the range of about 3000 to 20,000 ppm, and more particularly about 3000 to about 12,000 ppm. Within these ranges just mentioned, operation at metals levels of about 4000 or more, about 5000 or more, or about 7000 or more ppm can tend to reduce the rate of catalyst replacement required. The foregoing ranges are based on parts per million of Nickel Equivalents, in which the metals are expressed as metal, by weight, measured on and based on regenerated ecuilibrium catalyst. However, in the event that catalyst of adequate activity is available at very low cost, making feasible very high rates of catalyst replacement, the carbo-metallc oil could be converted to lower boiling liquid products with catalyst bearing less than 3,000 ppm Nickel Eeuiva- lents of heavy metals. For example, one might employ equilibrium catalyst from another unit, for example, an FCC unit which has been used in the cracking of a feed, e.g. vacuum gas oil, having a carbon residue on pyrolysis of less than 1 and containing less than about 4 ppm Nickel Equivalents of heavy metals. In any event, the equilibrium concentration of heavy metals in the circulating inventory of catalyst can be controlled (including maintained or varied as desired c needed) by manipulation of the rate of catalyst addition discussed above. Thus, for example, addition of catalyst m.v be maintained at a rate which will control the heavy metals accumulation on the catalyst in one of the ranges set forth above. In general, it is preferred to employ a catalyst having a relatively high level of cracking activity, providing high levels of conversion and productivity at low residence tines. The conversion capabilities of the catalyst may be expressed in terms of the conversion produced during actual operation of the process and/or in terms of conversion produced in standard catalyst activity tests. For example, it is preferred to employ catalyst which, in the course of extended operation under provailing process conditions, is sufficiently active for sustaining a level of conversion of at least about 504 and more preferably at least about 60t,. In this connection, conversion is expressed in liquid volume percent, based on fresh feed. Also, for example, the preferred catalyst may be defined as one which, in its virgin or equilibriun state, exhibits a specified activity expressed as a percentage in terms of MAT (micro-activity test) conversion. For purposes of the present invention the foregoing percentage is the volume percentage of standard feedstock which a catalyst under evaluation will convert to 4300F end point gasoline, light er products and coke at 9000F, 16 5E.SV (weight hourly space velocity, calculated on a moisture free basis, using clean catalyst which has been dried at 11000F, weighed and then conditioned, for a period of at least 8 hours at about 25 C and 50% relative humidity, until about one hour or less prior to contacting the feed) and 3C/O (catalyst to oil weight ratio) by ASTM D-32 MAT test D-3907-a0, using an appropriate standard feedstock, e.g. a sweet light primary gas oil, such as that used by Davison, Division of W.R. Grace, having the following analysis and properties: API Gravity at 600F, degrees 31.0 Specific Gravity at 600F, g,'cc 0.8708 Ramsbottom Carbon, wt. t 0.09 Conradson Carbon, wt. %(east.) 0.04 Carbon, wt. % 84.92 Hydrogen, wt. % 12.94 Sulfur, wt. % 0.68 Nitrogen, ppm 305 Viscosity at 100 0F, centistokes 10.36 Watson K Factor 11.93 Aniline Point 182 Bromine No. 2.2 Paraffins, Vol. % 31.7 Olefins, Vol. % 1.6 Naphthenes, Vol. z 44.0 Aromatics, Vol. % 22.7 Average molecular Weight 284 Nickel Trace Vanadium Trace Iron Trace Sodium Trace Chlorides Trace H S & W Trace Distillation ASTM D-1160 IBP 445 10% 601 30% 664 50% 701 70% 734 90% 787 FBP 834 The gasoline end point and boiling temperature-volume percent relationships of the product produced in the YT conversion test may for example be determined by simulated distillation techniques, for example modifications of gas chromate graphic "Sim-D", ASTM D-2887-73. The results of such simulations are in reascnable agreement with the results obtained by subjecting larger samples of material to standard laboratory distillation techniques. Conversion is calculated by subtracting from 100 the volume percent (based on fresh feed) of those products heavier than gasoline which remain in the recovered product. On page 935-937 of Hougen and Watson, Chemical Process Princinles, John Wiley & Sons, Inc., N.Y. (1947), the concent of "Activity Factors" is discussed. This concept leads to the use of "relative activity" to compare the effectiveness of an operating catalyst against a standard catalyst. Relative activity measurements facilitate recoc- nition of how the quantity resuirements of various catalysts differ from one another. Thus, relative activity is a ratio obtained by dividing the weight of a standard or reference catalyst which is or would be required to produce a given level of conversion, as compared to the weight of an operating catalyst (whether proposed or actually used) which is or would be recuired to produce the same level of conversion in the same or equivalent feedstock under the same or equivalent conditions. Said ratio of catalyst weights may be expressed as a numerical ratio, but preferably is converted to a percentage basis. The standard catalyst is preferably chosen from among catalysts useful for conducting the present invention, such as for example zeolite fluid cracking catalysts, and is chosen for its ability to produce a predetermined level of conversion in a standard feed under the conditions of temperature, aSSV, catalyst to oil ratio and other conditions set forth in the preceding description of the MAT conversion test and in ASTM D-32 MAT test D-j907-60. Conversion is the voltme percentage of feedstock that is converted to 4300F endpoint gasoline, lighter products and coke. For standard feed, one may employ the above-mentioned light primary gas oil, or equivalent. For purposes of conducting relative activity dcteriations, one may prepare a "standard catalyst curve", a chart or graph of conversion (as above defined) vs. reciprocal WHSV for the standard catalyst and feedstock. A sufficient num ber of runs i3 made under ASTRO D-3907-80 conditions (as modified above) using standard feedstock at varying levels of X*tSV to prepare an accurate curve of conversion vs. WHSV for the standard feedstock. This curve should traverse all or substantially all of the various levels of conversion including the range of conversion within which it is expected that the operating catalyst will be tested. From this curve, one may establish a standard IfGSV for test comparisons and a standard value of reciprocal WNISV correspoinding to that level of conversion which has been chosen to represent 1005 relative activity in the standard catalyst. For purposes of the present disclosure the aformentioned reciprocal tHSV and level of conversion are, respectively, 0.0625 and 75t. In testing an operating catalyst of unknown relative activity, one conducts a sufficient number of runs with that catalyst under D-3907-20 conditions (as modified above) to establish the level of conversion which is or would be produced with the operating catalyst at standard reciprocal 'SV. Then, using the above-mentioned standard catalyst curve, one establishes a hypothetical reciprocal t*tSV constituting the reciprocal WiSV which would have been required1 using the standard catalyst, to obtain the same level of conversion which was or would be exhibited, by the operating catalyst at standard WHSV. he relative activity may then be calculated by dividing the hypothetical reciprocal tfliSV by the reciprocal standard WHSV, which is 1/16, or .0625. The result is relative activity expressed in terms of a decimal fraction, which may then be multiplied by 100 to convert to percent relative activity. In anplying the results of this determination, a relative activity of 0.5, or 50%, means that it would take twice the amount of the operating catalyst to give the same conversion as the standard catalyst, i.e., the production catalyst is 50% as active as the reference catalyst. The catalyst may be introduced into the process-in its virgin form or, as previously indicated, in other than virgin form; e.g. one may use equilibrium catalyst withdrawn from another unit, such as catalyst that has been employed in the cracking of a different feed. Whether characterized on the basis of MAT conversion activity or relative activity, the preferred catalysts may be described on the basis of their activity "as introduced" into the process of the present invention, or on the basis of their "as withdrawn" or equilibrium activity in the process of the present invention, or on both of these bases. A preferred activity' level of virgin and non-virgin catalyst "as introduced" into the process of the prescnt invention is at least about 60% by MAT conversion, and preferably at least about 20t, more preferably at least about 40% and still more preferably at least about 605 in terms of relative activity. However, it will be appreciated that, particularly in the case of non-virgin catalysts supplied at high addition rates, lower activity levels may be acceptable. An acceptable "as withdrawn" or equilibrium activity level of catalyst which has been used in the process of the present invention is at least about 20% or more, but about 40% or more and preferably about 60% or more are preferred values on a relative activity basis, and an activity level of 60% or more on a MAT conversion basis is also contemplated. More preferably, it is desired to erploy a catalyst which will, under the conditions of use in the ur.it, establish an equilibrium activity at or above the indicated level. The catalyst activities are determined with catalyst having less than 0.01 coke, e.g. regenerated catalyst. One may employ any hydrocarbon cracking catalyst having the above indicated conversion capabilities. A particular- ly preferred class of catalysts includes those which have pore structures into which molecules of feed .mat2rial =ay enter for adsorption and/or for contact with active cataly- tic sites within or adjacent the pores. Various types of catalysts are available within this classification,- including for example the layered silicates, e.g. smectites. Although the most widely available catalysts within this classification are the well-known zeolite-containing catalysts, non-zeolite catalysts are also contemplated. The preferred zeolite-containing catalysts may include any zeolite, whether natural, semi-synthetic or synthetic, alone or in admixture with other materials which do not significantly impair the suitability of the catalyst, provided the resultant catalyst has the activity and pore structure referred to above. For example, if the virgin catalyst is a mixture, it may include the zeolite common nent associated with or dispersed in a porous refractory inorganic oxide carrier, in such case the catalyst may for example contain about 1% to about 60%, more Preferably about 15 to about 50%, and most typically about 20 to about 45% by weight, based on the total weight of catalyst (water free basis) of the zeolite, the balance of the catalyst being the porous refractory inorganic oxide alone or in combination with any of the known adjuvants for proofing or suppressing various desired and undesired reactions. For a general explanation of the genus of zeolite, mole- cular sieve catalysts useful in the invention, attention is drawn to the disclosures of the articles entitled "Refinery Catalysts Are a Fluid Business" and "Making Cat Crackers Work On Varied Diet", appearing respectively in the July 26, 1978 and September 13, 1978 issues of Chemical Week magazine. The descriptions of the aforementioned publications are incorporated herein by reference. For the most part, the zeolite components of the zeolitecontaining catalysts will be those which are known to be useful in FCC cracking processes. In general, these are crystalline aluminosilicates, typically made up of tetra coordinated aluminn atoms associated th¯¯ugh oxygen atoms with adjacent silicon atoms in the cr'stai structure. However, the term zeoliteS as used in this disclosure contemplates not only aluminosilicates, but also substances in which the aluminum has been partly or wholly replaced, such as for instance by gallium and/or other metal atoms, and further includes substances in which all or part of the silicon has been replaced, such as for instance by germanium. Titanium and zirconium substitution may also be practiced. Most zeolites are prepared or occur naturally in the sodium form, so that sodium cations are associated with the electro- negative sites in the crystal structure. The sodium cations tend to make zeolites inactive and much less stable when exposed to hydrocarbon conversion conditions, particularly high temperatures. Accordingly, the zeolite may be ion exchanged, and where the zeolite is a component of a catalyst composition, such ion exchanging may occur before or after incorporation of the zeolite as a component of the composition. Suitable cations for replacement of sodium in the zeolite crystal structure include arzlonium (decomposable to hydrogen), hydrogen, rare earth metals, alkaline earth metals, etc. Various suitable ion exchange procedures and cations which may be exchanged into the zeolite crystal structure are well known to those skilled in the art. Examples of the naturally occurring crystalline aluminosilicate zeolites which may be used as or included in the catalyst for the present invention are faujasite, mordenite, clinoptilote, chabazite, analcite, crionite, as well as levynite, dachiardite, paulingite, noselite, ferriorite, heulandite, scolccite, stibite, harmotome, phillipsite, brewsterite, flarite, datolite, gmelinite, caunnite, leucite, lazurite, scaplite, mesolite, ptolite, nephline, matrolite, offretite and sodalite. Examples of the synthetic crystalline alum-nosili¯te zeolites which are useful as or in the catalyst for carrying out the present invention are Zeolite X, U. S. Patent NO= 2;8S2,244, Zeolite Y, U. S Patent No. 3,130,007; and Zeolite A, U. S. Patent No. 2,882,243; as well as Zeolite B, U. S. Patent No. 3,008,803; Zeolite D, Canada Patent No. 661,981; Zeolite E, Canada Patent No. 614,495; Zeolite F, U. S. Patent No. 2,996,358; Zeolite H. U. S. Patent No. 3,010,789; Zeolite J., U. S. Patent No. 3,011,869; Zeolite L, Belgian Patent No. 575,177; Zeolite M., U. S. Patent No. 2,995,423, Zeolite O, U. S. Patent No. 3,140,252; Zeolite Q, U. S. Patent No. 2,991,151; Zeolite S, U. S. Patent No. 3,054,657, Zeolite T, U. S. Patent No. 2,950,952; Zeolite W, U. S. Patent No. 3,012,853; Zeolite Z, Canada Patent No. 614,495; and Zeolite Omega, Canada Patent No. 817,915. Also, ZK-4HJ, alpha beta and ZSM-type zeolites are useful. Moreover, the zeolites described in U. S. Patents Nos. 3,140,249, 3,140,253, 3,944,482 and 4,137,151 are also useful, the disclosures of said patents being incorporated herein by reference. The crystalline aluminosilicate zeolites having a faujasite-type crystal structure are particularly preferred for use in the present invention. This includes particularly natural faujasite and Zeolite X and Zeolite Y. The crystalline al minosil cate zeolites, such as synthetic faujasite, will under normal conditions crystallize as regularly shaped, discrete particles of about one to about ten microns in size, and, accordingly, this is the size range frequently found in commercial catalysts which can be used in the invention. Preferably, the particle size of the zeolites is from about 0.1 to about 10 microns and more preferably is from about 0.1 to about 2 microns or less. For eagle, zeolites prepared in s.'tu from calcined kaolin may be characterized by even smaller crystallites. Crstal- line zeolites exhibit both an inter-or and an exterior surface area, which we have defined as "portal" surface area, with the largest portion of the total surface area being internal. By portal surface area, we refer to the outer surface of the zeolite crystal through which reactants are considered to pass in order to convert to lower boiling products: Blockages of the internal channels by, for example, coke formation, blockages of entrance to the internal channels by deposition of coke in the portal surface area, and contamination by metals poisoning, will greatly reduce the total zeolite surface area. Therefore, to minimize the effect of contamination and pore blockage, crystals larger than the normal size cited above are preferably not used in the catalysts of this invention. Commercial zeolite-containing catalysts are available with carriers containing a variety of metal oxides and combination thereof, including for example silica, alumina, magnesia, and mixtures thereof and mixtures of such oxides with clays as e.g. described in U. S. Patent No. 3,034,948. One may for example select any of the zeolite-containing molecular sieve fluid cracking catalysts which are suitable for production of gasoline from vacuum gas oils. However, certain advantages may be attained by judicious selection of catalysts having marked resistance to metals. A metal resistant zeolite catalyst is, for instance, described in U. S. Patent No. 3,944,482, in which the catalyst contains 1-40 weight percent of a rare earth-exchanged zeolite, the balance being a refractory metal oxide having specified pore volume and size distribution. Other catalysts described as "metals-tolerant" are described in the above mentioned Cimbalo et al article. In general, it is preferred to employ catalysts having an over-all particle size in the range of about 5 to about 160, more preferably about 40 to about 120, and most preferably about 40 to about 80 icrons. For example, a useful catalyst may have a skeletal density of about 150 pounds per cubic foot and an average particle size of about 60-70 microns, with less than 10% of the particles having a size less than about 40 microns and loss than 80% having a size less than about 50-60 microns. Although a wide variety of other catalysts, including both zeolite-containing and non-zeolite-containing may be employed in the practice of the invention the following are exar,ples of commercially available catlysts which may be employed in practicing the invention: : TABLE 2 Specific weight Percent Surface Zeolite m2 /g Content A12O3 SiO2 Na2O Fe2O TiO2 AGZ-290 300 11.0 29.5 59.0 0.40 0.11 0.59 GRZ-1 162 14.0 23.4 69.0 0.10 0.4 0.9 CCZ-220 129 11.0 34.6 60.0 0.60 0.57 1.9 Super DX 155 13.0 31.0 65.0 0.80 0.57 1.6 F-87 240 10.0 44.0 50.0 0.80 0.70 1.6 FOX-90 240 8.0 44.0 52.0 0.65 0.65 1.1 RFZ 20 310 20.0 59.0 40.0 0.47 0.54 2.75 HEZ 55 210 19.0 59.0 35.2 0.60 0.60 2.5 The AGZ-290, GRZ-1, CCZ-220 and Super DX catalysts refered to above are products of W. R. Grace and Co. F-â7 and FOC-90 are products of Filtrol, while HFZ-20 and HEZ-55 are products of Engelhard/Houdry. The above are prcperTes of virgin catalyst and, except in the case of zeolite con- tent, are adjusted to a water free basis, i.e. based on material ignited at 1750 F. The zeolite content is derived by comparison of the X-ray intensities of a catalyst sam it and of a standard material composed of high purity sodium Y zeolite in accordance with draft =6, dated January 9, 1978, of proposed ASTM Standard Method entitled "Determina- tion of the Faujasite Content of a Catalyst.' Among the above mentioned commercially available catalysts, the Super D family and especially a catalyst designated GRZ-1 are particularly preferred. For example, Super DX has given particularly good results with Arabian light crude. The GRZ-l, although substantially more expensive than the Super DX at present, appears somewhat more metals tolerant. Although not yet commercially available, it is believed that the best catalysts for carrying out the present invention will be those which, according to proposals advanced by Dr. William P. Hettinger, Jr. and Dr. James E. Lewis, are characterized by matrices with feeder pores having large .-.;n¯.-um diameters and large mouths to facilitate diffusion of high molecular weight molecules through the matrix to the portal surface area of molecular sieve particles within the matrix. Such matrices preferably also have a relatively large pore volume in order to soak up unvaporized portions of the carbo-metallic oil feed. Thus significant nunbers of liquid hydrocarbon molecules can diffuse to active catalytic sites both in the matrix and in sieve particles on the surface of the matrix. In general it is preferred to employ catalysts with matrices wherein the feeder pores have diameters in the range of about 400-to about 6000 angstrom units, and preferably about 1000 to about 6000 angstrom units. It is considered an advantage that the process of the present invention can be conducted in the substantial absence of tin and/or antimony or at least in the presence of a catalyst which is substantially free of either or both of these metals. The process of the present invention may Se operated with the above described carbo-metailic oil and catalyst as sub stantially the sole materials charged te the reaction one But the charging of ad-¯t;onai materials is not e:cluded. The charging of recycled oil to the reaction zone has already been mentioned. As described in greater detail below, still other materials fulfilling a variety of functions may also be charged.. In such case, the carbo-metallic oil and catalyst usually represent the major proportion by weight of the total of all materials charged to the reaction zone. Certain of the additional materials which may be used perform functions which offer significant advantages over the process as performed with only the carbo-metallic oil and catalyst. Amona these functions are: controlling the effects of heavy metals and other catalyst contaminants; enhancing catalyst activity; absorbing excess heat in the catalyst as received from the regenerator; disposal of pollutants or conversion thereof to a form or forms in which they may be more readily separated from products and/ or disposed of; controlling catalyst temperature, diluting the carbo-metallic oil vapors to reduce their partial pressure and increase the yield of desired products; adjusting feed/catalyst contact time; donation of hydrogen to a hydrogen deficient carbo-metallic oil feedstock, for einle as disclosed in a United States application entitled "';se of :;aphtha in Carbo-Metallic Oil Conversion" and filed in the name of George D. Myers on larch 23, 1981, assisting in the dispersion of the feed; and possibly a so distillation of products. Certain of the petals in the heavy metals accumulation on the catalyst are mere active in promoting undesired reactions when they are in the form of elemental metal, than they are when in the oxidized form produced by contact with oxygen in the catalyst regenera- tor. However, the time of contact between catalyst and -:a- pors of feed and product in past convencional catalytic cracking was sufficient so that hydrogen released in the cracking reaction was able to reconvert a significant por- tion of the less harmful asides bacl: to the ore har:'.f"l elemental heavy metals. One can take advantage cf this situation through the introd'ct'cn of additional nateras which are in gaseous (including vaporous) form in the reaction zone in admixture with the catalyst and vapors of feed and products. The increased volume of material in the reaction zone resulting from the presence of such additional materials tends to increase the velocity of flow through the reaction zone with a corresponding decrease in the residence time of the catalyst and oxidized heavy metals borne thereby. Because of this reduced residence time, there is less opportunity for reduction of the oxidized heavy metals to elemental form and therefore less of the harmful elemental metals are available for contacting the feed and products. Added materials may be introduced into the process in any suitable fashion, some examples of which follow. For instance, they may be admixed with the carbo-metallic oil feedstock prior to contact of the latter with the catalyst. Alternatively, the added materials may, if desired, be admixed with the catalyst prior to contact of the latter with the feedstock. Separate portions of the added materials may be separately admixed with both catalyst and carbo-metallic oil. Moreover, the feedstock, catalyst and additional oatetlals may, if desired, be brought together substantially simultaneously. A portion of the added materials may be mixed with catalyst and/or carbo-metallic oil in any of the above described ways, while additional portions are subsequently brought into admixture. For example, a portion of the added materials may be added to the carbo-metallic oil and/or to the catalyst before they reach the reaction zone, while another portion of the added materials is introduced directly into the reaction zone. The added materials may be introduced at a plurality of spaced locations in the reaction "one or along the length thereof, if elongated. The amount of additional materials which may be present in the feed, catalyst or reaction zone for c-rr:lng out the above functions, and others, may be varied as desired; but said amount will preferably be sufficient to substantially heat balance the process. These materials may for example be introduced into the reaction zone in a weight ratio relative to feed of up to about 0.4, preferably in the range of about 0.02 to about 0.4, more preferably about 0.03 to about 0.3 and most preferably about 0.05 to about 0.25. For example, many or all of the above desirable functions may be attained by introducing H20 to the reaction zone in the form of steam or of liquid water or a combination thereof in a weight ratio relative to feed in the range of about 0.04 or more, or more preferably about 0.05 to about 0.1 or more. Without wishing to be bound by any theory, it appears that the use of H20 tends to inhibit reduction of catalyst-borne oxides, sulfites and sulfides to the free metallic form which is believed to promote condensation-dehydrogenation with consequent promotion of coke and hydrogen yield and accompanying loss of product. Moreover, H20 may also, to some extent, reduce deposition of metals onto the catalyst surface. There may also be some tendency to desorb nitrogen-containing and other heavy contaminant-containing molecules from the surface of the catalyst particles, or at least some tendency to inhibit their absorption by the catalyst. It is also believed that added H20 tends to increase the acidity of the catalyst by Bronsted acid formation which in turn enhances the activity of the catalyst. Assuming the S2 as supplied is cooler than the regenerated catalyst and/or the temperature of the reaction zone, the sensible heat involved in raising the temperature of the H20 upon contacting the catalyst in the reaction zone or elsewhere can absorb excess heat from the catalyst. Where the H2 0 is or includes recycled water that contains for example about 500 to about 5000 pm of H25 dissolved therein, a number of additional advantages may accrue. The ecologically unattractive H2 S need not be vented to the atmosphere, the recycled water does not require further treatment to remove H2S and the H25 may be of assistance in reducing coking of the catalyst by passivation of the heavy metals, i.e. by conversion thereof to the sulfide form which has a lesser tendency than the free metals to enhancc coke and hydrogen production. In the reaction zone, the presence of H2O can dilute the carbo-metallic oil vapors, thus reducing their partial pressure and tending to increase the yield of the desired products. It has been reported that H20 is useful in combination with other materials in generating hydrogen during cracking; thus it may be able to act as a hydrogen donor for hydrogen deficient carbo-metallic oil feedstocks. The H20 may also serve certain purely mechanical functions such as: assisting in the atomizing or dispersion of the feed; competing with high molecular weight molecules for adsorption on the surface of the catalyst, thus inter=upt- ing coke formation; steam distillation of vaporizable product from unvaporized feed material; and disengagement of product from catalyst upon conclusion of the cracking reaction. It is particularly preferred to bring together H20, catalyst and carbo-metallic oil substantially simul- taneously. For example, one may admix H20 and feedstock in an atomizing nozzle and ismediately direct the resultant spray into contact with the catalyst at the owmstrean end of the reaction zone. The addition of steam to the reaction zone is frequently mentioned in the literature of fluid catalytic cracking. Addition of liquid water to the feed is discussed relatively infrequently, compared to the Introduction of steam directly into the reaction zone. However, in accordance with the present invention it is particularly preferred that liquid water be brought into intimate admixture with the carbo-metallic oil in a weight ratio of about 0.04 to about 0.25 at or prior to the time of itrbducrion of the oil into the reaction zone, whereby the water (e.g., in the form of liquid water or in the form of steam produced by vaporization of liquid water in contact with the oil) enters the reaction zone as partof the flow of fecdstock which enters such zone. Although not wishing to be bound by any theory, it is believed that the foregoing is advantageous in promoting dispersion of the feedstock. Also, the heat of vaporization of the water, which heat is absorbed from the catalyst, from the feedstock, or from both, causes the water to be a more efficient heat sink than steam alone. Preferably the weight ratio of liquid water to feed is about 0.04 to about 0.2 more preferably about 0.05 to about 0.15. Of course, the liquid water may be introduced into the process in the above described manner or in other ways, and in either event the introduction of liquid water may be accompanied by the introduction of additional amounts of water as steam into the same or different portions of the reaction zone or into the catalyst and/or feedstock. For example, the amount of additional steam may be in a weight ratio relative to feed in the range of about 0.01 to about 0.25, with the weight ratio of total H20 (as steam and liquid water) to feedstock being about 0.3 or less. The charging weight ratio of licuid water relative to steam in such combined use of liquid water and steam may for example range from about 15 which is presently preferred, to about 0.2. Such ratio may be ".aintained at a predeter-..ined level within such range or varied as necessary or desired to adjust or maintain heat balance. Other materials may be added to the reaction zone to perform one or more of the above describe functions. For example, the dehydrogenation-condensation activity of heaxa metals may be inhibited by introducing hydrogen sulfide gas into the reaction zone. Hydrogen may be made available for hydrogen deficient carbo-metallic oil feedstocks by introducing into the reaction zone either a convent onal hydrogen donor diluent such as a heavy naphta or relative ly low molecular weight carbon-hydrogen fragment contributors, including for example: light paraffins; low molecular weight alcohols and other compounds which permit or favor intermolecular hydrogen transfer, and compounds that chemically combine to generate hydrogen in the reaction zone such as by reaction of carbon monoxide with water, or with alcohols, or with olefins, or with other materials or mixtures of the foregoing. All of the above mentioned additional materials (including water), alone or in conjunction with each other or in conjunction with other materials, such as nitrogen or other inert gases, light hydrocarbons, and others, may perform any of the above-described functions for which they are suitable, including without limitation, acting as diluents to reduce feed partial pressure and/or as heat sinks to absorb excess heat present in the catalyst as received from the regeneration step. The foregoing is a discussion of some of the functions which can be performed by materials other than catalyst and carbo-metallic oil feedstock intro duced into the reaction zone, and it should be understood that other materials may be added or other functions per formed without departing from the spirit of the invention. The invention may be practiced in a wide variety of appara tus. However, the preferred apparatus includes means for rapidly vaporizing as much feed as possible and efficiently admixing feed and catalyst (although not necessarily in that order), for causing the resultant mixture to flow as a dilute suspension in a progressive flow mode, and for separating the catalyst from cracked products and any cracked or only partially cracked feed at the end of a predetermined residence time or times, it being preferred that all or at least a substantial portion of the product should be abruptly separated from at least a portion of the catalyst. For example, the apparatus may include, along its elongated reaction chamber, one or more points for introduction of carbo-metallic feed, one or more points for introduction of catalyst, one or more points for introduction of additional material, one or more points for withdrawal of products and one or. more points for withdrawal of catalyst. The means for introducing feed, catalyst and other material may range from open pipes to sophisticated jets or spray nozzles, it being preferred to use means capable of breaking up the liquid feed into fine droplets. Preferably, the catalyst, liquid water (when used) and fresh feed are brought together in an apparatus similar to that disclosed in U. S. Patent Application Serial No. 969,601 of George D. Myers et al, filed December 14, 1978, the entire disclosure of which is hereby incorporated herein by refarence. According to a particularly preferred embodiment based on a suggestion which is understood to have emanated from Mr. Steven M. Kovach, the liquid water and carbo-metallic oil, prior to their introduction into the riser, are caused to pass through a propeller, apertured disc, or any appropriate high shear agitating means for forming a "homogenized mixture" containing finely divided droplets of oil and/or water with oil and/or water present as a continuous phase. It is preferred that the reaction chamber, or at least the major portion thereof, be more nearly vertical than horizontal and have a length to diameter ratio of at least about 10, more preferably about 20 or 25 or more. Use of a vertical riser type reactor is preferred. If tubular, the reactor can be of uniform diameter throughout or may be provided with a continuous or step-wise increase in diameter along the reaction path to maintain or vary the velocity along the flow path. In general, the charging means (for catalyst and feed) and the reactor configuration are such as to provide a relatively high velocity of flow and dilute suspension of catalyst. For example, the vapor or catalyst velocity in the riser will be usually at least about 25 and more typically at least about 35 feet per second. This velocity may range up to about 55 or about 75 feet or about 100 feet per second or higher. The vapor velocity at the top of the reactor may be higher than that at the bottom and may for example be about 80 feet per second at the top and about 40 feet per second at the bottom. The velocity capa- bilities of the reactor will in general be sufficient to prevent substantial build-up of catalyst bed in the bottom or other portions of the riser, whereby the catalyst loading in the riser can be maintained below about 4 or 5 pounds, as for example about 0.5 pounds, and below about 2 pounds, as for example 0.8 pound, per cubic foot, respectively, at the upstream (e.g. bottom) and downstream (e.g. top) ends of the riser. The progressive flow mode involves, for example, flowing of catalyst, feed and products as a stream in a positively controlled and maintained direction established by the elongated nature of the reaction zone. This is not to suggest however that there must be strictly linear flow. As is well known, turbulent flow and "slippage" of catalyst may occur to some extent especially in certain ranges of vapor velocity and some catalyst loadings, although it has been reported adviseable to employ sufficiently low catalyst loadings to restrict slippage and back-mixing. Most preferably the reactor is one which abruptly separates a substantial portion or all of the vaporized cracked products from the catalyst at one or more points along the riser, and preferably separates substantially all of the vaporized cracked products from the catalyst at the dox stream end of the riser. A preferred type of reactor embodies ballistic separation of catalyst and products; that is, catalyst is projected in a direction established by the riser tube, and is caused to continue its motion in the general direction so established, while the products, having lesser momentum, are caused to make an abrupt change of direction, resulting in an abrupt, substantially instantaneous separation of product from catalyst. In a preferred embodiment referred to as a vented riser, the riser tube is provided with a substantially unobstructed discharge opening at its downstream end for discharge of catalyst. An exit port in the side of the tube adjacent the downstream end receives the products. The discharge opening communicates with a catalyst flow path which extends to the usual stripper and regenerator, while the exit port communicates with a product flow path which is substantially or entirely separated from the catalyst flow path and leads to separation means for separating the products from the relatively small portion of catalyst, if any, which manages to gain entry to the product exit port. Examples of a ballistic separation apparatus and technique as above described, are found in U. S. Patents 4,066,533 and 4,070,159 to Myers et al, the disclosures of which patents are hereby incorporated herein by reference in their entireties. According to a particularly preferred embodiment, based on a suggestion understood to have emanated from Paul W. Walters, Roger M. Benslay and Dwight F. Barger, the ballistic separation step includes at least a partial reversal of direction by the product vapors upon discharge from the riser tube; that is, the product vapors make a turn or change of direction which exceeds 900 at the riser tube outlet. This may be accomplished for example by providing a cup-like member surrounding the riser tube at its upper end, the ratio of cross-sectional area of the cup-like member relative to the cross-sectional area of the riser tube outlet being low i.e. less than 1 and preferably less than about 0.6. Preferably the lip of the cup is slightly ee-.-n- stream of, or above the downstream end or top of the riser tube, and the cup is preferably concentric with the riser tube. By means of a product vapor line communicating with the interior of the cup but not the interior of the riser tube, having its inlet positioned within the cup interior in a direction upstream of the riser tube outlet, product vapors emanating from the riser tube and entering the cup by reversal of direction are transported away from the cup to catalyst and product separation equipment. Such an arrangement can produce a high degree of completion of the separation of catalyst from product vapors at the riser tube outlet, so that the required amount of auxiliary catalyst separation equipment such as cyclones is greatly - reduced, with consequent large savings in capital investment and operating cost. Preferred conditions for operation of the process are described below. Among these are feed,- catalyst and reaction temperatures, reaction and feed pressures, residence time and levels of conversion, coke production and coke laydown on catalyst. In conventional FCC operations with VGO, the feedstock is customarily preheated, often to temperatures significantly higher than are required to make the feed sufficiently fluid for pumping and for introduction into the reactor. For example, preheat temperatures as high as about 7000 or 800F have been reported. But in our process as presently practiced it is preferred to restrict preheating of the feed, so that the feed is capable of absorbing a larger amount of heat from the catalyst while the catalyst raises the feed to conversion temperature, at the same time minimizing utilization of external fuels to heat the feedstock. Thus, where the nature of the feedstock permits, it may be fed at ambient temperature. Heavier stocks may be fed at preheat temperatures of up to about 6000F, typically about 2000F to about ;00 F, but higher preheat temperatures are not necessarily excluded. The catalyst fed to the reactor may vary widely in temperature, for example from about 1100e to about 16000r, more preferably about 1200e to about 1500"F and most preferably about 13000 to about 14000F, with about 13250 to about 13750 being considered optimum at present. As indicated previously, the convcrsion of the carbometallic oil to lower molecular weight products may be conducted at a temperature of about 9000 to about 14000F, measured at the reaction chamber outlet. The reaction temperature as measured at said outlet is more preferably maintained in the range of about 965 to about 1300"F, still more preferably about 9750 to about 12000F, and most preferably about 980 to about 11500F. Depending upon the temperature selected and the properties of the feed, all of the feed may or may not vaporize in the riser. Although the pressure in the reactor may, as indicated above, range from about 10 to about 50 psia, preferred and more preferred pressure ranges are about 15 to abut 35 and about 20 to about 35. In general, the partial (or total) pressure of the feed may be in the range of about 3 to about 30, more preferably about 7 to about 25 and most preferably about 10 to about 17 psia. The feed partial pressure may be controlled or suppressed by the introduction of gaseous (including vaporous) materials into the reactor, such as for instance the steam, water and other additional materials described above. The process has for example been operated with the ratio of feed partial pressure relative to total pressure in the riser in the range of about 0.2 to about 0.8, more typically about 0.3 to about 0.7 and still more typically about 0.4 to about 0.6, with the ratio of added gaseous material (which may includes recycled gases and/or steam resulting from introduction of H20 to the riser in the form of steam and/or liquid water) relative to total pressure in the riser correspondingly ranging from about 0.8 to about 0.2, more typically about 0.7 to about 0.3 and still more typically about 0.6 to about 0.4. In the illustrative operations just described, the ratio of the partial pressure of the added gaseous material relative to the partial pressure of the feed has been in the range of about 0.25 to about 4.0, more typically about 0.4 to about 2.3 and still more typically about 0.7 to about 1.7. Although the residence time of feed and product vapors in the riser may be in the range of about 0.5 to about 10 seconds, as described above, preferred and more preferred values are about 0.5 to about 6 and about 1 to about 4 seconds, with about 1.5 to about 3.0 seconds currently being considered about optimum. For example, the process has been operated with a riser vapor residence time of about 2.5 seconds or less by introduction of copious amounts of gaseous materials into the riser, such amounts being sufficient to provide for example a partial pressure ratio of added gaseous materials relative to hydrocarbon feed of about 0.8 or more. By way of further illustration, the process has been operated with said residence time being about two seconds or less, with the aforesaid ratio being in the range of about 1 to about 2. The combination of low feed partial pressure, very low residence ti,rne and ballistic separation of products from catalyst are considered especially beneficial for the conversion of carbo-metallic oils. Additional benefits may be obtained in the foregoing combination when there is a substantial partial pressure of added gaseous material, especially H20 as described above. Depending upon whether there is slippage between the catalyst and hydrocarbon vapors in the riser, the catalyst riser residence time may or may not be the same as that of te vapors. Thus, the ratio of average catalyst reactor residence time versus vapor reactor residence time, i.e. slippage, may be in the range of about 1 to about 5, more preferably about 1 to about 4 and most preferably about 1 to about 3, with about 1 to about 2 currently being considered optimum. In practice, there will usually be a small amount of slippage, e.g., at least about 1.05 or 1.2. In an operating unit there may for example be a slippage of about 1.1 at the bottom of the riser and about 1.05 at the top. In certain types of known FCC units, there is a riser which discharges catalyst and product vapors together into an enlarged chamber, usually considered to be part of the reactor, in which the catalyst is disengaged from product and collected. Continued contact of catalyst uncracked feed (if any) and cracked products in such enlarged chamber results in an overall catalyst feed contact time appreciably exceeding the riser tube residence times of the vapors and catalysts. When practicing the process of the present invention with ballistic separation of catalyst and vapors at the downstream (e.g. upper) extremity of the riser, such as is taught in the above mentioned Myers et al patents, the riser residence time and the catalyst contact time are substantially the same for a major portion of the feed and product vapors. It is considered advantageous if the vapor riser residence time and vapor catalyst contact time are substantially the same for at least about 80g, more preferably at least about 90% and most preferably at least about 95 by volume of the total feed and product vapors passing through the riser. By denying such vapors continued contact with catalyst in a catalyst disengagement and collection chamber one may avoid a tendency toward re-cracking and diminished selectix ty. In general, the combination of catalyst to oil ratio, temperatures, pressures and residence times should be such as to effect a substantial conversion of the carbometallic oil feedstock. It is an advantage of the process that very high levels of conversion can be attained in a single pass; for example the conversipn may be in excess of 50 and may range to about 90% or higher. Preferably, the aforementioned conditions are maintained at levels sufficient to maintain conversion levels in the range of about 60 to about 90% and more preferably about 70 to about 35z. The foregoing conversion levels are calculated by subtracting from 100% the percentage obtained by dividing the liquid volume of fresh feed into 100 times the volume of liquid product boiling at and above 4300F (tbp, standard atmospheric pressure). These substantial levels of conversion may and usually do result in relatively large yields of coke, such as for example about 4 to about 14% by weight based on fresh feed, more com:nonly about 6 to about 13% and most frequently about 10 to about 13%. The coke yield can more or less quantitatively deposit upon the catalyst. At conter.latd catalyst to oil ratios, the resultant coke laydown may be in excess of about 0.3, more commonly in excess of about 0.5 and very frequently in excess of about 18 of coke by weight, based on the weight of moisture free regenerated catalyst. Such coke laydown may range as high as about 2t, or about 34, or even higher. In common with =onventional FCC operations on VGO, the present process includes stripping of spent catalyst after disengagement of the catalyst from product vapors. Persons skilled in the art are acquainted with appropriate stripping agents and conditions for stripping spent catalyst, but in some cases the present process may require somewhat more severe conditions than are commonly employed. This may result, for example, from the use of a carbo-metallic oil having constituents which do not volatilize under the conditions prevailing in the reactor, which constituents deposit themselves at least in part on the catalyst. Such adsorbed, unvaporized material can be troublesome from at least two standpoints. First, if the gases (including vapors) used to strip the catalyst can gain admission to a catalyst disengagement or collection chamber connected to the downstream end of the riser, and if there is an accumulation of catalyst in such chamber, vaporization of these unvaporized hydrocarbons in the stripper can be followed by adsorption on the bed of catalyst in the chamber. More particularly, as the catalyst in the stripper is stripped of adsorbed feed material, the resultant feed material vapors pass through the bed of catalyst accumulated in the catalyst collection and/or disengagement chamber and may deposit coke and/or condensed material on the catalyst in said bed. As the catalyst bearing such deposits moves from the bed and into the stripper and from thence to the regenerator, the condensed products can create a demand for more stripping capacity, while the coke can tend to increase regeneration temperatures and/or demand greater regeneration capacity. For the foregoing reasons, it is preferred to prevent or restrict contact between stripping vapors and catalyst accumulations in the catalyst disengagement or collection chamber. This may be done for example by preventing such accumulations from forming, e.g. with the exception of a quantity of catalyst which essentially drops out of circulation and may remain at the bottom of the disengagement and/or collection chamber, the catalyst that is in circulation may be removed from said chamber promptly upon settling to the bottom of the chamber. Also, to minimize regeneration temptratures and demand for regeneration capacity, it may be desirable to employ conditions of time, temperatura and atmosphere in the stripper which are sufficient to reduce potentially volatile hydrocarbon material borne by the stripped catalyst to about 104 or less by weight of the total carbon loading on the catalyst. Such stripping may for example include preheating of the catalyst, extensive stripping with steam, the use of gases having a temperature considered higher than normal for FCC/VGO operations, such as for instance flue gas from the regenerator, as well as other refinery stream gases such as hydrotreater off-gas (H25 containing), hydrogen and others. For example, the stripper may be operated at a temperature of about 3500F using steam at a pressure of about 15C psig and a weight ratio of steam to catalyst of about 0.002 to about 0.003. On the other hand, the stripper may be operated at a temperature of about 10250F or higher. Substantial conversion of carbo-metallic oils to lighter products in accordance with the invention tends to produce sufficiently large coke yields and coke laydown on catalyst to require some care in catalyst regeneration. In order to maintain adequate activity in zeolite-and non-zeolite catalysts, it is desirable to regenerate the catalyst under conditions of time, temperature and atmosphere sufficient to reduce the percent by weight of carbon remaining on the catalyst to about 0.25% or less. The amounts of coke which must therefore be burned off of the catalysts when processing carbo-metallic oils are usually substantially greater than would be the case when cracking VGO. The term coke when used to describe the present invention, should be understood to include any residual unvaporized feed or cracking product, if any such material is present on the catalyst after stripping. Regeneration of catalyst, burning away of coke deposited on the catalyst during the conversion of the feed, may be performed at any suitable temperature in the range of about 11000 to about 16000F, measured at the regenerator catalyst outlet. This temperature is preferably in the range of about 12000 to about 15000r, more preferably about 12750 to about l425F and optimally about 1325 to about 12750F. The process has been operated, for example, with a fluidized regenerator with the temperature of the catalyst dense phase in the range of about 13000 to about 1400"F. In accordance with the invention, regeneration is conducted while maintaining the catalyst in one or more fluidized beds in one or more fluidization chambers. Such fluidized bed operations are characterized, for instance, by one or more fluidized dense beds of ebulliating particles having a bed density of, for example, about 25 to about 50 pounds per cubic foot. Fluidization is maintained by passing gases, including combustion supporting gases, through the bed at a sufficient velocity to maintain the particles in a fluidized state but at a velocity which is sufficiently small to prevent substantial entrainment of particles in the gases. For example, the lineal velocity of the fluidizing gases may be in the range of about 0.2 to about 4 feet per second and preferably about 0.2 to about 3 feet per second. The average total residence time of the particles in the one or more beds is substantial, ranging for example from about 5 to about 30, more preferably about 5 to about 20 and still more preferably about 5 to about 10 minutes. Heat released by combustion of coke in the regenerator is absorbed by the catalyst and can be readily retained thereby until the regenerated catalyst is brought into contact with fresh feed. When processing carbo-metallic oils to the relatively high levels of conversion involved in the present invention, the amount of regenerator heat which is transmitted to fresh feed by way of recycling regenerated catalyst can substantially exceed the level of heat input which is appropriate in the riser for heating and Vaporz- ing the feed and other materials, for supplying the endothermic heat of reaction for cracking, for making up the heat losses of the unit and so forth. Thus, the amount of regenerator heat transmitted to fresh feed may be controlled, or restricted where necessary1 within certain approximate ranges. The amount of heat so transmitted may for example be in the range of about 500 to about 1200, more particularly about 600 to about 900, and more particularly about 650 to about 850 BTUs per pound of fresh feed. The aforesaid ranges refer to the combined heat, in BTUs per pound of fresh feed, which is transmitted by the catalyst to the feed and reaction products (between the contacting of feed with catalyst and the separation of product from catalyst) for supplying the heat of reaction (e.g. for cracking) and the difference in enthalpy between the products and the fresh feed. Not included in the foregoing are the heat made available in the reactor by the adsorption of coke on the catalyst, nor the heat consumed by heating, vaporizing or reacting recycle streams and such added materials as water, steam naphtha and other hydrogen donors, flue gases and inert gases, or by radiation and other losses. One or a combination of techniques may be utilized for controlling or restricting the amount of regeneration heat transmitted via catalyst to fresh feed. For example, one may add a combustion modifier to the cracking catalyst in order to reduce the temperature of combustion of coke to carbon dioxide and/or carbon monoxide in the regenerator. Moreover, one may remove heat from the catalyst through heat exchange means, including for example heat exchangers (e.g. steam coils) built into the regenerator itself, whereby one may extract heat from the catalyst during regeneration. Heat exchangers can be built into catalyst transfer lines, such as for instance the catalyst return line from the regenerator to the reactor, whereby heat may be removed from the catalyst after it is regenerated. The amount of heat imparted to the catalyst in the regenerator may be restricted by reducing the amount of insulation-on the regenerator to permit some heat loss to the surrounding atmosphere, especially if feeds of exceedingly high coking potential are planned for processing; in general, such loss of heat to the atmosphere is considered economically less desirable than certain of the other alternatives set forth herein. One may also inject cooling fluids into portions of the regenerator other than those occupied by the dense bed, for example water and/or steam, whereby the amount of inert gas available in the regenerator for heat absorption and removal is increased. Another suitable and preferred technique for controlling or restricting the heat transmitted to fresh feed via recycled regenerated catalyst involves maintaining a specified ratio between the carbon dioxide and carbon monoxide formed in the regenerator while such gases are in heat exchange contact or relationship with catalyst undergoing regeneration. Still another particularly preferred technique for controlling or restricting the regeneration heat imparted to fresh feed via recycled catalyst involves the diversion of a portion of the heat borne by recycled catalyst to added materials introduced into the reactor, such as the water, steam, naphtha, other hydrogen donors, flue gases, inert gases, and other gaseous or vaporizable materials which may be introduced into the reactor. In most circumstances, it will be important to insure that no adsorbed oxygen containing gases are carried into the riser by recycled catalyst. Thus, whenever such action is considered necessary, the catalyst discharged from the regenerator may be stripped with appropriate stripping gases to remove oxygen containing gases. Such stripping may for instance be conducted at relatively high temperatures, for example about 13500 to about 1370 oF, using steam, nitrogen or other inert gas as the stripping gas(es). The use of nitro gen and other inert gases is beneficial from the standpoint of avoiding a tendency toward hydro-thermal catalyst deactivation which may result from the use of steam. The following comments and discussion relating to metals management, carbon management and heat management may be of assistance in obtaining best results when operating the invention. Since these remarks are for the most part directed to what is considered the best mode of operation, it should be apparent that the invention is not limited to the particular modes of operation discussed below. Moreover, since certain of these comments are necessarily based on theoretical considerations, there is no intention to be bound by any such theory, whether expressed herein or implicit in the operating suggestions set forth hereinafter. Although discussed separately below, it is readily apparent that metals management, carbon management and heat management are inter-related and interdependent subjects both in theory and practice. While coke yield and coke laydown on catalyst are primarily the result of the relatively large auantities of coke precursors found in carbo-metallic oils, the production of coke is exacerbated by hi metals accumu lations, which can also significantly affect catalyst performance. Moreover, the degree of success experienced in metals management and carbon management-will have a direct influence on the extent to which heat management is necessary. Moreover, some of the steps taken in support of metals management have proved very helpful in respect to carbon and heat management. As noted previously the presence of a large teavy metals ac cu..ulation on the catalyst tends to aggravate the problem of dehydrogenation and aromatic condensation, resulting in incrcased production of gases and coke for a feedstock of a given Ramsbottom carbon value. The introduction of substantial quantities of 1120 into the reactor, either in the form of steam or liquid water, appears highly beneficial from the standpoint of keeping the heavy metals in a less harmful form, i.e. the oxide rather than metallic form. This is of assistance in maintaining the desired selectivity. Also, a unit design in which system components and residence times are selected to reduce the ratio of catalyst reactor residence time relative to catalyst regenerator residence time will tend to reduce the ratic of the times during which the catalyst is respectively under reduction conditions and oxidation conditions. This too can assist in maintaining desired levels of selectivity. Whether the metals content of the catalyst is being manaced successfully may be observed by monitoring the total hydrogen plus methane produced in the reactor and/or the ratio of hydrogen to methane thus produced. In general, it is considered that the hydrogen to methane mole ratio should be less than about 1 and preferably about 0.6 or less, with about 0.4 or less being considered about optimum. In actual practice the hydrogen to methane ratio may range from about 0.5 to about 1.5 and average about 0.8 to about 1. Careful carbon management can improve both selectivity (the ability to maximize production of- valuable products), and heat productivitv. In general, the techniques of metals control described above are also of assistance in carbon manaqe- ment. The usefulness of water addition in respect to carbon management has already been spelled out in considerable detail in that part of the specification which relates to added materials for introduction into the reaction zone. In general, those techniques which improve dispersion of the feed in the reaction zone should also prove helpful, these include for instance the use of fogging or misting devices to assist in dispersing the feed. Catalyst to oil ratio is also a factor in heat manaacment. In common with prior FCC practice on VGO, the reactor tem perature may be controlled in the practice of the present invention by respectively increasing or decreasing the flow of hot regenerated catalyst to the reactor in response to decreases and increases in reactor temperature, typically the outlet temperature in the case of a riser type reactor. Where the automatic controller for catalyst introduction is set to maintain an excessive catalyst to oil ratio, one can expect unnecessarily large rates of carbon production and heat release, relative to the weight of fresh feed charged to the reaction zone. Relatively high reactor temperatures are also beneficial from the standpoint of carbon management. Such higher tem pera=res foster more complete vaporization of feed and disengagement of product from catalyst. Carbon management can also be facilitated by suitable restriction of the total pressure in the reactor and the partial pressure of the feed. In general, at a given level of conversion, relatively s.. ll decreases in the aforementioned pressures can substantially reduce coke production. This may be due to the fact that restricting total pressure tends to enhance vaporization of high boiling components of the feed, encourage cracking and facilitate disengagement of both unconverted feed and higher boiling cracked products from the catalyst. It may be of assistance in this regard to restrict the pressure drop of equipment downstream of and in communi cation with the reactor. But if it is desired or necessary to operate the system at hicher total pressure, such as for instance because of operating limitations (e.g. pressure drop in downstream eauiomcnt) the above described benefits may be obtained by restrictinq the feed partial pressure. Suitable ranges for total reactor pressure and feed partial pressure have been set forth above. and in general it is desirable to attempt to minimize the pressures within these ranges. The abrupt separation of catalyst from product vapors and unconverted feed (if any) is also of great assistance, It is for this reason that the so-called vented riser-apparatus and technique disclosed in U.S. Patents 4,070,159 and 4,066,533 to George D. Myers et al is the preferred type of apparatus for conducting this process. For similar reasons, it is beneficial to reduce insofar as possible the elapsed time between separation of catalyst from product vapors and the commencement of stripping. The vented riser and prompt stripping tend to reduce the opportunity for coking of unconverted feed and higher boiling cracked products adsorbed on the catalyst. A particularly desirable mode of operation from the standpoint of carbon management is to operate the process in the vented riser using a hydrogen donor if necessary, while maintaining the feed partial pressure and total reactor pressure as low as possible, and incorporating relatively large amounts of water, steam and if desired, other diluents, which provide the numerous benefits discussed'in greater detail above. Moreover, when liquid water, steam, hydrogen donors, hydrogen and other gaseous or vaporizable materials are fed to the reaction zone, the feeding of these materials provides an opportunity for exercising additional control over catalyst to oil ratio. Thus, for example, the practice of increasing or decreasing the catalyst to oil ratio for a given amount of decrease or increase in reactor temperature may be reduced or eliminated by substituting either appropriate reduction or increase in the charginq ratios of the water, steam and other caseous or vaporizable material, or an appropriate reduction or increase in the ratio of water to steam and/or other gaseous matcrials introduced into the reaction zone. Heat management includes measures taken to control the amount of heat released in various parts of the process and/or for dealing successfully with such heat as may be released. Unlie conventional FCC practice using VGO, wherein it is usu ally a problem to generate sufficient heat during regeneration to heat balance the reactor, the processing of carbometallic oils generally produces so much heat as to require careful management thereof. Heat management can be facilitated by various techniques associated with the materials introduced into the reactor. Thus, heat absorption by feed can be maximized by minimum preheating of feed, it being necessary only that the feed temperature be high enough so that it is sufficiently fluid for successful pumping and dispersion in the reactor. When the catalyst is maintained in a highly active state with the suppression of coking (metals control), so as to achieve higher conversion, the resultant higher conversion and greater selectivity can increase the heat absorption of the reaction. In general, higher reactor terseratures promote catalyst conversion activity in the face of more refractory and higher boiling constituents with high coking potentials. 'hile the rate of catalyst deactivation may thus be increased, the higher temperature of operation tends to offset this loss in activity. Higher temperatures in the reactor also contribute to enhancement of octane number, thus offsetting the octane depressant effect of high carbon lay down Other techniques for absorbing heat have also been discussed above in connection with the introduction of water, stem, and other gaseous or vaporizable materials into the reactor. As noted above, the invention can be practised in the abovedescribed mode and in many others. An illustrative, nonlimiting example is described by the accompanying schematic diagrams in the figures and by the description of these figures which follows. Referring in detail to the drawings, in Fig. 1 petroleum feedstock is introduced into the lower end of riser reactor 2 through inlet line 1 at which point it is mixed with hot regenerated catalyst coming from regenerator 9 through line 3. The feedstock is catalytically cracked in passing up riser 2 and the product vapors are separated from spent catalyst in vessel 8. The catalyst particles move upwardly from riser 2 into the space within vessel 8 and fall downwardly into dense bed 16. The cracking products together with some catalyst fines pass through horizontal line 4 into cyclone 5. The gases are separated from the catalyst and pass out through line 6. The catalyst fines drop into bed 16 through dipleg 19. The spent catalyst, coated with coke and vanadium in a reduced state, passes through line 7 into upper dense fluidized bed 18 within regenerator 9. The spent catalyst is fluidized with a mixture of air, CO and CO2 passing through porous plate 21 from lower zone 20. The spent catalyst is partially regenerated in bed 18 and is passed into the lower portion of vented riser 13 through line 11. Air is introduced into riser 13 through line 12 where it is mixed with partially regenerated catalyst. The catalyst is forced rapidly upwards through the riser and it falls into dense settled bed 17. Line 14 provides a source of reducing gas such as CO for bed 17 to keep the regenerated catalyst in a reducing atmosphere and thus keep vanadium present in a reduced oxidation state. Regenerated catalyst is returned to the riser reactor 2 through line 3, which is provided with a source of a reducing gas such as CO through line 22. In Fig. 2, spent catalyst coatea with coke and vanadium in a reduced state fiows into dense fluidized bed 32 of regenerator 31 through inlet line 33. Air to combust the coke and fluidize the catalyst is introduced through line 34 into air distributor 35. Coke is burned and passes upwardly into riser regenerator 36. The partially regenerated catalyst which reaches the riser 35 is contacted with air from line 37 which completes the regeneration. The regenerated catalyst passes upwardly from the top of the riser 36 and falls down into dense settled bed 37. Dense bed 37 and the zone above 37 through which the regenerated catalyst falls are supplied with a reducing gas such as CO through lines 40 and 41. The regenerated catalyst is returned to the cracking reactor through line 38. The COrich flue gases leave the regenerator through line 39. Having thus described this invention, the following Example is offered to illustrate it in more detail. Example A carbo-metallic feed at a temperature of about 4000F is fed at a rate of about 2000 pounds per hour into the bottom of a vented riser reactor where it is mixed with a zeolite catalyst at a temperature of about 12750F and a catalyst to oil ratio by weight of about 11. The carbo-metallic feed has a heavy metal content of about 5 ppm Nickel Equivalents, including 3 ppm vanadium, and has a Conradson carbon content of about 7 percent. About 85 percent of the feed boils above 6500F and about 20 percent of the feed boils above 10250F. The temperature within the reactor is about 1000F and the pressure is about 27 psia. About 75 percent of the feed is converted to fractions boiling at a temperature less than 430"F and about 53 percent of the feed is converted to gasoline. During the conversion, about 11 percent of the feed is converted to coke. The catalyst containing about one percent by weight of coke contains about 20,000 ppm Nickel Equivalents including about 12,000 ppm vanadium. The catalyst is stripped with steam at a temperature of about 10000F to remove volatiles and the stripped catalyst is introduced into the upper zone of the regenerator as shown in Fig. 1 at a rate of about 23,000 pounds per hour, and is partially regenerated to a coke concentration of about 0.2 percent by a mixture of air CO and CO2. The CO/CO2 ratio in the fluidized bed in the upper zone is about 0.3. The partially regenerated catalyst is passed to the bottom of a riser reactor where it is contacted with air in an amount sufficient to force the catalyst up the riser with a residence time of about 1 second. The regenerated catalyst, having a coke loading of about 0.05 percent exits from the top of the riser and falls into a dense bed having a reducing atmosphere comprising CO. The regenerated catalyst is recycled to the riser reactor for contact with additional feed.

Claims: Claims 1. A process for converting a vanadium-containing hydrocarbon oil feed to lighter products comprising: contacting said oil feed under conversion conditions with a cracking catalyst to form lighter products and coke, whereby vanadium in an oxidation state less than +5 and coke are deposited on said catalyst; separating said lighter products from the spent catalyst carrying vanadium in an oxidation state less than +5 and coke; regenerating said spent catalyst by contacting it with an oxygen-containing gas under conditions whereby said coke on the spent catalyst is combusted, forming gaseous products comprising CO and CO2, said regeneration being carried out under conditions whereby vanadium is maintained in an oxidation state less than +5; and recycling the regenerated catalyst to the regenerator to contact fresh feed. 2. A process according to claim 1 wherein said feed contains 6500±+ material characterized by a carbon residue on pyrolysis of at least about 1 and a Nickel Equivalent content of heavy metals of at least about 4. 3. A process according to claim 2 wherein said 6500F+ material represents at least about 70% by volume of said feed and includes at least about 10% by volume of material which will not boil below about 10000F. 4. The process of claim 1 wherein the feed contains at least about 1 ppm vanadium. 5. The process of claim 1 wherein the cracking catalyst comprises a zeolite molecular sieve catalyst containing from about 1 to about 60% by weight of sieve. 6. The process of claim 1 wherein said catalyst is regenerated in at least two stages, in the first stage of which said spent catalyst is contacted in a dense fluidized bed with a gas containing less than a stoichiometric amount of oxygen to convert the hydrogen in said coke to H2O and the carbon in said coke to CO2, and in the final regeneration stage of which partially regenerated catalyst is contacted with a stoichiometric excess of oxygen for a period of time of less than about 2 seconds. 7. The process of claim 6 wherein the catalyst in said final stage comprises a dispersed phase having a density less than about 4 pounds per cubic foot, the residence time of the catalyst in said dense fluidized bed is at least about 5 minutes and said fluidized bed has a density from about 25 to about 50 pounds per cubic foot. 8. The process of claim 6 wherein the partially regenerated catalyst is contacted with at least a stoichiometric amount of oxygen in a riser regenerator, the residence time of the catalyst in the riser regenerator is less than about 2 seconds, and the regenerated catalyst is separated from the gaseous products. 9. The process of claim 8 wherein the separated, regenerated catalyst is immediately contacted with a reducing gas and is then collected in a dense bed maintained under a reducing atmosphere. 10. The process of claim 8 wherein the regenerated catalyst is separated from the gaseous products by being projected in a direction established by the riser regenerator, or an extension thereof, while the gaseous products are caused to make an abrupt change of direction resulting in an abrupt, substantially instantaneous ballistic separation of gaseous products from regenerated catalyst.

Inventors: CARRUTHERS, JAMES D., HETTINGER, WILLIAM P., JR., WATKINS, WILLIAM D.
Applicants: ASHLAND OIL, INC.
Application year: 1982
Main classification: C10G11
IPC codes: C10G11