LARGE PORE CATALYSTS FOR HEAVY HYDROCARBON CONVERSION

Abstract: A catalyst, a method for its preparation, and a process for its use in the production of high octane gasoline and/or other valuable liquid hydrocarbon products from carbo-metallic oils contaminated with heavy metals and coke precursors. The catalyst is prepared from a shapable composition containing reticulated carbon black particles which are subsequently removed by burning, as in a catalyst regenerator, to provide large feeder pores in a shaped catalyst structure. The catalyst and process conditions for its use facilitate conversion of the contaminated oils to lighter products at catalytic sites within the catalyst structure which may contain zeolitic materials.

Description: LARGE PORE CATALYSTS FOR HEAVY HYDROCARBON CONVERSION Technical Field This invention relates to an improved catalyst, a method for its preparation, and a process for its use in the conversion of petroleum oil feeds containing hydrocarbon molecules of high molecular weight (heavy) . More particularly, the invention is related to a catalyst composition comprising a catalytically active crystalline aluminosilicate zeolite uniformly dispersed within a matrix component having large feeder pores for convey¬ ing reactants to and reaction products from the zeolitic com¬ ponent. Background of the Invention In general, gasoline and other liquid hydrocarbon fuels boil in the range of about 100°F to 650°F. However, the crude oils fro which these fuels are made contain mixtures of hydrocarbons whi boil over wider temperature ranges, the boiling point of each h drocarbon depending upon its molecular weight. As an alterna¬ tive to discarding or otherwise using the higher boiling hydro¬ carbons, the petroleum refining industry has developed a variet of processes for breaking or cracking the large molecules of hi molecular weight into smaller molecules which boil within the above boiling range for hydrocarbon fuels. The cracking pro¬ cess which is most widely used for this purpose at the present time is known as fluid catalytic cracking (FCC) and may employ a fluidized bed reactor with backmixing and/or a riser reactor with progressive flow. In a typical FCC process, feedstock oil is mixed with particulate catalyst at an elevated temperature in the lower portion of an elongated reaction vessel called a "rise Contact of the hot-catalyst with the oil rapidly generates large volumes of gases which propel the stream of feed and catalyst as a suspension through the reaction zone at high velocity, giving relatively short contact times. The initial propelling gases are comprised of vaporized oil, the major portion of which boils below 1,025°F and is immediately vaporized by contact with the hot catalyst which enters the rise at a higher temperature. As the suspension travels up the riser a large fraction of the feedstock hydrocarbons is converted to lower boiling hydrocarbons by catalytic cracking and these crac products form part of the propelling gases. The velocity of th suspension is sometimes increased further by introducing diluen materials into the riser either along with the feed or separate The conversion reaction initiated in the lower portion of the riser continues until the catalyst and gases are separated, whi may take place as the suspension leaves the riser reaction zone or in an upper, larger diameter vessel for collecting the cata¬ lyst. Upon being separated from the catalyst, the gases are us ally referred to as "product vapors". Crude oil in its natural state contains a variety of materials which, unless removed prior to the cracking reaction, tend to have troublesome effects on FCC processes. These include coke precursors, such as asphaltenes, polynuclear aromatics and high boiling nitrogen containing molecules; and metals, such as sodiu and small amounts of other .alkali or alkaline earth metals, nick vanadium, iron and copper, which are detrimental to the convers process and/or to the catalyst. ' During the cracking operation, coke precursors either tend to d posit as solid aromatic structures having some residual hydroge or are high boiling and do not vaporize but lay down on the cat lyst as a liquid. These coke deposits block the catalytically active acid sites of the catalyst and thereby reduce its conver sion activity. While it is believed that both the solid and li quid components of coke may cover and thereby block acidic site the liquid components may also fill pores of the matrix and the by retard diffusion of lower boiling components to the zeolite. Although the carbonaceous material formed by the conversion pro cess is referred to as coke, it may have hydrogen to carbon ratios of 1.0 or greater and may contain in addition to hydroge various amounts of other element depending upon the composition of the feed. The coke formed is deposited on the catalyst part cles and thereby reduces the conversion activity of the catalys In order to restore conversion activity, the contaminated cata¬ lyst is regenerated by burning off the coke by contacting thar. J the catalyst particles at relatively high temperatures with an oxidizing gas such as air. The regenerated catalyst may then be returned to the reaction zone for additional passes or con¬ version cycles in contact with fresh feed. In general, the coke-forming tendency or coke precursor content of a feedstock oil can be ascertained by determining the weight percent of residue remaining upon pyrolyzing a sample of the feed. Two tests presently recognized by the industry are the Conradson carbon residue test described in ASTM D189-76 and the Ramsbottom carbon test described in ASTM D524-76. In conven¬ tional FCC practice, Conradson carbon residues of about 0.05 to 1.0 are regarded as indicative of relatively contaminate free gas oil feeds. Unless removed by careful desalting of the crude oil, the sodiu and other alkali or alkaline earth metals can diffuse to the acti i.e., acidic, sites of the catalyst and poison or kill their catalytic activity. Vanadium, and to a lesser extent nickel an other metals, may also " migrate to and poison acidic sites. The metals will therefore be referred to collectively as poison met Nickel, vanadium, copper and iron are also known as "heavy meta and catalyze undesirable dehydrogenation reactions which increa the amount of coke deposited on the catalyst and the amounts of hydrogen and normally gaseous hydrocarbons to be handled by pro cess equipment. During the cracking process, the heavy metals transfer almost quantitively from the feedstock oil to the cata lyst particles and tend to deposit on interior. and exterior sur faces of the particles where they can block and/or retard dif¬ fusion to active sites. Since the various heavy metals are not of equal poisoningactivi relative to catalytic acid sites, it is convenient to express the poisoningactivity of an oil containing one or more of these metals in terms of the amount of a single metal which is esti mated to have equivalent poisoning activity. Thus, the heavy metals content may be expressed by the following formula in which the content of each metal present is expressed in parts per million by weight based on the weight of the oil: Nickel Equivalents = Ni + V/4.8 + Fe/7.1 + Cu/1.23. In conventional FCC practice, crude oils are carefully fractionated to provid a gas oil with a relatively low level of heavy metal contami¬ nants, namely, 0.25 ppm Nickel Equivalents or less. The above formula can also be used as a measure of the heavy metals accumulated on the cracking catalyst itself, the quant ty of metal used in the formula being based on the weight of moisture-free catalyst. In FCC practice, equilibrium catalys is removed and fresh, contaminant-free catalyst is added at a rate sufficiently high to control the heavy metal content of the catalyst at relatively low levels, namely, 1,000 ppm Nick Equivalents or less. Some crude oils contain from about 10% to about 30% by volume of heavy hydrocarbons which will not boil below about 1,025°F at atmospheric pressure. Atmospheric bottoms and vacuum bot¬ toms may contain even higher percentages of this highest boil ing fraction. The coke precursor and poison metal components of the crude are for the most part concentrated in this frac¬ tion. Accordingly, many of the problems presented by these components have been avoided by sacrificing the yield of liqu fuel fractions which is potentially available from cracking the highest boiling fraction. More particularly, in conven¬ tional FCC practice, the crude oil has been vacuum fractionat to provide a FCC feedstock boiling between about 650°F and about 1,000°F, this fraction being referred to as a vacuum ga oil and being relatively free of coke precursors and poison metals. Vacuum gas oil is generally prepared from crude by d tillinc off the fraction boiling below about 650°F at atmcA-. _ pheric pressure and then separating by vacuum distillation on or more fractions boiϋinσ between about 650°F and about 1,025 from the heaviest fraction boiling above 1,025°F. The hea est fraction is normally not used as a source of catalytic co version feedstock, but instead is employed for other purposes such as the production of asphalt, which represents a waste the potential value of this portion of the crude oil as a sou of liquid fuels. Due to the continually increasing demand for gasolines, relative to heavier liquid fuels, coupled with shrink¬ ing supplies of normally used gas oil cracking stocks, more attention has recently been given to the catalytic crack ing of heavier chargestocks, such as residuals from which the highest boiling fraction has not been separated. In addition, consideration has been given to blending the heaviest or "resi fraction with various lower boiling fractions in order to in¬ crease overall conversion of crude oil to liquid fuels. In vi of the high potential value of the heaviest fraction of crude, oils, a number of methods have been proposed in the past to overcome the problems associated with the cracking of feedstoc contaminated with metals and coke precursors and thereby in¬ crease the overall yield of gasoline and other hydrocarbon fuels from a given quantity of crude oil. Suggestions have been made to pretreat the contaminated feed to reduce the meta content to below about 4 ppm nickel equivalents and the Conrad son carbon residue to below about 1. Various demetalization techniques -have also been suggested for removing the metal con taminants once they have been deposited on the catalyst. Most of these prior art techniques, however, require expensive ad¬ ditional equipment and materials and cannot be justified from an economic standpoint. Attention has also been given to developing improved catalysts for cracking more contaminated feeds. However, many problems have been encountered in the use of. prior art catalysts for cracking feeds containing a resid fraction. A catalyst comprised o f crystalline zeolite particles embedded within a larger matrix particle has numerous passages leading from the outer peripheral surface of the matrix particle to the smaller zeolite particles supported within the matrix . In this speci ication , these matrix passages are referred to as " feeder pores " . Feeder pores in effect provide access passageways from the surface of each catalyst particle to those zeolite particles at locations internal to the matrix . There also may be a small but finite number of zeolite particles exposed at the surface of the matrix . Generally , the pores of zeolitic sieves fall within the range of 4 to 13A° . Accordingly , any pores larger than 13A° are usually in the matrix . In prior art catalysts of this type , the average diameter of feeder pores in fresh catalyst usually falls within the range of about 30A° to about 400A° . Alumina-silica matrices generally have pores in this range , although a relatively small proportion may be larger . Eowever , after an extended period of use , the effective average pore diameter of these prior art catalysts may be decreased significantly because of coke and metal accumulations . These prior art catalysts have proven in¬ efficient for cracking res id containing feedstocks for a number of reasons , including both a low zeolite utilization factor and undesirable reaction diffusion limitations . The low zeolite utilization effect is a consequence primarily of the deposition of both coke and heavy metals in and/or across the mouths of the working pores of the zeolite . These components in effect block off zeolitic pore vo lume containing unused or incompletely used acidic sites . " Z eolitic pore vo lume " refers to the free volume of the micro- pores in the zeolite component rather than the matrix. The term "pore volume" as applied to the catalyst compo sition as a who le re fers to the free volume in the matrix and zeolite combined which is provided by both acropores (pores having a minimum diameter above 30A° ) and icropores (pores having a minimum diameter of 30A° ore les s ) . The pore volume fraction for pores greater tha 30A° in diameter may be determined by mercury methods , such as the method of U. S . Patent No. The pore volume fraction in the 0 to 3ΘA° range may be determined by the BET nitrogen adsorption method described by Brunauer, Em ett, and Teller in the Journal of the American Chemical Society, 60, 309 (1938). The pore volume of fresh hydrocar¬ bon conversion catalyst may vary widely depending upon the s of the pores in the matrix and, where used, the relative amount of catalytic promoter, such as zeolite, and the size of its pores. Diffusion limitations may result from a number of different mechanisms. One such limitation is a consequenc of high molecular weight molecules in the feed and the absenc of a sufficient number of feeder pores of the size range re¬ quired for transporting these large molecules to the acidic sites of the catalyst, some of which may be in the matrix but the majority of which are in the zeolite. Another diffusion limitation in the processing of reduced crude or other resid containing feeds is due to what may be called "pore plugging" Pore plugging is caused by the absorption of unvaporized hy¬ drocarbons " in the catalyst pores so that they are impractical to remove by stripping operations prior to regeneration. The trapping of heavy hydrocarbons which cannot be removed by con ventional stripping operations can lead to excessive coke and r eration temperatures and increased air consumption. Pore plugging and the deposition of coke and/or heavy metals with in or over the pores also leads to decreased diffusion of re- actants to and products from acidic sites. Slow diffusion rates may result in thermal cracking predominating over cata¬ lytic cracking, which in turn causes loss of selectivity. Thus, catalysts possessing relatively small or restricted feeder pores will show relatively poor cracking characteristi when cracking resid containing feeds, including low conver¬ sion, poor selectivity, increased air consumption during re¬ generation, and higher regeneration temperatures. Hot spots also occur more readily during regeneration and cause cata¬ lyst deactivation through sintering of the matrix and loss of zeolite crystalline structure and acidity. Furthermore, low catalyst utilization factors and diffusion limitations both require hiσh catalvst to oil ratios which necessitate tively low oil feed rates . In order to provide economic levels of conversion activity an more importantly the selectivity required in processing the very refractory hydrocarbons found in resid fractions, it is desirable to run the riser at a relatively high temperature. In addition, large amounts of coke accumulate on the catalyst The primary problem with this increased coke make is that the reactions in the regenerator which convert coke to carbon monoxide and carbon dioxide are highly exothermic. Since the regeneration reactions areexothermic, the regeneration step is normally carried out at a temperature much above the cracking temperature in the riser. This makes it necessary to run the regenerator at maximum temperatures in order to burn the coke off the catalyst to the relatively low levels required for restoration of its activity. To achieve a heat balance in cracking resid, it is therefore necessary to operate the re¬ generator at very severe hydrothermal conditions, which can cause rapid degradation of many prior art catalysts. At high regenerator temperatures, excess heat and localized h spots may develop within catalyst particles, especially in places where pore plugginghas occurred or excessive coke de¬ posits have accumulated. These localized hot spots result in sintering and collapse of the matrix pore structure, thus ren ering a large portion of the acidic sites in the matrix unava able for further reactant contact. Where a catalytic promote is used, the promoter will necessarily be entrapped within th collapsed pores of the matrix and blocked off from further re active contact. Coke from resid molecules can also cover and block portal surface areas of both the matrix and the ' zeolite. The crystalline structure of zeolites is susceptible to degra dation by high regenerator temperatures per se. Zeolites are crystalline alumino-silicates made up of tetra-coordinated aluminum atoms associated through oxygen atoms with silicon atoms in an ordered crystalline structure. Localized hot spo in or near the zeolite particles can cause destruction of i- aluminosilicate crystalline structure, at least to the ex¬ tent of destroying portal area of the zeolite, with a re¬ sulting loss of its catalytic action. Furthermore, both sodium and vanadium contaminates accelerate sintering and collapse of pore structures in both the matrix and zeolite components. Such degradation permanently deactivates the catalyst so that it must be removed from the system, re¬ sulting in high make-up rates that may prove uneconomical because of the high cost of zeolite in the catalyst. There is a need therefore for a heat resistant zeolite catalyst suitable for use in cracking resid containing oil feeds with improved overall utilization of acidic sites and a minium of diffusion limitations. While it has been recognized in the past that the physical structure of catalyst particles plays an important role in their effectiveness, the extent to which such structure is im¬ portant generally has been obscured by the lack of analytical techniques for isolating the complex mechanisms involved in catalytic cracking. In this connection, some attention has been given in the prior art to increasing the pore size of catalyst matrices. Thus, it has been suggested that extremely large pores, such as those with diameters above 1,000A°, migh be introduced into a catalyst by incorporating a removable ma¬ terial and subsequently removing that material during catalyst preparation. See, for example, the catalysts described in U.S Patent No. 2,890,162 to Anderson, et al. , and U.S. Patent No. 3,944,482 to Mitchell, et al. However, the removable material suggested for this purpose have not been easily controllable and have resulted in poorly defined pore structures and a wide variety of pore sizes, relatively few, if any, of the pore diameters being in the actual size range needed for resid cracking. Disclosure of the Invention The present invention provides a catalyst and a process for selective conyersion of heavy hydrocarbons. It is therefore a principal object of the invention to provide a special hy¬ drocarbon conversion catalyst resistant to deactivation by se¬ vere hydro-thermal conditions and by accumulations of coke pre cursors and poison metals. Another object of the invention is to provide an improved process for catalytic conversion of hig boiling carbo-metallic oil feeds containing relatively high co centrations of coke precursors and poison metals. The catalys and process of the invention areparticularly useful for crack¬ ing oil feeds which contain significant quantities of residual hydrocarbons, e.g., at least ten percent, boiling above about 1,025°F, and significant quantities of heavy metals, e.g., at least about 4 ppm of Nickel Equivalents. Feeder pores having large minimum diameters and large mouths are provided in the catalyst of the inventionso as to facilitate diffusion of high molecular weight molecules through the matrix to the portal surface area of the sieve particles. The catalyst matrix also has a relatively large pore volume in order to soak up unvapor ized portions of the oil feed. The feeder pores through the matrix are sufficiently large so that significant numbers of hydrocarbon molecules can diffuse to active catalytic sites both in the matrix and n the sieve. On the other hand, the open channels of these feeder pores are somewhat tortuous and tend to trap molecules having molecular weights of or greater and average diameters of 200A° or greater (for exam¬ ple asphaltenes, porphyrins and polynuclear aromatics) . These very large molecules can effectively cover and block the rela¬ tively small pores of the zeolitic sieve. The size and struc¬ ture of the feeder pores of the invent!on.are such that they remain open without diffusion restrictions even when matrix su faces and pore channel walls become loaded with very large as- phaltene type molecules. * It has been discovered that the optimum average diameter for feeder pores in the catalyst matrix is in the range of 400 to 6,000A°, preferably 1,000 to 6,O00A°. It has also been dis¬ covered that these pore size ranges can be easily and readily obtained with the use of carbon black particles having a cros sectional diameter in the range of about 100 to 1,000A° and a length to diameter ratio greater than about 2:1, preferably greater than about 5:1. Upon firing a catalyst composition containing these carbon black particles, such as during the out of the matrix, leaving large feeder pores within the desire size range. The carbon black may be added at almost any time during the catalyst preparation prior to drying the final cata¬ lyst composite. The fired catalyst has a relatively large and controlled number of feeder pores within the desired size range, the number and size distribution of these pores being enhanced to a significant extent by the type and amount of dispersant use for suspending the carbon black in the aqueous suspension from which the unfired composite- is made. The large feeder pores of the invention extend from the surface the interior of thematri -particles and provide access channels to the much smaller zeolite particles supported within the matri material. Carbon black may be incorporated in the catalyst matr by adding it to a composite catalyst slurry along with the zeo¬ litic component prior to spray drying the slurry to form final microspheres of matrix supported zeolite. Where a kaolin clay c position is spray dried, calcined and then treated to generate z lite crystals in situ, the carbon black.may be added to the clay slurry fed to the spray drier so as to be present in the composi tion during and after in situ formation of the zeolite within th matrix material. A competing consideration for the selectivity desired in crackin residual feeds is that there must be sufficient acid sites prese so that catalytic cracking dominates the conversion reaction. T as average pore size and pore volume increase, catalytic surface area may decrease. To compensate for this decreas in surface ar the amount of zeolitic promoter can be increased and/or smaller zeolite particles (crystals) employed to increase the portal sur face area of this component. A preferred catalyst of the inven¬ tion therefore contains relatively high concentrations of very small zeolitic sieve particles supported within an alumina, sili and/or alumina-silica matrix. Both the sieve and the matrix sho have good steam and thermal stability. A "Y" type zeolite sieve relatively free of sodium and stabilized with hydrogen and/or am¬ monium ions and/or rare earth ions is preferred. According to a preferred embodiment of the invention, spray dried microspheres containing carbon black are partially exchanged with rare earths calcined to remove the carbon black and stabilize the zeolite, a further exchanged with rare earths to provide a catalyst having superior hydrothermal stability. Acidic sites may also be provided in the matrix material so that at least some of the heavier hydrocarbons, both liquid and vapor ous, can be cracked on the surfaces and in the passages of the matrix to provide a means by which molecules larger than the sie pores can be converted to smaller molecules of a size small enou to enter and be cracked in the sieve. For example, alkyl fragme can be stripped from large aromatic molecules so that the fragme may enter the highly active pores of the zeolite. The large feeder pores and process conditions of the present inv tion provide increased conversion and improved selectivity in th cracking of reduced crude and other resid containing oil feeds. this connection, it is believed that these larger pores are capa ble of maintaining adequate diffusion of reactants and reaction ducts while providing sufficient pore volume and surface area in matrix for absorption and retention of poison metals, coke and u vaporizable hydrocarbons. Although not wishing to be bound by a particular theory or hypothesis as to the reasons for the improv ments afforded by the invention, the following matters may con¬ tribute to its success. According to the literature, a normal C . - hydrocarbon molecule h a boiling point above 1500°F, a cross-sectional diameter of abou 4A° and a maximum length of about 50A°. Therefore, with a conve tional catalyst matrix having an effective pore, diameter at equi brium of less * than 50A°, this heavy molecule can enter the avera pore based on its minimum cross-sectional dimension but cannot ent based on its length dimension. Thus some net orientation would required for heavy molecules of a resid fraction to traverse the pores of the matrix and reach the acid sites of the zeo- lite. Statistically, heavy molecules, such as those having mol cular wieghts in the range of 1,000 to 10,000, would exist in a relatively disordered state so that some of the molecules would not be able to enter the matrix of catalysts having relatively small feeder pores. Since the reactants and the products of catalytic conversion must necessarily use the same feeder pore passages and diffuse in opposite directions, the rate of dif¬ fusion into a pore must equal the rate of diffusion out under steady state conditions. It is therefore believed that the larg feeder pores of the invention decrease the necessity for a parti cular molecular orientation and therefore increase diffusion rate through these pores, particularly at equilibrium conditions where effective pore diameters necessarily reflect deposits of coke and heavy metals on the wall of pore channels. The larger feeder pores of the matrix allow coke and metal de¬ position near the surface of the catalyst particles without pore blockage and absorption of heavy liquid hydrocarbons with¬ out pore plugging. Large vaporized hydrocarbon molecules can enter and exit these pores at equilibrium conditions without special orientation and can therefore more easily reach acidic sites which are concentrated in the zeolite but also may be present to a lesser extent in the matrix ' material. In a sense, the large pores are able to "soak up" both metal poisons and liquid coke precursors and in effect neutralize at least a por¬ tion of these contaminants. The high pore volume provides physical space in which to '.'load" these types of contaminants without unduly restricting diffusion, even in the presence of partial sintering. The large pores also facilitate reactions between the trapped material and elements or compounds that may be added to the riser or regenerator in order to passivate the poison-- metals accumulated on the matrix. Unusually high temperatures are needed to crack resid containing OMPI feedstock for the reasons given previously. However, at the hig temperatures proposed, thermal cracking reactions compete with catalytic cracking reactions. The product distribution (selecti ty) for the thermal reaction is quite different from the catalyt reaction and highly undesirable, the thermal reaction yielding much lighter gases, more coke, high boiling gas oils with high Conradson carbon values, and relatively low octane gasoline rang products. The activation energy for catalytic crakcing is con¬ siderably lower than for thermal cracking. With prior art cata¬ lysts, as the coke and metals deposit on the catalyst and the liquid asphaltene components fill the relatively small matrix pores, diffusion of the hydrocarbon reactants to the acidic site of the zeolite is retarded to the extent that diffusion through the matrix becomes rate limiting. As this occurs, the apparent activation energy for catalytic cracking rises and diffusion be¬ comes rate limiting, as this occurs, the activation energy for - thermal cracking, which remains constant, results in greater pro portions of ' the undesirable thermal produce distribution. Becau the hydrocarbon molecules of a residual feed have relatively eas access to the zeolite in the catalyst disclosed, the catalytic reaction is much faster than the thermal reaction, giving higher conversion and the desired catalytic product distribution. In addition, improved diffusion rates for regeneration gases and combustion products within the particles help avoid hot spots du ing catalyst regeneration and thereby reduce sintering of both t matrix and the promoter materials. Coke also tends to deposit nearer to the particle surface where it can be more easily reach by regeneration gases and burned off at lower temperatures. While the pore volume is large, the surface area of the matrix m be correspondingly low. This relatively low surface area, toget with large feeder pores, minimizes the physical and/or chemical tention of vaporizable hydrocarbon molecules so as to facilitate stripping these.molecules from the catalyst and reducing the amo of coke carried into the regenerator. The amount of carbon burn is less and the amount of air necessary for regeneration is ther by reduced. Although the carbon black (CB) catalysts disclosed may be used in a variety of conversion processes employing a wide variety of contacting equipment, it is particularly useful in the catalytic cracking apparatus and process of the inven¬ tion. The apparatus disclosed comprises a progressive flow riser with a ballistic separator at the upper end of the rise for causing a sudden and substantially instantaneous separa¬ tion of catalyst particles from product vapors. The catalyst is then transferred to a stripping vessel for removal of re¬ sidual hydrocarbons removable in the presence of high tempera ture steam and/or other stripping gases. Stripped catalyst is then transferred to a two-stage combuster having insuffi¬ cient oxygen in the first stage to convert all of the carbon to carbon dioxide and an excess of oxygen in the second stage to almost completely burn off the carbon remaining after the first stage. The regenerated catalyst particles attain rela¬ tively high temperatures in the range of about 1,300°F to 1,500°F and have very low levels of residual carbon, namely, 0.05 weight percent or less. The regenerated catalyst parti¬ cles are then returned to the bottom of the riser where they are contacted with fresh feedstock. The conversion process of the invention is particularly effec tive in utilizing the CB catalysts for cracking resid contain ing, carbo- etallic feeds of the type described. Fresh oil feed and a diluent, such as steam, are mixed with the hot regenerated cracking catalyst at or near the bottom of the riser, vaporiz¬ ing and/or fluidizing the feed and diluent substantially in¬ stantaneously to form a vaporous suspension that flows rapidl upward to €he ballistic separator. The temperature and catalyst oil ratio in the suspension is sufficiently se¬ vere to convert approximately 50 to 90% of the carbo-metal¬ lic oil feed to gasoline per pass and produce relatively high levels of coke on the catalyst. The velocity of the suspension combined with ballistic separation is such as to provide very short contact times and avoid overcracking the desired molecular species of the product, notwith¬ standing the high teπuoeratures and very active catalvst. The - 16 - regenerator also is operated at relatively high temperatures which provides rapid and effective coke removal and the heat necessary for the endothermic riser reaction. In view of th high temperatures, the regenerator configuration is such tha the average catalyst hold up time for regeneration is rela¬ tively short, namely, on the. order of about 3 to 5 minutes or le Since relatively small amounts of catalyst are held up in other portions of the system, the overall catalyst inventory is significantly low. Relatively high " oxygen partial pressures are maintained in t regenerator, either in the last stage of multistage regenera tors or in a zone immediately upstream of discharge conduits from single stage regenerators to keep heavy metals on the catalyst in their less active oxide form. The substantially instantaneous fluidization of the oil feed and the very shor residence times employed tend to inhibit reduction of these metals in the riser to their more catalytically active free metal state. It is also contemplated that certain elements, such as antimony, may be added to the regenerator or riser to more permanently tie up accumulated heavy metals. The large pores and smaller surface areas of the CB catalysts facilitate such metals deactivation reactions. Although economically prohibitive at the present time, demetalization techniques also should prove to be more effective with CB catalysts. Because the catalyst is resistant to degradation and process apparatus and parameters are adapted for its effective utili¬ zation, the make-up rates at which fresh catalyst must be introduced into the system are well within acceptable limits, namely, in the range of about 0.1 to 3.0 pounds of catalyst per barrel of fresh feed. More particular process para¬ meters of the invention, together with further particulars on the catalyst and apparatus thereof, are given in the de¬ scription below of the best mode for carrying out the in¬ vention. Brief Description of the Drawings The invention may be further understood by reference to the description of the best mode taken in conjunction with the accompanying drawings in which: Fig. 1 is a diagramatic illustration of the reticulated stru ture of one type of carbon black employed in the catalyst co position of the invention. Fig. 2 is a diagramatic representation of a catalyst particl made according to the invention. Fig. 3 is a diagramatic representation of a feeder pore of t invention enlarged relative to the view of Fig. 2. Fig. 4 is a diagramatic representation of cracking a polynu- clear aromatic hydrocarbon within the matrix of the catalyst of the invention. Fig. 5 is a diagramatic representation of cracking the ali¬ phatic fragment of Fig. 4 in the zeolitic component of the catalyst of the invention. Fig. 6 is a graph showing changes in pore volume ranges rela¬ tive to changes in the amount of carbon black used in making the catalyst of the invention. Fig. 7 is a schematic diagram of an apparatus for utilizing the catalyst and carrying out the process of the invention. Fig. 8 is a schematic diagram of another apparatus for utiliz ing the catalys and carrying out the process of the inventio CMPI - 17B - Figs. 9A, 10A, 11A, and 12A are micrographs made by a scanning electron microscope (SEM) showing reticulated carbon black part cles of the type employed in the catalytic composition of the invention. Figs. 9B, 10B, 11B, and 12B are contour plots and other data fr photo-grammetric topography and correspond to Figs. 9A through 12A, respectively. Fig. 13 is a SEM micrograph showing a catalyst matrix made ac¬ cording to the invention. Fig. 14 is a graph showing the change in catalyst relative acti ty with decreasing vol.% MAT conversion. O - 18 - Best Mode for Carrying Out the Invention THE FEEDSTOCK The catalyst of the invention may be used for treating any h drocarbon feedstock suitable for cracking to lower boiling components or for reforming or other hydrocarbon conversion processes. It is especially useful for cracking oil feeds containing an appreciable amount of high molecular weight co ponents, for example a feedstock with at least 5 weight per¬ cent, preferably at least 10 weight percent, not boiling below 1,025°F. Diffusion limitation problems, in general, become increasingly troublesome with increasing fractions of high molecular weight components in the feedstock. Other factors involved include molecular configuration and the like. The terms "high molecular weight" and/or "heavy" components refe to those hydrocarbon fractions -having a normal boiling point of at least 1,025°F and include non-boiling hydrocarbons, i. materials which may not boil under any conditions. These heavy components have relatively large molecular dimen sions, the largest average dimension of each component being referred to here as its "average diameter". For purposes of selecting a catalyst having an optimum most frequent feeder pore size and distribution, some knowledge of the range of average diameters of the feedstock is desirable. One method of evaluating molecular sizes of large organic molecules is that of J. J. Hermans, et al. (J. Che . Phys, , Vol. 20, page 1360, 1952) , in which the average effective molecular diame¬ ter in Angstroms (A° ) is equal to 7.4 N 1 ' 2 where N is the number of carbon atoms in the molecule. Since various mole¬ cules have a variety of configurations and shapes, this pro¬ cedure gives a relative indication of molecular sizes and provides a useful tool for estimating average molecular sizes and spreads of sizes in feedstocks. Such data may be used for selecting the predominant feeder pore sizes and size dis¬ tribution characteristics of the catalyst to be prepared for given types of feedstocks. This method therefore may be used as the basis for calculatinσ average molecular diameters—*a ? i- OMPI - 19 - ranges of diameters for heavy feeds. Generally, high boiling hy drocarbon feedstock has a relatively wide range of molecular dia meters for the molecules present with significant frequency. Fo example, the average difference between the lower and higher sig nificant diameters may be as high as 250 A°. It is preferred th the ratio of average feeder pore size of the catalyst to average molecular diameter of the 1025°F portion of the feedstock be at least about 2, preferably at least about 5. The composite catalysts disclosed have a high tolerance both to metals and to coke precursors and these catalysts will economi¬ cally crack feedstocks containgin high concentrations of such co taminants. A high metals feedstock for purposes of this inven¬ tion is one having a heavy metal content of at least about 4 ppm of Nickel Equivalents. A high coking feedstock for purposes of this invention is one having a Conradson carbon residue value greater than about 1. The feedstocks for which the invention is particularly useful will have aheavy metal content of at least about 5 ppm Nickel Equivalents and a Conradson carbon residue of at least about 2. The greater the heavy metal content and the Conradson carbon residue, the more advantageous the catalyst and process of this invention becomes. A particularly preferred fee stock for treatment by the process of the invention includes a reduced crude and comprises 70 percent or more of a 650°F+ ma¬ terial having a fraction greater than 20 percent boiling above 1025°F at atmospheric pressure, a metals content of greater than 5.5 ppm Nickel Equivalents and a Conradson carbon residue of greater than 4. This feed may have a hydrogen to carbon ratio of less than about 1.8 and coke precursors in an amount suffi¬ cient to yield about 4 to 14 percent coke by weight based on fresh feed. If the metals and/or Conradson carbon are higher th these values, the feed may be pretreated (but preferably is not) by a hydrotreating step to saturate unsaturated hydrocarbons and by contacting adsorbent particles to remove a portion of the poison metals and carbon precursors. OMPI - 20 - Representative feedstocks contemplated for use with the inve tion include whole crude oils, fractions of crude oils such as topped crude, reduced crude, vacuum fractionator bottoms and other fractions containing heavy residua, coal-derived oils, shale oils, waxes, untreated or deasphalted residua, and blends of such fractions with gas oils and the like. In addition to feedstock per se, added diluent materials may also be charged to the riser to lower the vapor pressure of the oil feed. Diluents will increase the spacevelocity of the process by accelerating the velocity of the oil and de¬ creasing catalyst contact time. Any diluent which is a vapo or becomes a vapor under the conditions in the conversion zone can be used. If the diluent is a hydrocarbon, it shoul desirably have a boiling point below about 650°F, and more preferably it should be a gasoline range hydrocarbon, e.g., naphtha, or lighter, which fractions boil at about 430°F or b low. If added hydrocarbon boils above 650°F, it will itself be considered a portion of the cracking feedstock. Other diluen include various gases such as hydrogen, nitrogen, methane and ethane, and water, which may be charged either as liquid or steam. Such diluents may be added at or near the bottom of or at one or more locations along the riser conversion zone so as to assist in dispersal and fluidization of the cata¬ lyst, dispersal and vaporization of the liquid feedstock, quenching of the catalyst and/or oil suspension, and under some conditions, may increase the cracking -rate and/or im¬ prove the selectivity of the cracking process. With respect to the tolerance levels of the catalyst itself, the heavy metals may accumulate on the catalyst to levels in the range of from about 3,000 to about 20,000 ppm of Nickel Equivalents. Where about 4 to 14 percent of the feedstock is converted to coke, this coke is generally deposited on the catalyst in .amounts in the range of about 0.3 to 3.0 percent by weight of the catalyst. O PI - 21 - THE CATALYST The present invention includes a method of making a catalyst wit a relatively large percentage of pores in size ranges above 100 preferably in the range of 400 to 6,000 A°, more preferably 1,00 to 6,000 A°. The catalyst is especially useful in the processin of reduced crudes and other heavy feeds. The method involves in corporating a relatively large amount of selected carbon black solids into the catalyst matrix material. The carbon black is s sequently removed by converting it to gaseous carbon oxides by oxidation at elevated temperatures. Its removal provides a por volume greater than " 0.10 cc/gm in pores greater than 400 A°, pre erably 0.15 cc/gm with at least 0.10 cc/gm in pores greater than 1,000 A°, more preferably at least 0.20 cc/gm with at least 0.15 cc/gm in pores greater than 1,000 A°. Relative to theprior art, the invention reliably and predictably increases the pore volume comprises of these large pore sizes. Carbon black may be added, along with other solids such as zeoli and clay, during the dormation of the catalyst matrix material. Alternately, the carbon black may be added after the basic cata¬ lyst composition has been prepared but before final drying, as for example, by forming a slurry of prepared catalyst compositio followed by dispersing the carbon black within the slurry and then spray drying the resulting suspension. Carbon black may al be added to the zeolite component either prior to or during the formation of the aluminosilicate crystals. Carbon black differs from other types of carbon, such as char¬ coal, and from other types of fillers such as flour or cellulose fibers, in that it exists as very small reticulated particles and is substantially non-porous in most of its forms. Each reticulated particle is itself comprised of smaller "primary" carbon black particles. Primary carbon black particles are essentially spherical, the diameter of the sphere varying de¬ pending upon the method of manufacture. The primary particles in turn are composed of several thousand microcrystallite bun¬ dles stacked together in a random order and each bundle consists of several polynuclear aromatic platelets which are stacked in a - 22 - not quite parallel manner. The carbon blacks of the present in vention possess to a greater or lesser extent a basic property called "structure". Structure refers to the degree to which th primary particles are bound together into a 3-dimensional prima chain network making up the reticulated particle. An idealized reticulated particle of carbon black is illustrated in Fig. 1. While not intending to be bound by any one theory, it is believ that the primary particles may be fused together or share commo microcrystallite bundles or planes to make up the primary chain This primary reticulated structure is to be distinguished from secondary or reversible structures which result from van der Wa forces between individual reticulated particles. Primary carbo black structures exhibit a pronounced tendency to agglomerate in secondary reticulated structures when dispersed in almost all me The carbon black of the present invention is preferably produced in refractory-lined furnace reactors by pyrolysis of highly aro¬ matic refinery by-product oils. These oils are subjected to tem eratures of about 1,400 to 1,650°C in a reaction zone maintained at conditions producing an endothermic reaction which strips ato hydrogen from the aromatic hydrocarbon molecules to leave aromat carbon nuclei. The resulting reticulated particles in the form a black "smoke" are quenched in a downstream tunnel by water inj tion at a point several feet from the reaction zone. In this method of manufacturing carbon black, the primary particle size can be closely controlled and produces particle diameters in the range of about 200 to 900 A°. These primary particles are simul taneously bound together to form primary reticulated chains havi lengths in the range of about 500 to 30,000 A°. Both structure particle zie may be closely controlled through the design of the oil injection nozzle, reaction chamber geometry, pyrolysis temp¬ erature, residence time, and the intensity of gaseous turbulence Another important feature of carbon black is that it contains - 23 - less than about half the amount of hydrogen theoretically need to bond all edge portions of the polynuclear aromatic platelet and it is believed that the particles contain many unsatisfied valences or free radicals. A number of elements, such as oxy¬ gen and sulfur, may interact at some of the peripheral posi¬ tions of these micrographic platelets so as to form complexes which are generally analogous to the functional groups of or¬ ganic compounds. The principal surface groups on carbon black have been identified in the literature as carboxylic acid, phe nolic hydroxyl and quinone groups, and possibly peroxide and lactone groups. In addition to accepting or trapping free radicals, it has been suggested that functional groups at the carbon black surface can also generate free radicals, or at least can initiate free radical reactions. It is believed that these organic functional groups may assist in forming both the unfired and fired catalyst compositions of the pres¬ ent invention. It should also be noted that carbon black is an amorphous form of carbon as opposed to graphite which is a soft crystalline form of carbon that differs greatly in properties from amor¬ phous carbon. The thermal conductivity of amorphous carbon is relatively high and is equivalent to some metals. Carbon blac also has a very low co-efficient of thermal expansion and a high resistance to thermal shock. It is believed that these features contribute to relatively rapid rates of carbon burn¬ out and large feeder pores of substantially uniform diameter. A preferred type of carbon black meets the specifications of ASTM No. N-219. These blacks have relatively low structure and are made using an intermediate super-abrasion furnace. Such blacks are available from Ashland Chemical (United N-219) , Cabot (Regal 600) , Columbian (Neotex 130) , Continental (Conti- nex ISAF-LS) , and Phillips (Philblack N-219) . United N-219L is preferred as it is not compacted but supplied loose at rela tively low bulk density compared to pelleted or compacted blacks. - - The average primary particle diameter of this black is about 300A and it has an ASTM Iodine Number of about 115 (a measurement of surface area per unit weight correlating well with nitroge absorption measurements for furnace blacks) . The relatively low structure of this black is indicated by a low DBP Absorp¬ tion value of about 0.78 cubic centimeters per gram (the DBP Absorption value is indicative of the degree of linkage be¬ tween primary carbon black particles) . One of the principal objects of the invention is to disperse the carbon black sufficiently in the catalyst forming media so that formation of secondary CB structures through agglomer tion of primary CB structures can be significantly controlled. Some of the more effective dispersants for carbon black in aqueous media are hexadecyltrimethylammonium bromide, an ethoxylated alcohol sulfate sold under the brand name Triton X-100, a sodium lignosulfonate sold under the brand name Marasperse CBO-3, and mixtures thereof. These dispersants ar used in amounts generally ' proportional to the weight of carbo black added, preferred proportions being in the range of abou 0.05 to 1.0 weight percent of carbon black, more preferably about 0.1 percent. Quaternary surfactants such as Quaternary 0, succinates such as Aerosol, and other ethoxylated alcohol sulfates may also be used. OMPI - 25 - With carbon black, it has been found that the most predominant feeder pore size can be controlled to a significant extent bot by the amount of carbon black used and the amount and effec¬ tiveness of the dispersant used to suspend the carbon black in an aqueous medium. It is believed that the degree of disper¬ sion versus the degree of agglomeration of the primary reticu¬ lated CB particles is a controlling factor in determining whether the predominant feeder pore sizes are in the lower or the upper portion of the preferred pore size range of 500 to 6,000A°. Thus, lower carbon black concentrations in combina¬ tion with the most effective, dispersants provide an increased number of pores with effectivediameters in the range of 400 to l,000A°when using primary carbon black particles having an aver age diameter of about 300A° . Larger concentrations of carbon black and lower and/or less effective concentrations of carbon black dispersants provide predominant feeder pore sizes in the range above 1,000A°. This phenomenon is believed to be due to the formation of feeder pores by agglomerates of two or more individual reticulated CB particles. The amount of carbon black used in preparation of the unfired catalyst will therefore depend on the extent to which large feeder pores are desired in the final catalyst structure. Other factors include the final attrition resistance desired. Generally, the amount of carbon black should be in the range of 1 to 35 percent by weight of unfired product. Too little carbon black will not produce a sufficient number of large pores and too- much carbon black will result in a catalyst having relatively low attrition resistance. Accordingly, pre¬ ferred amounts of carbon black are in the range of about 2 to 30 percent by weight, more preferably 5 to 15 percent by weigh of the unfired product. Another factor to be considered in selecting the amount and ty of carbon bla ' ck is the average effective diameter of the hydrocarbon molecules in the feedstock. The average diameter referred to here is the statistical average of the largest ef- - 26 - fective dimension of the molecules boiling above 1025°F. The feeder pores should have an average effective diameter at least equal to this average feedstock dimension, but should not exceed about 10 times this dimension so as not to decrease unduly the surface area and, correspondingly, the number of catalytic sites available for the cracking reaction. Accordingly, the ratio of average feeder pore diameter to average hydrocarbon diameter should be in the range of 2 to10 more preferably 4 to 8, most preferably 5 to 8. Another important feature of the invention is that the carbon black suspension, preferably an aqueous medium, can be mixed uniformly with a catalyst slurry and the resulting composite suspension spray dried to form substantially uniform micro- spheres within the preferred range of particle sizes described below. By comparison, prior art techniques using a decompos¬ able solid for introducing large pores into a matrix involve the formation and extrusion of a viscous paste which then has to be dried and broken up. This results in catalyst particles having a wide range of diverse shapes and sizes which have to be sifted in order to provide a catalyst of any uniformity. In addition, the types of . decomposable solids used in the prio art produce low activity catalysts with excessively large pore the pore size range in any given mix being virtually uncontrol lable. After being intimately mixed with the matrix material and any other ingredients, such as zeolite and/or filler components, the composite is shaped and dried to produce an unfired cata¬ lyst composite. This shaped composite is then heated to burn out the carbon black and produce a final catalyst product con¬ taining a significant volume of large feeder pores within the desired size ranges. The temperature experienced by the cata¬ lyst particles should not cause objectionable changes in the structure of either the zeolite or the supporting matrix. Wher carbon black is removed during manufacture, burn out is initiat at about 500°F and the firing time varies in accordance with th temperature selected, higher temperatures requiring shorter firing times. Where carbon black is burned out in the process unit during regeneration, the temperature should not exceed ^ - 27 - about 1,500°F to avoid damage to the zeolite. Preferred firing temperatures are in the range of about 1,000 to 1,450°F, with corresponding firing times from about three hours to as low as a few minutes, such as associated with catalyst hold-up in a regenerator. Where carbon black is burned out of the compositio prior to in situ formation of the zeolite, temperatures as high as 2,000°F may be tolerated. The invention is not restricted to the use of contact agents containing any specific matrix components or catalytic promoter Any matrices and/or promoters of the prior art may be used in combination with carbon black for. the production of feeder pore in solid catalysts which may be of any suitable shape an^ «* \-2. For example, carbon black may be used to provide feeder pores in synthetic silica-alumina catalysts of the type described in U.S. Patent No. 3,034,994 to Braithwaite, et al., which patent is incorporated herein by reference. However, it is preferred to use this or a similar s.ilica-alumina composition as a matrix for supporting a superactive zeolite component. It is to be further understood that other promoters, such as catalytically active metals or metal compounds, may be used in place of or along with a zeolitic promoter. A preferred catalyst of the pr ent invention therefore comprises three main components, namely catalytically active or inactive matrix material, a superactive catalytic promoter dispersed in the matrix material, and a car¬ bon black initially dispersed in the matrix but removable there from by combustion. Virtually any refractory oxide material capable of maintaining stable pore characteristics may be used as the matrix for the catalyst of this invention. A preferred matrix composition is one having sufficient acid sites to provide significant crack¬ ing activity, particularly for the high molecular weight com¬ ponents of the feed. It is therefore a further object of the invention to employ a catalyst matrix in which significant catalysis of the heavy hydrocarbon molecules boiling above 1,025°F is affected in the feeder pores. Catalysis in these - 28 - macropores may be affected by acidic sites either in the ma¬ trix itself or on exposed outer surfaces of the superactive zeolite. It is believed that conversion and selectivity is significantly improved if these feed components are initially cracked in the matrix to smaller size molecules capable of en tering the much smaller zeolite pores. Less reliance on ther mal cracking for these types of reactions gives improved pro¬ duct distribution and gasoline yield and better overall produ quality, i.e., less methane, ethane, ethylene, thermal coke, thermal gas oil, and thermal gasoline. A catalyst microsphere made according to the invention is il¬ lustrated in Fig. 2. The mechanism for trapping heavy liquid asphaltenes in the matrix feeder pores so as to reduce the nu ber reaching and blocking zeolitic pores is illustrated in Fig 3. Fig. 4 shows diagramatically the cracking of a heavy poly¬ nuclear aromatic at an acidic site (H+) in the matrix. The straight chain fragment from the matrix cracking of Fig. 4 mayth be cracked and reformed by the zeolite as illustrated in Fig. 5 The matrix material should possess good hydro-thermal stabilit Examples of materials exhibiting relatively stable pore charac teristics are alumina, silica-alumina, silica, silica-magnesia magnesia-alumina, silica-zirconia, ' clays such as kaolin, meta- kaolin, halloysite, anauxite ' , dickite and/or macrite, and comb nations of these materials. Other clays, such as natural mont morillonite, synthetic mica montmorillonite (SMM) , and/or pil¬ lared layered clays (PLC) may be added to increase the acidity of the matrix. Clay may be used in its natural or thermally modified states. The preferred matrix of U.S. Patent No. 3,034,994 is a se isynthetic combination of clay and silica-alumina. Preferably the clay is mostly a kaolinite and is combined with a synthetic silica-alumina hydrogel or hydrosol. This synthetic component forms preferably about 15 to 75 percent, more preferably about 20 to 25 per¬ cent, of the fired catalyst in weight. The proportion of clay is such that the catalyst preferably contains after firing abo 10 to 75 percent, more preferably about 30 to 50 percent , clay by weight. A most preferred comoosition of the matrix coni - 29 - approximately twice as much clay as synthetically derived silic alumina or silica-alumina. Synthetically derived silica-alumin should contain 55 to 95 percent by weight of silica (SiO-) , mor preferably about 75 percent. After introduction of the zeolite and/or other promoters, the composition is preferably slurried and spray dried to form catalyst microspheres. The particle size of the spray dried matrix is generally in the range of about 5 to 160 microns, preferably 40 to 80 microns. Various processes may be used in preparing the synthetic silica alumina, such as those described in U.S. Patent No. 3,034,994. One of these processes involves gelling an alkali metal silicat with an inorganic acid while maintaining the pH on the alkaline side. An aqueous solution of an acidic aluminum salt is then intimately mixed with the silica hydrogel so that the aluminum salt solution fills the silica hydrogel pores. The aluminum is thereafter precipitated as a hydrous alumina by the addition of an alkaline compound. As a specific example of this method of preparation, a silica hydrogel is prepared by adding sulfuric acid with vigorous ag¬ gitation and controlled temperature and concentration condition to a sodium silicate solution. Aluminum sulfate in water is th added to the silica hydrogel with vigorous aggitation to fill t gel pores with the aluminum salt solution. An ammonium solutio is then added to the gel with vigorous aggitation to precipitat the aluminum as hydrous alumina which combines with the silica to produce silica-alumina on the surface of the silica hydrogel after which the hydrous gel is processed, for instance, by separating a part of the water on vacuum filters and then dryin or more preferably, by spray drying the hydrous gel to produce microspheres. The dried product is then washed to remove sodiu and sulfate ions τ either with water or a very weak acid solutio The resulting product is then dried to a low moisture content, usually less tjian 25 percent by weight, e.g., 10 percent to 20 percent by weight, to provide the finished catalyst product. The silica-alumina hydrogel slurry may be filtered and washed - 3.0 - in gel form to affect purification of the gel by the removal of di solved salts . This may enhance the formation of a continuous phas in the spray dried microspheric particles . If the slurry is prefil tered and washed and it is- desired to spray dry the filter cake , t latter may be res lurried with enough, water to produce a pumpable mi ture for spray drying . The spray dried product may then be washed a gain and given a final drying in the manner previously described. Spray driable compositions to which carbon black can be added and spray drying techniques usable with the present invention are also described in U.S. Patent No. 3,867,308 to Elliott and U.S. Patent No. 4,126,579 to Flaherty, et al. , which patents are incorporated herein by reference. Suitable catalytically active promoters other than zeolites may be used and include metals or a catalytic compound of metals such as Pb, Sn, Bi, Ge Sc, Ti, Cr, Mn, Co, Zn, Y, Nb, Mo, Ma, Ru, Rh, Pd, La, Hf, Ta, W, Re, Os, Ir, Pt, Zr, Ac, Th, Pa, and U and the. like. Additional contact agents for the matrix in¬ clude metal oxides and sulfides, selenide, teluride, halides, phosphates, manga- nates, molybdates, chromates, bi-chromates and the like. These components may be used alone or in addition to a superactive zeolite. In the latter case, these elements and/or compounds may increase the catalytically active sites available in the matrix. These additional promoters may be used in concentration ranges generally from about 0.5 percent to about 20 percent, more preferably about 1 to 5 percent by weight of fired catalyst. The catalytically active promoter for a preferred catalyst compositio is a - crystalline aluminosilicate zeolite , commonly known as molecular sieves . Molecular sieves are initially formed as alkalimetal alumino silicates , which are dehydrated forms of crystalline hydrous siliceous zeolites . However ,, since the alkali form does not have appreciable ac tivity and alkali metal ions are deleterious to cracking processes , the aluminosilicates are ion exchanged to replace sodium with some other ion such as , for example , ammonium and/or rare earth metal ion The silica and alumina making up the structure of the zeolite are ar ranged in a definite crystalline pattern ' containing a large number o small uniform cavities interconnected by smaller uniform channels o pores . The effective size of these pores is usually between about 4A° and 13A°. The zeolites which can be employed in accordance with this in¬ vention include both natural and synthetic zeolites. The nat¬ ural occurring zeolites include gmelinite, clinoptilolite, chabazite, dechiardite, faujasite, heulandite, erionite, anal- cite, levynite, sodalite, cancrinite, nepheline, lcyurite, scolicite, natrolite, offretite, mesolite, mordenite, brew- sterite, ferrierite, and the like. Suitable synthetic zeolites include zeolites Y, A, L, ZK-4B, B, E, F, H, J, M, Q, T, W, X, Z, ZSM-types, alpha, beta and omega. The term "zeolites" as used herein contemplates not only aluminosilicates but sub¬ stances in which the aluminum is replaced by gallium and sub¬ stances in which the silicon is replaced by germanium. The zeolite materials utilized in the preferred embodiments of this invention are synthetic faujasites which possess silica to alumina ratios in the range from about 2.5 to 7.0, prefera¬ bly 3.0 to 6.0 and most preferably 4.5 to 6.0. Synthetic faujasites are widely known as crystalline aluminosilicate zeo¬ lites and common examples of synthetic faujasites are the X and Y types commercially available from the Davisσn Division of W. R. Grace and Company and the Linde Division of Union Carbide Corporation. The ultrastable hydrogen exchanged zeolites, such as Z-14XS and Z-14US from Davison, are particularly suitable. In addition to faujasites, other preferred types of zeolitic materials are mordenite and erionite. The preferred synthetic faujasite is zeolite Y which may be pre pared as described in U.S. Patent No. 3,130,007 and U.S. Pat¬ ent No. 4,010,116, which patents are incorporated hereby by ref erence. The aluminosilicates of this latter patent have high silica (Si0 2 ) to alumina (Al 2 O. molar ratios,- preferably above 4, to give high thermal stability. The following is an example of a zeolite produced by the silica tion of clay. A reaction composition is produced from a mix¬ ture of sodium silicate, sodium hydroxide, and sodium chloride formulated to contain 5.27 mole percent Si0 2 , 3.5 mole percent Na 2 0, 1.7 mole ' ercent chlorine and the balance water. 12.5 parts of this solution are mixed with 1 part by weight of cal- - 32 - cined kaolin clay. The reaction mixture is held at about 60°F to 75°F for aperiod of about four days. After this low temper ture digestion step, the mixture is heated with live steam to about 190°F until crystallization of the material is complete, for example, about 72 hours. The crystalline material is fil¬ tered and washed to give a silicated clay zeolite having a sil ca to alumina ratio of about 4.3 and containing about 13.5 per cent by weight of Na 2 0 on a volatile free basis. Variation of the components and of the times and temperatures, as is usual in commercial operations, will produce zeolites having silica to alumina mole ratios varying from about 4 to about 5. Mole ratios above 5 may be obtained by increasing the amount of Si0 2 in the reaction mixture. The sodium form of the zeolite is the exchanged with polyvalant cations to reduce the Na 2 0 content to less than about 5 percent by weight, and preferably less than 1.0 percent by weight. Procedures for removing alkali metals and putting the zeolite in the proper form are well-known in the art as described in U.S. Patent Nos. 3,293,192; 3,402,996; 3,446,727; 3,449,070; and 3,537,816; which patents are incor¬ porated herein by reference. As previously indicated, the carbon black ingredient of the present invention may be admixed with the zeolite forming com¬ position prior to crystallization of the zeolite and the carbon black component later burned out to form large feeder pores between the zeolite crystals in the resulting agglomerate. The zeolites and/or other promoters can be suitably dispersed in matrix materials for use as .cracking catalysts by methods well-known in the art, such as those disclosed, for example, in U.S. Patent Nos. 3,140,249 and 3,140,253 to Plank, et al.; U.S. Patent No. 3,660,274 to Blazek, et al.; U.S. Patent No. 4,010,116 to Secor, et al.; U.S. Patent No. 3,944,482 to Mit¬ chell, et al.; ' and U.S. Patent No. 4,079,019 to Scherzer, et al.; which patents are incorporated herein by reference. ■ ro RE OMPI - 33 - The amount of zeolitic material dispersed in the matrix based on the final fired product should be at least about 10 weight percent, preferably in the range of about 25 to 40 weight per¬ cent, most preferably about 35 to 40 weight percent. More tha one type of zeolitic particles, such as zeolites with differen functions and/or selectivities, and zeolitic particles in com¬ bination with particles of metals or other catalytic materials may be employed together to make up these amounts of total pro moter ingredient. For example, particles of hydrdesulfurizing catalyt as in U.S. Patent No. 3,770,617 may be mixed with part cles of a cracking catalyst as in 4,010,116. The upper ranges of zeolite concentrations, together with amuch less but signi¬ ficantly active matrix, are preferred to provide a catalyst wi ultrahigh cracking activity even at large pore volumes and rel tively low surface areas. Crystalline aluminosilicate zeolites exhibit acidic sites on both interior and exterior surfaces with the largest proportio of total surface area and cracking sites being internal to the particles within the crystalline micropores. These zeolites are usually crystallized as regularly shaped, discreet particl of approximately 0.1 to 10 microns in size and, accordingly, this is the size range normally provided by commercial catalys suppliers. To increase exterior (portal) surface area, the particle size of the zeolites for the present invention is pre erably in the range of less than 0.1 to about 1 micron and mor preferably in the range of less than 0.1 micron. The preferre zeolites are thermally stabilized by heating to temperatures o at least 1,000°F and then further exchanged with hydrogen ions and/or rare earth ions. These zeolites are steam stable to about 1,650°F. Feeder pore content of the catalyst within the size range of about 400 to 6,000A° is provided in the main by the preparativ method of the invention. Macropores of less than 400A° may be produced by this method and may also be introduced into the catalyst matrix by prior art methods for preparing such matric The micropore * content of the catalyst in the size range below 20A° is provided principally by the zeolite particles per se. - - Since the amount of zeolite used may vary widely, the micropor content of the catalyst will be variable by a similar amount. The fraction of total pore volume provided by micropores of 20 or less also will vary depending upon the fraction of large feeder pores introduced by the carbon black component. The surface area of the zeolitic component in the catalyst may be estimated by multiplying the surface area of the pure zeoli (usually about 800 to 900 m 2/gm) by the percentage of zeolite the final catalyst. Total pore volume including the feeder po volume of the matrix should be at least 0.2 cc/gm, preferably more than 0.4 cc/gm. The upper limit for total matrix pore volume is best expressed as the amount of contact surface area provided in the final catalyst matrix. In general, the minimu surface area of a satisfactory catalyst matrix is about 20 squ meters per gram, preferably at least 30 square meters per gram, more preferably at least 40 square meters per gram. These par meters are believed to provide sufficient volume in the form o large feeder pores to reasonably minimize diffusion limitation effects, yet not so large a volume as to unduly reduce the availability of acidic sites or adversely affect the physical properties, e.g., attrition resistance, of the catalyst. The introduction of CB feeder pores, as a practical matter, causes no appreciable change in the volume and size distributio of micropores in the zeolite and macropores produced in the matrix by other mechanisms. Thus, the macropore content of the catalyst particles produced by the carbon black of the inventio is readily obtained by the difference in pore volume between th unfired and fired catalyst composition, provided there is no significant sintering of the matrix material. The determinatio of feeder pore volume is by the conventional mercury porosimetr method referenced above. In order to test the activity of a catalyst, it is the practice in the petroleum cracking art to measure the catalyst activity by a bench-scale test. Although various tests have been pro- OMPI - 35 - posed, one test widely accepted by the industry at the present time is known as the microactivity test or "MAT". The micro- activity test and standard procedures for obtaining "MAT Activi are described below. Because the MAT test range is appropriate for equilibrium (used) zeolite catalyst instead of virgin (unus catalyst, virgin catalyst of thepresent invention is treated with 97-100% steam at 1,450°F for 5 hours to provide a standard reduction in its activity before it is tested by the MAT proce¬ dure. Having thus described the catalyst of the invention broadly, th following examples are offered to illustrate methods for its preparation in more detail. Example 1 2.1 gms of carbon black pellets (N-339 from Ashland Chemical Co pany) were placed in 350 ml of deionized water containing 0.03 of hexadecyltrimethylammonium bromide and 0.03 gm of Triton X-100 dispersants. This mixture was agitated ultrasonically with an Artex Sonic probe at 300 watts for 30 minutes, and the resulting carbon black suspension added to 700 ml of water along with 22.2 gm of Kaopaque 10 Kaolinte. This slurry was mixed for 30 minutes with a Premier High Viscosity Dispersa- tor and then 23.4 gm of rare earth zeolite was added, along with 5 ml of PQ N-Brand sodium metasilicate as a dispersant for the clay and zeolite. The resulting slurry was mixed for an additional 30 minutes with the Dispersator and then 120 ml of sodium metasilicate and 200 ml of water were added with mix¬ ing for 10 more minutes. 180 ml of a 11.5 wt% sulfuric acid so tion were added slowly while mixing continued to partially neu¬ tralize the slurry and precipitate silica gel. This silica ge contained uniformly dispersed carbon black and clay and was aged for 1 hour at 110°F to increase the size of interstitial cavities within the gel. Mixing was then resumed and a 18 wt% solution of aluminum sulfate, prepared from 108.4 gm A1 2 (S0.).-,: 18H 2 0 and 200 ml of water at 120°F, was added to the silica OMPI - 36 - gel and mixing continued for 15 minutes at 110°F. This slurry was then neutralized by adding 68 ml of concentrated NH.OH to precipitate alumina gel. The silica and alumina gel mixture was filtered and the resulting gel cake washed three times wit 3L of water at 150°F. The washed gel cake was then dried at 500°F for 16 hours and the dried cake ground into small solid particles. These particles were then exchanged at 200°F for 1 hour with IL of 1.0M rare earth chloride solution. The ex¬ changed particles were filtered and washed three times with 2L of water at 150°F and dried at 500°F for 16 hours. After dry¬ ing, the carbon black was burned out by firing the catalyst particles in air at 1,000°F until the catalyst became white in color (about 3 hours) . This procedure produced about 100 ~~ of dry catalyst particles consisting essentially of 20 wt% zeo¬ lite and 20 wt% clay substantially uniformly dispersed in 60 wt% of a silica-alumina matrix containing about 75 wt% Si0 2 - Examples 2-5 The quantity of carbon black used in Example 1 above provided about 2 wt% carbon black in the unfired catalyst composite. In Examples 2-5, the same preparation procedure was followed but the amount of carbon black was changed to 4.3, 11.6, 27.1 and 44.6 gms giving a carbon black wt% in the unfired composites of 4, 10, 20, and 30, respectively. In these examples, the amounts of the other ingredients remained the same except that the amount of carbon black dispersant was varied in proportion to the amount of black used. Thus, the weights of hexadecyl- trimethlyammonium bromide and Triton X-100 used in these ex¬ amples were each approximately 1% of the weight of carbon black added to the deionized water. The changes in pore volume produced by the carbon black of Ex¬ amples 1 through 5 are shown in Table 1 relative to a fired control sample containing no carbon black. In order to show the effects of firing alone on pore size distribution, the con- trol sample without carbon black was also fired and the result¬ ing pore volume changes for this sample are reflected in the - 37 - first column of Table 1 under the heading "Control" . The data shown in Table 1 indicates that firing of the control sample containing no carbon black caused a slight decrease in pore volume in all pore size ranges. By comparison, the firing of samples containing carbon black resulting in significant increases of porosity in the ranges above both 400 A° and 1,000 Several composites also showed significant increases in porosity near 60 A° for reasons not yet explained, but which could involv differences in aging the hydrogel. Sample No. 5, which containe 30 wt% carbon black, did not form an adhering cake when filtered and dried, but instead produced a fine powder as an unfired prod TABLE 1 - POROSITY EFFECTS OF CARBON BLACK ADDITIONS Sample No. - Control 1 2 3 4 5 Carbon Black, wt% - 0 2 4 10 20 30 Pore Volume Changes in cc/gm as Compared to fired control Pore Diameter Range In A 0 sample without carbon black. 30-35 -0.012 0.017 -0.017 0.040 -0.013 0.014 35-40 -0.009 0.006 0.045 0.015 0.007 0.008 40-60 -0.032 0.184 0.139 0.040 0.135 0.024 60-80 -0.007 0.044 0.005 0.006 0.026 0.005 80-100 -0.010 0.007 0.013 0.009 0.008 0.009 100-200 -0.015 0.004 0.036 0.034 0.029 0.026 200-400 -0.002 0.004 0.012 0.078 -0.010 -0.026 400-1000 -0.010 -0.009 0.027 0.026 0.073 0.137 1000-6000 -0.002 0.020 0.015 0.037 0.154 0.222 Pore Volume Change For 100-6000A 0 , cc/g -0.029 0 .019 0. 090 0.175 0.256 0.359 Total Pore Volume - 39 - Examples 6-9 The catalyst samples of Examples 6 through 9 were prepared fol¬ lowing the same basic procedure of Examples 1 through 5 except the amount of clay was increased and the amount of silica and alumina gel decreased so that the amounts of clay and of silic alumina gel were approximately equal in the final fired produc with the resulting silica-alumina matrix containing about 75 w Si0 2 . In addition, instead of filtering and drying to produce a gel cake, the final neutralized slurry was spray dried to for microspheres of catalyst. These microspheres were then washed, exchanged with rare earth chloride solution, rewashed, and drie at 300°F for 16 hours in a manner similar to the gel cake and ground particles cf Ξxcmples 1-5. Spray drying ro uce" * smalle and more uniform particle sizes than grinding gel cake and is preferred for providing a more uniform and fluidizable catalyst The average size of the spray dried particles is preferably within the range of about 40 to 80 microns. The changes- in pore volume produced by the carbon black of Ex¬ amples 6 through 9 are given in Table 2. In this comparison, the changes in pore volumeweremeasured relative to a fired con trol sample containing no carbon black instead of the unfired type of control sample forming the basis for the comparison in Table 1. The changes in porosity exhibited by the samples of Table 2 are consistent with those of " Table 1 in that the use of carbon black increased catalyst porosity in the ranges above 400A° and 1,000A°. However, no significant increase in porosi occurred in the 40 to 80A° range, the reason for this variance with the data of Table 1 being uncertain at the present time. TABLE 2 - POROSITY EFFECTS OF CARBON BLACK ADDITIONS Sample No. - Carbon Black, wt% - 11.6 22.0 23.4 27.0 Pore Volume Change in cc/gm as compare to fired control sample without carbon black. Pore Diameter. " . Range in A 0 . Group A 30-35 + .004 0.022 0.017 +0.006 35-40 -0.001 0.004 0.001 0.001 40-60 - .007 -0.051 -0.042 -0.068 60-80 - .003 -0.035 -0.026 0.042 80-100 0.006 -0.009 -0.013 -0.011 100-200 0.020 -0.005 -0.003 -0.010 200-400 0.027 0.009 0.008 0.004 400-1000 0.058 0.044 0.022 0.014 1000-6000 0.109 0.233 0.227 0.225 5*6000 * 0.233 -0.032 0.481 0.225 Group B - 1000->6000 * 0.342 0.201 0.708 0.450 400->6000 * 0.400 0.245 0.730 0.464 200->6000 * 0.427 0.254 0.738 0.468 100->6000 * 0.447 0.249 0.735 0.458 Group C 400 - 6000 0.167 0.277 0.249 0.239 200 - 6000 0.194 0.286 0.257 0.249 100 - 6000 0.214 0.281 0.251 0.233 Grouo D 200 - 1000 0.085 0.053 0.030 0.018 100 - 1000 • 0.105 0.048 0.027 0.008 * Includes pore sizes above 6000 A° . ■ OMPI - 41 - In Fig. 6 of the drawings, the changes in pore volume are plotted against carbon black content for several of the pore diameter ranges in Table 2. The curves of Fig. 6 correspond to the data of Table 2 as follows: curve A-7 represents the pore volume change in the 7th pore diameter range of group A; curve A-8 represents the pore volume change in the 8th pore diameter range of group A; curve D-2 represents the pore vol¬ ume change in the second pore diameter range of group D; and curve C-3 represents the pore volume change in the third pore diameter range of group C. The curves of Fig. 6 suggest that the maximum number of feeder pores having diameters in th<? r of about 100 to 6,000 ° occur with 18 to 19 wt% carbon black in the unfired catalyst composite. This amount of carbon black, which equals about 50% by weight of the dry silica-alumina gel in the unfired com¬ posite, increases the volume of pores having diameters greater than 100A° by about 0.3 cc per gm. Similarly, a carbon black concentration of about 12 wt% in the unfired composite, which is about 30% by weight of the silica-alumina gel present, pro¬ vides the maximum number of feeder pores in the 100 to 1,000A° . range. This corresponds to an increase in pore volume of abou 0.1 cc per gm. As also evident from Table 2, most of the poros ity increase at 12 wt% carbon black is due to feeder pores in the 400 to 1,000A° range. Those feeder pores increased the pore volume by about 0.07 cc/gm, which is about the volume of the carbon black burned out of the sample. _The maximum porosi increase in the 200 to 400A° pore diameter range occurred at about 10 wt% carbon black. The curves of Fig. 6 further indicate that there is an optimum carbon black concentration to attain maximum porosity in a given pore size range and that lower concentrations are requir to give the maximum porosity for smaller diameter pores. This suggests that the pore sizing attainable with carbon black is closely related to the degree to which individual reticulated carbon black oarticles are separated from each other and dis- - 42 - persed within the silica-alumina matrix, full dispersion be¬ coming more difficult to attain and maintain as concentration of carbon black in the gel slurry increase. In addition to using lower concentrations of carbon black, more effective ca bon black dispersants also increase the proportion of pore si in the smaller diameter ranges of feeder pores. One example a more effective dispersant for carbon black is sodium ligno- sulfonate which may be used in place of hexadecyltrimethlam- monium bromide and/or Triton X-100 dispersants, particularly where more feeder pore diameters approximating the diameter of primary carbon black particles are desired. Lignin is derived from wood pulp and varies in molecular weig between 1,000 and 50,000. It's basic organic structural unit is a substituted phenylpropane. Lignin dispersants are com¬ mercially available under the name Marasperse which is a sodi lignosulfonate of relatively low sulfonation. While not in¬ tending to be bound by any one theory or hypothesis, the acti of lignin as a dispersant is believed to be .electrochemical i nature. When lignin molecules are adsorbed on the solid carb black or clay particles in aqueous suspension, they impart a negative charge to the particles, causing them to repel each other. Adsorption of lignin molecules on a particle may also create a film which then acts as a physical barrier against direct contact between the particle and the surrounding aque¬ ous media, including silica colloids. These effects are be¬ lieved to contribute to the production of feeder pore sizes corresponding generally to the transverse diameters of pri¬ mary reticulated carbon black particles and to improved pore size control. Sodium lignosulfonate is a preferred dispersan for carbon black in the catalyst composition of the invention and is added to the matrix slurry with the carbon black prior to the spray drying or other forming step for shaping the sol catalyst particles. The catalyst compositions containing lig¬ nin and the processes for making those compositions are at¬ tributed to William P. Hettinger, Jr., James E. Lewis and H. Wayne Beck, all of Ashland Oil, inc. - 43 - Examples 10-11 In a Kady Mill homogenizer, 12 liters of tap water was mixed with 960 grams of aluminum sulfate hydrate and 1.2 liters of H 2 SO . Using an addition rate of 200 ml/min, 9 liters of "N" Brand sodium metasilicate was added to this acid solution whil mixing to provide a silica hydrosol. Ten grams of a lignin dispersant called Marasperse and 800 grams of N-219 carbon black were added to the homogenizer and mixed for 5 minutes to disperse carbon black in the silica hydrosol. In a separate container, 10 grams of sodium pyrophosphate (Na.P 2 0_) was adde to 11 liters of 2pH water made with H_S0.. With vigorous mix¬ ing, 11 kg of fine kaolinite clay was added to form a clay slur ry. The clay slurry was then combined with the silica hydro¬ sol in the homogenizer and mixed for 5 minutes. A slurry of NaY zeolite made from 4 liters of 2pH H 2 0 and 4 kg of zeolite was quickly added to the homogenizer and mixed for 10 minutes. The resulting slurry was immediately spray dried at 400°C inlet and 125°C outlet temperatures in a Niro Atomize Model V Spray Drier to form catalyst microspheres. Air pres¬ sure was 30 psig. One kilogram of the microspheres from the spray drier was washed three times with 4 liters of 150°F wate and filtered. The filter cake was exchanged twice with 3 lite of 1.25 Molar NH 3 C1 for 15 minutes each at 150°F. After fil¬ tering again, the cake was exchanged 3 times at 150°F for 30 minutes each with 3 liters of 0.33N mixed rare earth chloride solution. The solid particles were then finally washed 4 time with 3 liters of 150°F water and dried at 300°F for 16 hours to produce Sample No. 10. A portion of this sample was oxi¬ dized in air at 1,000°F in a shallow bed for 4 hours to burn out the carbon black and produce Sample No. 10R. Samples 11 and 11R were prepared by the same procedure except twice as much carbon black was used, the designation "R" mean¬ ing regenerated (fired) to burn out the carbon black particles - 44 - In the preferred embodiment of the invention, the fired micro spheres are further. exchanged after firing with a 0.5N soluti of mixed rare earth chlorides at 150°F for 30 minutes, and then washed 4 times with 150°F water and dried for 16 hours at 300°F. The firing step to remove carbon black also cal¬ cines the zeolite in the composition. This calcination in combination with the further ion exchange provides an ultra- stable, hydrothermally resistant zeolite catalyst. A portion of Sample 11R was further exchanged in this manner to produce Sample 11RE shown in Table 5. ■ ^3RE QMPI - - - 45 - Table 3 gives an analysis of the compositions of Samples 10R an 11R as compared to a commercial Super DX catalyst available fro the Davison Division of W. R. Grace and Company. The Davison catalyst has a comparable total pore volume and surface area relative to the sample catalysts. Table 4 gives porosity mea¬ surements for Samples 10, 10R, 11, and 11R over various pore diameter ranges. This table also shows the change in pore vol¬ ume, both in cc/gm and as a percentage change, upon burning out the carbon black from Samples 10 and 11 to produce Samples 10R and 11R. Table 5 gives some of the catalytic properties of Samples 10R, 11R and llRE as compared to the Davison catalyst. The data given was determined by a micro-activity test (MAT) based on the procedure found in ASTM test method No. D-3908-80 which is described in more detail below. Although an FCC type feed is used in this standard test so that it is not fully indica¬ tive of the performance of the catalyst samples with a carbo- metallic feed, the data is believed to show the advantage of the invention in reducing the carbon make. Thus, the carbon producing factor (CPF as defined below) and the weight per¬ cent coke (relative to feed weight) are significantly lower for Samples 10R, 11R and 11RE of the invention relative to the Davison catalyst as seen in Table 5.. It is also significant that the MAT conversion obtained with Sample 10R was 80 as com¬ pared to 73 obtained with the Davison catalyst. Referring to Table 3, it will be seen that the higher conversion was ob¬ tained with less zeolite, namely, 6.6 percent intensity for Sample 10R relative to 13.1 percent intensity for Super DX. The relatively high conversion and low coke producing factors o Sample 10R demonstrate clearly that the heavier hydrocarbons of even a MAT feed can be cracked by and stripped more effec¬ tively from the catalyst of the present invention. The lower conversion experienced with Samples 11R and USE are attributed to the very low level of zeolite measured for those samples, namely, 3.8 percent intensity. - 46 - Table 3 - Composition of Samples 10R and 11R Compared to Davison Super DX Catalyst Sample 10R Sample 11R Davison Super Amount of Carbon Black Burned Out, wgt% (1) 4.5 ■ 8.3 None A1 2 0_, wgt% 29.6 27.0 28.3 Si0 3 , wgt% 63.2 55.3 64.2 Na D, wgt% 0.33 0.22 0.84 Zeolite Intensity, % NaY (2) 6.6 3.8 13.1 Hg (<6000 A 0 ) Pore ϋiϋTι" * 2, cc/gm 0.25 0.33 .0.22 2 Surface Area, m /gm 198 197 197 (1) Carbon Black wgt% based on unfired Samples 10 and 11. All other wgt% are o ignition basis (IB) of final fired catalyst. (2) Zeolite X-ray intensity of sample as percentage of X-ray intensity of high purity sodium Y zeolite. _QMPI _ Table 4 - Porosity Measurements • of Sampli 2S 10 , 10R, 11 and 11R Pore Dia. S-10 S-10R Change, S-ll S-llR Change, Range (A°) cc/gm (1) cc/gm (2) cc/gm % cge (3) cc/gm (1) cc/gm (2) cc/gm % cge (3 6000-1000 0.039 0.049 0.010 25.6 0.071 0.086 0.015 21.0 1000-400 0.055 0.071 0.016 29.1 0.071 0.080 0.009 12.7 400-200 0.034 0.044 0.010 29.4 0.043 0.051 0.008 18.6 200-100 0.025 0.040 0.015 60.0 0.031 0.057 0.026 83.9 100-80 0.008 0.012 0.004 50.0 0.014 0.014 0.000 0.0 80-60 0.008 0.010 0.002 25.0 0.011 0.014 0.003 27.3 60-40 0.013 0.012 0.001 7.7 0.017 0.017 0.000 0.0 40-30 0.011 0.011 0.000 0.0 . 0.012 0.009 -0.003 -25.0 100-6000 0.153 0.204 0.051 33.3 0.216 0.274 0.058 26.9 TOTAL 0.193 0.249 0.058 30.1 0.270 0.328 0.058 21.5 (1) With,Carbon Black. (2) Regenerated. - 48 - Table 5 - Catalytic Properties After Steaming (1) of Samples 10R, 11R and 11RE Compared to Davison Super DX Catalyst Sample Sample Sample Super 10R 11R LIRE DX Mat Conversion, Vol.% 80 59.7 58.3 73 HPF 1.3 2.1 0.7 0.6 CPF 0.7 0.8 0.5 1.1 Coke, wgt% Feed 2.6 1.3 0.7 2.8 (1) Each sample was treated with steam (97%) for 5 hours at 1450°F to simulate a used catalyst at equilibrium FCC ccnditiσns.. - 49 - For comparison purposes with reference to the data of Table 5, two additional samples, designated as Sample Nos. 12 and 13, were prepared following the preparation procedure described above for Sample No. 10, but without the addition of carbon black. These samples yielded MAT volume percent conversions of 73.5 and 79.7, respectively, and carbon producing factors (CPF) of 0.92 and 0.84, respectively, with zeolite intensities- of 7.3 and 12.5, respectively. The catalyst of the invention has a relatively high level of cracking activity and is capable of providing high levels of conversion and selectivity at low residence times in the riser. The conversion capabilities of the catalyst may be expressed in terms of the conversion produced during actual operating of an RCC cracking or other conversion process and/or in terms of con version produced in standard catalyst activity tests. For example, it is preferred to employ a catalyst which, in the course of extended operation under prevailing process condition is sufficiently active for sustaining a level of conversion of at least about 50 percent and more preferably at least about 60 percent. In this connection, conversion is expressed as liquid volume percent based on volume of fresh feed. Conversion is th volume percentage of feedstock that is converted to 430°F end- point gasoline, lighter products and coke, and is calculated by subtracting from 100 the volume percentage -of those products heavier than the gasoline which remain in the recovered product Also, for example, the preferred catalyst may be defined as one which, in its virgin or equilibrium state, exhibits a specified activity expressed as a percentage in terms of MAT (micro- activity test) conversion. For purposes of the present inven¬ tion, the foregoing percentage is the volume percentage of stan ard feedstock which a catalyst under evaluation will convert to 430°F end point gasoline, lighter products and coke at 900°F, 16 WHSV (weight hourly space velocity) , and 3 C/O (catalyst to oil weight ratio) using the equipment and procedures specified - 50 - by ASTM D-32 MAT test D-3908-80 and an appropriate standard FCC feedstock. The WHSV is calculated on a moisture free basis using clean catalyst which has been dried at 1100°F, 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. The feedstock is preferably a sweet light primary gas oil, such as that used by the Davison Division of W. R. Grace and defined as follows: API Gravity at 60 °F, degrees 31.0 Specific Gravity at 60°F, g/cc 0.8708 Ramsbottom Carbon , wt.% 0.09 Conradson Carbo , wt.% ( St.) 0.04 Carbon, wt.% 84.92 Hydrogen, wt.% 12.94 Sulfur, wt.% 0.68 Nitrogen, ppm 305 Viscosity at 100°: F, centistokes 10.36 Watson K Factor 11.93 Aniline Point 182 Bromine No. 2.2 Paraffins, Vol.% 31.7 Olefins, Vol.% 1.8 Naphthenes, Vol.% 44.0 Aromatics, Vol.% 22.7 Average Molecular Weight 284 Nickel Trace Vanadium " Trace Iron Trace Sodium .Trace Chlorides Trace B S _ W Trace Distillation ASTM D-1160 IBP 445 10% 60 ' 1 30% 664 50% 701 70% 734 90% 787 FBP 834 The gasoline end point and the boiling temperature volume per¬ centage relationships of the products produced in the MAT con¬ version test may be determined by simulated distillation tech- - 51 - niques, for example by modification of the gas chromatographic "Sim-D" technique of ASTM D-2887-73. The results of such simu lations are in reasonable agreement with the results obtained by subjecting larger samples of material to standard laborator distillation techniques. The" catalyst may be introduced into theprocess ofthe invention in its virgin form or in other than its virgin form, e.g., one may use equilibrium catalyst withdrawn from another unit such as catalyst that has been employed in the cracking of an FCC feed. Whether characterized on the basis of MAT activity or relative activity, the preferred catalysts may be described on the basis of their activity "as introduced" into the process o the present invention, or on the basis of their "as withdrawn" or equilibrium activity in the process of the present inventio or on both of these bases. A preferred activity level of virgin and non-virgin catalyst "as introduced" into the process of the present invention is at least about 50%, preferably at least about 60% by MAT con¬ version. 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 50% and an activity level of 60% or more on a MAT conversion basis is also contemplated. More preferably, it is desired to employ a catalyst which will, under the conditions of use in the unit, establish an equilibrium activity at or above the indicated levels. Catalyst activities are determine with the catalyst having less than 0.01 coke, i.e., fully re¬ generated catalyst. In Table 5, "CPF" stands for Carbon Producing Factor and is de fined as the ratio of the amount of coke produced by the test catalyst to the amount of coke produced by a standard catalyst at the same conversion level. "HPF" stands for Hydrogen Pro- - 52 ducing Factor and is defined as the ratio of the amount of hy¬ drogen produced by the test catalyst to the amount of hydrogen produced by a standard catalyst at the same conversion level. The standard catalyst is chosen from among conventional FCC catalysts, such as for example, zeolite fluid cracking cata¬ lysts, and is chosen for its ability to produce a predetermine level of conversion in a standard feed under the conditions of temperature, WHSV (weight hourly space velocity), catalyst to oil ratio and other conditions set forth in the preceding de¬ scription of the MAT conversion test and in ASTM D-32 MAT test D-3907-80. For standard feed, one may employ the. above-mentio light primary gas oil, or an equivalent FCC feed. Although th equipment employed for the referenced MAT test is not capable of processing RCC feeds, applicants are undertaking to develop an equivalent test for evaluating RCC catalysts using RCC type feeds. On pages 935-937 of Hougen and Watson, "Chemical Process Princ ples", John Wiley & Sons, Inc., N.Y. (1947), the concept of "Activity Factors" is discussed. This concept leads to the us of "relative activity" to compare the effectiveness of an op¬ erating catalyst against a standard catalyst. Relative activi ty measurements facilitate recognition of how the quantity re¬ quirements 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 re¬ quired to produce a given level of conversion, as compared to the weight of an operating catalyst (whether proposed or actu¬ ally used) which is or would be required to product 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. For purposes of conducting relative activity determinations, on may prepare a "standard catalyst curve", a chart or graph of con version (as above defined) vs. reciprocal WHSV for - 53 - catalyst and feedstock. A sufficient number of runs is made under ASTM D-3907-80 conditions (as modified above) using standard feedstock at varying levels of WHSV to prepare an ac¬ curate "curve" of conversion vs. WHSV for the standard feed¬ stock. This curve should traverse all or substantially all of the various levels of conversion including the range of con¬ version within which it is expected that the operating cata¬ lyst will be tested. From this curve, one may establish a standard WHSV for test comparisons and a standard value of re¬ ciprocal WHSV corresponding to that level of conversion which has been chosen to represent 100% relative activity in the standard catalyst. For purposes of thepresent disclosure, the aforementioned reciprocal WHSV and level of conversion are, respectively, 0.0625 and 75%. In testing an operating cata¬ lyst of unknown relative activity, one conducts a sufficient number of runs with that catalyst under D-3907-80 conditions (as modified above) to establish the level of conversion which is or would be produced with the operating catalyst at standard reciprocal WHSV. Then, using the above-mentioned standard catalyst curve, one es tablishes a hypothetical reciprocal WHSV constituting the re¬ ciprocal WHSV which would have been required, using the standar catalyst, to obtain the same level of conversion which was or would be exhibited by the operating catalyst at standard WHSV. The relative activity may then be calculated by dividing the hypothetical reciprocal WHSV of the standard catalyst by the actual reciprocal WHSV of the test catalyst. The result is rela tive activity expressed in terms of a decimal fraction, which may then be multiplied by 100 to convert to % relative activity Relative Activity may also be expressed as follows: relative activity at constant conversion is equal to the ratio of the WHSV of the test catalyst divided by the WHSV of the standard catalyst. To simplify this calculation, a MAT conversion vs. relative activity curve was developed utilizing a standard catalyst of 75 vol.% conversion to represent 100% relative activity. One such curve is shown in Fig. 14. In applying the results of this determination, a relative activityof 0.5, - 54 - or 50%, means that it would take twice the amount of the opera ing or test catalyst to give the same conversion as the standa catalyst, i.e., the production catalyst is 50% as active as th reference catalyst. The relative activity level of virgin and non-virgin catalyst "as introduced" into the process of the present invention is a least about 20%, preferably at least about 40% and more pref¬ erably at least about 60%. However, it will be appreciated th 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 relative activity level of catalyst which has been used in the process of the present invention is at least about 20% or more, preferably about 40% or more, and more preferably about 60% or more. More preferably, it is desired to employ a catalyst which will, unde the conditions of use in the unit, establish an equilibrium ac¬ tivity at or above the indicated levels. O PI - 55 - Referring to the idealized carbon particle illustrated in Fig. 1, the individual reticulated CB particle "R" is made up of a series of primary particles "P" of roughly spherical shape hel together by atomic and/or interfacial binding forces. Attache to a central chain C-C are a number of side chains SI, S2 and S3, the actual number of side chains varying from particle to particle and the extent of such branching being indicated by a "branching factor", such as given in Figs. 9B through 12B. The diameter of the particle P determines the transverse diame ter of the central chain C-C and is usually about 100 to 1,000 preferably about 300 to 600A°. The length of the longest chai such as from point L to point M in Fig. 1, is usually about 500 to 30,000A°, preferably about 1,000 ' to 10,000A°. The aver age length to diameter ratio for such particles is preferably in the range of 2 to 5 or greater. It is believed that such reticulated particles yeild feeder pores roughly of the same dimensions as illustrated by the dotted outline of a feeder pore F in Fig. 1. The central chain and these side chains are believed to overla with those of adjacent reticulated particles and thereby pro¬ vide correspondingly tortuous channels and feeder pore branche in the matrix leading to the zeolite particles supported there in. In addition, two or more reticulated particles may adhere closely together to form an agglomerate mass so that the re¬ sulting pore size corresponds to the dimensions of this mass. Such agglomerates are believed to produce pore sizes of 1,000A and greater. The degree of agglomeration depends not only upo the particular dispersant used, but also upon the mixing means used and the mixing energy imparted to the carbon black sus¬ pension and resulting catalyst slurry. A greater percentage of pore sizes in the range of 400 to 600A° may be realized with catalyst compositions similar to those of Examples 1 to 9 but using increased agitation in combination with lignosul¬ fonate as the dispersant. OMPI - 56 - A scanning electron microscope (SEM) was used to make micro¬ graphs of some of the carbon black particles useful in the catalyst of the present invention. Copies of those micrograph are presented as Figs. 9A, 10A, 11A and 12A. Contour plots were also prepared of these reticulated carbon black particles by plotting the particle contours with an_X-Y recorder attached to a computer read-out from photogrammetric topography. The contourplots, together with tables of topographical measurement for each particle, are presented in Figs. 9B, 10B, 11B and 12B which correspond to carbon black types N-339, N-550, N-220 and N-326 shown in Figs. 9A-12A, respectively. Fig. 13 is a photograph showing a catalyst matrix from which t carbon black has been removed, the catalyst composition having been made and fired in accordance with the teaching of the in¬ vention. The photograph has been marked with horizontal co¬ ordinates 1 through 10 and vertical coordinates A-H. The dark areas of the photograph are pores and the horizontal and verti cal coordinates may be used to identify the location of those pores. The location and approximate size of the minor trans¬ verse dimension of some of the carbon black pores are identifi as follows: Dl-2, 300A°; G4, 1500A 0 ; B6-7, 250A°; C7-8, 700A°; D9, 600A°. As shown in the upper left hand corner of Fig. 13, the scale of this photograph is in micrometers (microns) , one- tenth of a micrometer being equal to 1,000 Angstrom units. The dimensional units of Fig. 13 are comparable to those of Figs. 9 through 12 where the scale is in nanometers (nm) , one nano¬ meter being equal to ten Angstrom units. - 57 - THE CONVERSION APPARATUS Although the above catalyst may be used in various types of con tacting operations within the contemplation of this invention, the catalyst is particularly useful in the catalytic cracking o residuum or carbo-metallic oil feeds. The catalytic cracking operation may be conducted in various types of reactors ' and as¬ sociated equipment, such as fixed bed systems or fluidized sys¬ tems. The catalyst is particularly useful in a fluidized bed type of operation where the catalyst in a finely divided flui¬ dized state is suspended in vaporized feedstock and the suspen¬ sion of catalyst and feedstock is passed upward through an elon gated riser in a progressive flow arrangement without signifi¬ cant backmixing. In general, riser operations are carried out at conditions conducive to achieving the improved results de¬ sired, bearing in mind the specific feedstock, catalyst composi tion and process equipment being used. The process is prefera¬ bly carried out without added hydrogen. Fig. 7 is a schematic diagram of one apparatus for carrying out the process of the present invention. The carbon-metallic oil feed is supplied through a feed supply pipe 10 having a control valve 11 to a wye 12 in which the feed is mixed with a flow of catalyst delivered through a catalyst supply pipe 13 as control by a valve 14. When used, water or some other diluent may be introduced to the wye through a diluent supply pipe 9. The mix ture of catalyst and feed, with or without additional diluent materials, then flows upward through a riser reactor 18. The riser 19 is an elongated conduit in which the length to diame¬ ter ratio may vary widely. The relatively narrow riser conduit provides high linear velocities, such as in the range of 25 to 90 feet per second. The length to diameter ratio of the riser is preferably in the range of 10 to 30, more preferably 20 to 25. - 58 - At the upper end of riser 18 is a chamber 19 which receives catalyst from the riser. While chamber 19 may be a conven¬ tional disengagement and collection chamber, often referred to as a reactor vessel, it is preferred that means be provided at the riser exit for causing product vapors to undergo a rapid change of direction relative to the direction traveled by the catalyst particles so that the vapors are suddenly or "ballis- tically" separated from the catalyst particles. Within disen¬ gagement chamber 19 of Fig. 7 is an upward extension 20 of riser pipe 18 having an open top 21 through which the catalyst particles are discharged. The product vapors are caused to un dergo a sudden change of direction into lateral port 22 in the side of riser extension 20, the catalyst particles because of their momentum being, for the most part, unable to follow the product vapors into port 22. This embodiment of the cracking apparatus makes use of the vented riser concept described in U.S. Patent Nos. 4,066,533 and 4,070,159 to Myers, et al. , whi patents are incorporated herein by reference. The vented rise thus provides a substantially instantaneous separation between hydrocarbon and catalyst in disengaging vessel 19. Because of the relatively high severity required to crack the more refrac tory components of carbo-metallic feeds, rapid disengagement o catalyst from cracked hydrocarbons by ballistic separation pre vents overcracking of desirable liquid products, such as gaso¬ line, to light undesirable products, such as hydrogen and metha gases and carbon. The product vapors and a relatively small amount of catalyst en trained with them are transferred by a cross pipe 23 to a cy¬ clone separator 24 for removing the entrained catalyst. The cy clone separator 24 may be connected via a transfer pipe 17 to an optional secondary cyclone separator 25. Product vapors are discharged from disengagement chamber 19 through overhead pro¬ duct discharge pipe 26. The catalyst particles which discharge from open top 21 of rise pipe extension 20, and those entrained catalyst particles which < g J E OMPI - 59 - are discharged from dip legs 27 and 28 of the primary and sec¬ ondary cyclones 24 and 25, respectively, drop to the bottom of disengagement chamber 19. Catalyst spilling over from the bot tom of disengagement chamber 19 passes via a drop leg 31 into a stripping chamber 32 equipped with baffles 33 and a steam je 34. Other stripping gases known in the art may be introduced through jet 34 and employed with or in place of steam. Accumulated carbon is burned from the catalyst in a combuster 38 which receives stripped catalyst via a downcomer pipe 39 with a control valve 40. Combustion air is supplied from an air supply system, generally indicated by 44 and including blo ers 41 and 42 and filter bank 43, to combustion air jets 48 at the bottom of the combuster and to fluffing air jets 49 at an elevated position therein. Regenerated catalyst, with most of the carbon burned off, departs the combuster 38 through an up¬ per outlet 50 and cross pipe 51 to a secondary combustion cham 52, where it is deflected into the lower portion of chamber 52 by baffle 53. Although the use of two-stage regeneration is contemplated and preferred, in this particular embodiment, the secondary chamber 52 is operated primarily as a separator cham for separating regenerated catalyst from combustion gases. Ad ditional oxygen containing gas may be introduced through lines 59 and secondary chamber 52 operated with an excess of oxygen to insure removal of carbon deposits down to residual carbon levels of 0.05 weight percent or less. Catalyst may optionally move in up to three different directio from secondary chamber 52. Fully regenerated catalyst is dis¬ charged through bottom outlet 69 to catalyst supply pipe 13 through which it is recirculated to wye 12 for contact with fresh feed as previously described. A portion of thehot cata¬ lyst may be circulated back to combuster 38 via catalyst recir culation line 55 with control valve 56 for heat control in the combuster and to raise the temperature of incoming air with ho catalyst so as to insure that combustion of coked catalyst is prop erly initiated. Since some of the catalyst will be entrained the combustion gases, such as carbon oxides produced by burning - 60 - erally indicated by 57 and 58, are provided in chamber 52 for separating these catalyst fines from the combustion gases. Catalyst collected in cyclones 57, 58 is discharged through their dip legs to the bottom of chamber 52 where the catalyst is kept in suspension by air, inert gas and/or steam from line 59 and by a baffle arrangement 54. Combustion product gases produced by regeneration of the catalyst and separated from en trained catalyst fines are discharged through effluent pipes 61, 62 and heat exchangers 60,63. If such gases contain suffi cient amounts of carbon monoxide, they may be sent via gas sup ply pipe 64 to an optional CO boiler in which the CO is burned in order to heat heating coil 66 connected with a steam boiler 67. The amount of heat passed from the regenerator back to the ris is preferably controlled, at least in part, by controlling the flow of catalyst through catalyst standpipe 13 by control valv 14 which is preferably a type of slide valve operated by suit- able automatic control equipment (not shown) responsive to the temperature of product vapors downstream of riser outlet 22. If additional heat removal from the combustion chambers is re¬ quired, conventional direct or indirect cooling techniques may be employed in the manner known to persons skilled in the art of designing and operating regenerators. Since the preferred conversion reactor is of the progressive flow type and is operated so that no dense bed of catalyst bui up within the riser, space velocities in the riser are usually high and will generally fall within the range of about 100 to 600 weight of hydrocarbon per hour per instantaneous weight of catalyst occupying the riser volume. In the absence of sig¬ nificant catalyst build up, the instantaneous catalyst inven¬ tory within the riser volume is represented by the catalyst particles suspended with the oil feed. - - A particularly preferred emobidment is described in Fig. 15 where reference numeral 80 identifies a feed control valve in feedstock supply pipe 82. Supply pipe 83 (when used) introduc liquid water into the feed. Heat exchanger 81 in supply pipe acts as a feed preheater, whereby preheated feed material may delivered to the bottom of riser type reactor 91. Catalyst is delivered to the reactor through catalyst standpipe 86, the fl of catalyst being regulated by a control valve 87 and suitable automatic control equipment (not shown) with which persons ski in the art of designing and operating riser type cracking units are fa liar. The reactor is equipped with a disengagement chamber 92 simila to the disengagement chamber 19 of the r?=ctor _r.z-. ~ -. i Fig. 1 Catalyst departs disengagement chamber 92 through stripper 94. Spent catalyst passes from stripper 94 to regenerator 101 via spent catalyst transfer pipe 97 having a slide valve 98 for controlling flow. Regenerator 101 is divided into upper chamber 102 and lower chamber'103 by a divider panel 104 intermediate the upper and lower ends of the regenerator. The spent catalyst frorn ^ trans¬ fer pipe 97 enters upper chamber 102 in which the catalyst is partially regenerated. A funnel-like collector 106 having a bias-cut upper edge receive partially regenerated catalyst fro the upper surface of the dense phase of catalyst in upper cham ber 102 and delivers it via drop leg 107 having an outlet 110 beneath the upper surface of the dense phase of catalyst in lo er chamber 103. Instead of internal catalyst drop let 107, on may use an external drop leg. Valve means in such external dr let can control the residence time and flow rate in and betwee the upper and lower chambers. Air is supplied to the regenerator through an air supply pipe1 A portion of the air travels through a branch supply pipe 114 bayonet 115 extending upwardly in the interior of plenum 111 a long its central axis. Catalyst in chamber 103 has access to the space within plenum 111 between its walls and bayonet 115. A small bayonet (not shown) in the aforementioned space fluf s the catalyst and urges it upwardly toward a horizontally arran - 62 - ring distributor (not shown) where the open top of plenum 111 opens into chamber 103. The remainder of the air passing thr air supply pipe 113 may be heated in air heater 117 (at least during start-up with VGO) and is then introduced into inlet 1 of the ring distributor, which may be provided with holes, no zles or other apertures which produce an upward flow of gas t fluidize the partially regenerated catalyst in chamber 103. The air in chamber 103 completes the regeneration of the part regenerated catalyst received via drop let 107. The amount o air supplied is sufficient so that the resultant combustion g are still able to support combustion upon reaching top of cha 103. Drop leg 107 extends through an enlarged aperture in pa 104, to which is secured a gas distributor 120 which is concen tric with and surrounds the drop leg. Combustion supporting gases, which have been partially depleted, are introduced via gas distributor 120 into upper regenerator chamber 102 where they contact incoming coked catalyst from coked catalyst trans fer pipe 97. Apertured probes 121 in gas distributor 120 assi in achieving a uniform distribution of the partially depleted bustion supporting gas into upper chamber 102. Supplemental a or cooling fluids may be introduced into upper chamber 102 thr a supply pipe 122, which may discharge through gas distributor Fully regenerated catalyst with less than about 0.25% carbon, preferably less than about 0.1% andmore preferably less than a bout 0.05%, is discharged from lower regenerator chamber 103 through regenerated catalyst stripper 128, whose outlet feeds to catalyst standpipe 86. Thus, regenerated catalyst is retur to riser 91 for contact with additional fresh feed. The division of the regenerator into upper and lower regenerat chambers 102 and 103 not only smooths out variations in cataly regenerator residence time but is also uniquely of assistance resticting the quantity of regeneration heat which is imparted to the fresh feed while yielding a regenerated catalyst with l levels of coke and/or carbon black for return to the riser. - .3 - Because of the arrangement of the regenerator, coked catalyst transfer line 97 and/or virgin carbon black catalyst from addi tion line 99, with a high loading of carbon, contacts in chamb 102 combustion supporting gases which have already been at lea partially depleted of oxygen by the burning of carbon from par tially regenerated catalyst in lower chamber 102. Because of this, it is possible to control both the combustion of carbon the quantity of carbon dioxide produced in upper regenerator c ber 102. Although regenerating gas introduced through air sup pipe 113 and branch conduit 114 may contain relatively large quantities of oxygen, the partially regenerated catalyst which contacts in lower -chamber 103 has already had a major portion its carbon removed. The high oxygen concentration and tempera in chamber 103 combine to rapidly remove the remaining carbon the catalyst, thereby achieving a clean, regenerated catalyst with a minimum of heat release. Thus, here again, the combust temperature and CO to CO.-, ratio in the lower chamber are readi controlled. The regeneration off gases are discharged from up chamber 102 via gas pipe 123, regulator valve 124, catalyst fi trap 125 and outlet 126. The vapor products from disengagement chamber 92 may be proces in any convenient manner such as by discharge through vapor li 131 to fractionator 132. Fractionator 132 includes a bottoms outlet 133, side outlet 134, flush oil stripper 135, and stripp bottom line 136 connected to pump 137 for discharging flush oi Overhead product from stripper 135 returns to fractionator 132 via line 138. The main overhead discharge line 139 of the fractionator is co nected to receiver 142 having a bottoms line 143 feeding into pump 144 for discharging gasoline product. A portion of this product may be returned to the fractionator via recirculation line 145, the flow being controlled by valve 146. The- receive also includes a water receiver 147 and a water discharge line The gas outlet 150 of the overhead receiver discharges a strea which is mainly below C g , but containing some C 5 , C β and C-, ma terial. If desired, the C g and above material in the gas stre may be separated by compression, cooling and fractionation, an recycled to receiver 142. - 64 - The oxidizing gas, such as air, introduced into regeneration zone 103 through line 114 may be mixed with a cooling spray of water from a conduit 109. The mixture of oxidizing gas and atomized water flows through bayonet 115 and thus into the lower bed of catalyst particles. The apertures in distributor 120 are large enough so that the upwardly flowing as readily passes therethrough into zone 102 However, the perforations are sized so that the pressure dif¬ ference between the upper and lower zones prevents catalvst particles from passing downwardly through the distributor. The bayonet 115 and distributor are similarly sized. Gases exiting the regenerator comprise combustion products, nitro¬ gen, steam formed by combustion reactions and/or from vaporiz¬ ing water added with the oxidizing gas, and oxides of sulfur and other trace elements. These gases are separated from sus¬ pended catalyst particles by a cyclone separator (not shown) and then pass out of the regenerator through discharge con¬ duit 123. While this invention may be used with single stage regenerator or with the multiple stage regenerator of Fig. 7, which has basically concurrent instead of countercurrent flow, it is es¬ pecially useful in a regenerator of the type shown in Fig. 8, which is well-suited for producing combustion product gases having a high ratio of CO to C0 2 , which helps lower regenera¬ tion temperatures in the presence of high carbon levels. .THE CONVERSION PROCESS The catalysts described in this specification may be employed the processes and apparatuses for carbo-metallic oil conversio described in co-pending U.S. applications Serial Nos. 94,091; 94,092; 94,216; 94,217; and 94,277 of Ashland Oil, Inc., each of said co-pending applications having been filed on November 14, 1979, and being expressly incorporated herein by reference QMPI - b _- - Furthermore, the adverse effects of vanadium deposited on the catalyst may be controlled as described in co-pending interna¬ tional application No. PCT/US81/00356 of Ashland Oil, Inc., filed on March 19, 1981, and entitled IMMOBILIZATION OF VANA- DIA DEPOSITED ON CATALYTIC MATERIALS DURING CARBO-METALLIC OIL CONVERSION, which application is incorporated herein by reference. A preferred method for cracking feeds with heavy fractions, re¬ ferred to as Reduced Crude Conversion (RCC) after a particularl common and useful carbo-metallic feed, is disclosed in co-pend¬ ing application S.N. 94,216 filed on November 14, 1979, for "Carbo-Metallic Oil Conversion", which application is incorpor¬ ated herein by reference. The preferred feeds capable of being cracked by this method are comprised of 100 percent or less of the 650°F+ material previously described. The cracking reactio according to the method disclosed in S.N. 94,216 is sufficientl severe to convert 50 to 90 percent of the carbo-metallic oil feed to gasoline per pass and produce coke in amounts of 4 to 14 percent by weight based on weight of fresh_ feed. This coke is laid down on the catalyst in amounts in the range of about 0.3 to 3 percent by weight of catalyst, depending upon the cata lyst to oil ratio (weight of catalyst to weight of feedstock) in the riser. The feed, with or without pretreatment, is introduced as shown in Fig. 7 into the bottom of the riser along with a suspension of hot cracking catalyst prepared in accordance with this in¬ vention. Steam, naphtha, water and/or some other diluent is preferably introduced into the riser along with the feed. Thes additives may be from a fresh source or may be recycled from a process stream in the refinery. Where recycle additive streams are used, they may contain hydrogen sulfide and other sulfur compounds which may help passivate adverse catalytic activity by heavy metals accumulating on the catalyst. It is to be understood that the water additive may be introduced either as a liquid or as steam. Water is added primarily as a -^DRE > OMPI - 66 - source of vapor for accelerating the feed and catalyst to achieve the vapor velocity and residence times desired. Other diluents as such need not be added but where used, the total amount of diluent specified includes the amount of water used. Extra diluent would further increase the vapor velocity and further lower the feed partial pressure in the riser. As the feed travels up the riser, it is catalytically cracked to form basically five products known in the industry as dry gas wet gas, cat naphtha., light, cycle oil, and heavy cycle and/or slurry oil. At the upper end of the riser, the catalyst particles are ballistica separated from product vapors as previously described. The catalyst which then contains the coke formed in the riser is sent to the regenerator to burn off the coke and the separated product vapors are sent to a fractionator for further separa¬ tion and treatment to provide the five basic products indicate The preferred process conditions for the riser conversion re¬ action are set forth in Table 3, in which the abbreviations used have the following meanings: "Temp." for temperature, "Dil." for diluent, "pp" for partial pressure, "wgt" for weigh "V" for vapor, "Res." for residence, "C/O" for catalyst to oil ratio, "Cat." for catalyst, "bbl" for barrel, "MAT" for micro- activity by the MAT test using a standard Davison feedstock, "Vel." for velocity, "cge" for charge, "d" for density, and "Reg." for regenerated. - ϋREΛ> QMPI 67 - TABLE 6 - RCC RISER CONDITIONS Board Operating Preferred Parameter Range Range Feed Temp. - 400-800°- F 400-650° F Steam Temp. - 200-500° F 300-400° F Reg. Catalyst Temp. - 1100-1500° F * 1275-1450° F Riser Exit Temp. - 900-1400° F 950-1100° F Pressure - 0-100 psia 10-50 psia Water/Feed - 0.05-0.30 0.05-0.15 Dil. pp/Feed pp - 0.25-3.0 1.0-2.5 Dil. wgt/Feed wgt - <p.4 0.1-0.3 V. Res. Time - 0.1-5 0.5-3 sec. C/O, wgt. - 3-18 5-12 " Lbs. Cat./bbl Feed - 0.1-4.0 0.2-2.0 Inlet Cat. MAT - >50 vol. 7β >60- Outlet Cat. MAT -" vol. X >40 V. Vel. - 25-90 ft./sec. 30-60 V. Vel./Cat. Vel. ■ >1.0 _ 1.2-2.0 Dil. Cge. Vel. - 5-90 ft./sec. 10-50 . Oil Cge. Vel. - 1-50 ft./sec. 5-50 Inlet cat. d. - 1-9 lbs./ft. 3 2-6 Outlet cat. d. - 1-6 lbs./ft. 3 1-3 Coke, wgt.%, coked cat. 0.6-1.5 0-8-1.3 O PI - 68 - In cracking carbo-metallic feedstocks in accordance with the method of S.N. 94,216, the regenerating gas may be any gas which can provide oxygen to convert carbon to carbon oxides. Air is highly suitable for this purpose in view of its ready availability. The amount of air required per pound of coke for combustion depends upon the desired carbon dioxide . to car bon monoxide ratio in the effluent gases and upon the amount of other combustible materials present in the coke, such as hydrogen, sulfur, nitrogen and other elements capable of form ing gaseous oxides at regenerator conditions. The regenerator is operated at temperatures in the range of about 1,000 to 1,600°F, preferably 1,275 to l,45υ°F, to acnie adequate combustion while keeping catalyst temperatures below those at which significant catalyst degradation can occur. I order to control these temperatures, it is necessary to con¬ trol the rate of burning which in turn can be controlled at least in part by the relative amounts of oxidizing gas and carbon introduced into the regeneration zone per unit time. The rate of introducing carbon into the regenerator may be controlled by regulating the rate of flow of coked catalyst through valve 40 in conduit 39, the rate of removal of regen¬ erated catalyst by regulating valve 14 in conduit 13, and the rate of introducing oxidizing gas by the speed of operation of blowers 41, 42. These parameters are regulated such that the ratio of carbon dioxide to carbon monoxide in the efflu¬ ent gases is equal to or less than about 4.0, preferably 1.5 o less. In addition, water, either as liquid or steam, may be added to the regenerator to help control temperatures and to influence the carbon dioxide to carbon monoxide ratio. The regenerator combustion reaction is carried out so that the amount of carbon remaining on regenerated catalyst is less tha about 0.25, preferably less than about 0.05 percent and most preferably less than about 0.01 percent on a substantially moi ture-free weight basis. The residual carbon level is ascer¬ tained by conventional techniques which include drying the - 69 - catalyst at 1,100°F for about four hours before measuring car¬ bon content. In the method of regeneration of the present invention, the amount of oxidizing gas and catalyst are controlled so that the amount of oxidizing gas passing into the second zone is greater than that required to convert all of the coke remain¬ ing on the catalyst reaching this oxygen rich zone to carbon dioxide. On the other hand, the amount of oxidizing gas pass¬ ing into the first zone from the second zone and the additiona oxidizing gas added to the first zone, such as through line 116 and distributor 118 in Fig. 8, is insufficient to convert all of the coke in this zone to carbon dioxide. As the first zone is therefore oxygen deficient, significant afterburning and excessive temperatures in the first zone are prevented, thereby protecting both the catalyst and regeneration equip¬ ment from damage. Industrial Applicability The present invention is particularly useful in the catalytic cracking of high boiling carbo-metallic feedstocks to lower boiling hydrocarbon fractions in the liquid fuel " range. Exam¬ ples of these oils are reduced crudes and other crude oils or crude oil fractions containing residua as hereinabove defined. However, it is to be understood that the catalyst and processe of the present invention are useful for cracking almost any crude oil or crude oil fraction, including conventional light and heavy gas oils! Although the process is preferably conducted in a riser reacto of the vented type, other types of risers with either upward o downward flow or non-riser type reactors may be employed. Thu the cracking operation may be conducted with a fluid bed of catalyst through which the feedstock is passed under suitable contact conditions of pressure, temperature and feed rate. Al ternately, the catalyst may be used in the form of a moving be passing through or otherwise contacted with the feedstock ma- OMPI - 70 - ■ terials to be cracked. Although the preferred contacting operation is catalytic crac ing, the catalyst and processes of the invention may be emplo in various other types of hydrocarbon conversion operations, such as dehydrocyclization, hydrocracking, hydroforming of naphthene hydrocarbons and the like," polymerization of olefin depolymerization of polymers, alkylation, dealkylation, dis- proportionation, reforming of naphthas, isomerization of para fins and the like, aromatization of paraffins and the like, hydrogenation, dehydrogenation, various types of hydrofining operations in which one or more characteristics of the feed¬ stock are improved by treatment with hydrogen in the presence of a catalyst, oxidation of organic compounds with an oxida¬ tion media such as air, adsorption and absorption operations, and like types of other contacting, conversion and/or separa¬ tion processes. With the present invention, greater cracking selectivity to gasoline boiling range products and less carbon make for a given level of conversion may be achieved because of the grea er accessibility of reaction sites to reactants and reduction in undesirable side reactions. Also the processes may be op¬ erated at higher space velocities and/or lower reaction temp¬ eratures to achieve a given selectivity, conversion and oc¬ tane number as compared to conventional processes employing conventional catalysts. The catalyst is generally easier to regenerate than prior art catalyst whether by conventional regeneration processes or those of the invention because dif¬ fusion limitations on oxidation reactions within the catalyst are minimized. The large pore CB catalyst has high tolerances to both metals and coke and are therefore highly suitable for cracking any charge stock containing metal contaminates and/or coke pre¬ cursors. By ' "high tolerance" is meant that the catalyst can accumulate significantly greater quantities of poison metals - 71 and coke as compared to prior art catalysts and still have ef¬ fective catalytic activity in the processes of the invention. During the conversion reaction, metals and coke may deposit on the walls of the large feeder pores, rather than on the zeo lite where they can block substantial numbers of catalytic sites. The large feeder pore matrix thereby prolongs the use¬ ful life of the zeolite promoter. The large feeder pores are also capable of absorbing asphaltenes and other liquid hydro¬ carbons boiling above reactor temperatures without becoming filled to a degree significantly retarding diffusion of lower boiling reactant molecules. Because -che large feeder pores provide better diffusion cf . - large molecules they may be cracked at acidic sites instead of thermally at or near the catalyst surface. There is therefore less coke laydown for a given level of conversion. Diffusion during regeneration is also improved so that temperatures of localized hot spots will be considerably less and correspond¬ ingly there is less matrix and zeolite pore collapse due to sintering. Furthermore, with larger pore diameters, partial collapse of feeder pores restrict hydrocarbon diffusion to a lesser degree than with conventional catalysts. ' The large feeder pores are also less likely to be blocked by the disposi tion of heavy metals and coke on particle surfaces. One factor reducing the quantities of coke to be removed in the regenerator is the relative ease with which both uncracked and cracked hydrocarbons can be stripped from the large feeder pores. Another factor reducing coke deposits, is that asphal¬ tenes and other hydrocarbons remaining in a liquid state are more likely to be cracked by reaching an acid site either in the matrix material or on an exposed surface of the zeolite. A further industrial advantage of the invention is the ease with which a gel slurry of the catalyst composite can be spray dried and formed into attrition resistentmicrospheres of a controlled, fluidizable size. After formation, these - 72 - microspheres may be easily washed and redried and then fired to remove the carbon black without significant alteration of either the zeolite or the support structure of the matrix. An especially important feature of the invention is that the carbon black can be removed during initial circulation of the catalyst through the reactor and regenerator of an operating conversion unit, rather than requiring removal during catalyst manufacture. OMFI

Claims: - 73 - Claims What is claimed is: 1. A process for preparing a hydrocarbon conversion catalyst which comprises introducing particles of carbon black into a shapable catalyst composition having hydrocarbon conversion characteristics, shaping said catalyst composition into solid shapes of hydrocarbon conversion catalyst, and heating said solid catalyst shapes in the presence of an oxygen containing gas at combustion conditions sufficient to remove said carbon black particles and provide large feeder pores in said solid catalyst shapes. 2. The process of claim 1 in which a promoter containing cat lytic sites for hydrocarbon conversion is provided in said solid catalyst shapes and said shapable catalyst composition includes a matrix material for supporting said promoter in sai solid catalyst shapes. 3. The process of claim 2 in which said promoter is a natural or synthetic zeolite. 4. The process of claim 1 in which said shapable catalyst co position includes silica and alumina materials capable of for ing crystalline aluminosilicates, and in which said aluminosil cates are crystallized after at least a portion of said carbon black particles are mixed with said silica alumina macerials. 5. The process of claim 1 in which the effective diameter of at least a portion of said feeder pores provided by removal of said carbon black particles is in the range of 100 to 6,000, or 400 to 6,000, or 1,000 to 6,000 Angstroms. 6. The process of claim 1 in which the pore volume provided b removal of said carbon black particles is at least 0.1 cc/gm. 7. The process of claim 1 in which said solid catalyst shapes after said heating contain cation exchanged zeolite in an amou - 74 - in the range of about 5 to 45 weight percent. 8. The process of claim 1 in which said shaped catalyst com tion contains carbon black in the range of 5 to 25 weight pe cent prior to said heating step. 9. The process of claim 1 in which said shaped catalyst com position contains carbon black in the range of 10 to 20 or 15 to 18 weight percent prior to said heating step. 10. The catalyst composition prepared according to the proce of claim 1. 11. A cracking catalyst comprising a cation exchanged zeolit in an amount in the range of about 10 to 40 weight percent, a a silica, alumina or silica-alumina matrix in an amount in th range of about 15 to 90 weight percent, said catalyst having a pore volume of at least 0.10 cc per gm formed by introducin particles of carbon black into said matrix and subsequently heating said matrix in the presence of an oxygen containing g under combustion conditions sufficient to remove said carbon black particles from said matrix and provide said pore volume 12. A process for preparing a hydrocarbon conversion catalys comprising the steps of: preparing an aqueous- slurry of a hydrogel suitable for f ing a solid catalyst matrix; adding carbon black particles to said hydrogel slurry; processing said mixture to form solid shapes of hydro¬ carbon conversion catalyst containing said carbon black parti cles in a matrix formed from said hydrogel slurry; and, heating said solid catalyst shapes in the presence of an oxygen containing gas under combustion conditions sufficient to remove said carbon black particles and provide large feede pores in said solid catalyst shapes. 13. The process of claim 12 in which said aqueous slurry con tains a siliceous gel or a colloidal dispersion of silica__σel, - 75 - 14. The process of claim 12 in which said aqueous slurry con¬ tains an alumina gel. 15. The process of claim 12 in which said aqueous slurry con¬ tains a silica gel or alumina gel or both a silica gel and a alumina gel. 16. The process of claim 12 in which said hydrogel slurry con tains a lignin compound as a dispersant for said carbon black particles. 17. The process of claim 16 in which said lignin compound is a sodium lignosulfate and the amount of said lignin compound in said hydrogel slurry is in the range of 0.05 to 3.0 or 0.1 to 2.0 or 0.5 to 1.5 weight percent based on total weight of solids in said hydrogel slurry after said addition of carbon black particles. 18. The catalyst composition prepared according to the proces of claim 12. . .. 19. The catalyst composition of claim 18 which further contai a cation exchanged Y or X or Y-X zeolite in an amount in the range of about 3 to 45 weight percent, said composition being " further exchanged with rare earth ions following said combus¬ tion of carbon black particles to provide a zeolite catalyst of improved hydrothermal stability. 20. The process of claim 12in which a crystalline aluminosili cate zeolite is added to said mixture prior to processing sai mixture to form said solid catalyst shapes. 21. The process of claim20 wherein said zeolite is a Type Y or Type X zeolite. 22. The process of claim21 wherein said zeolite is exchanged with hydrogen, ammonium or rare earth ions or a mixture of sai ions. OMPI_ 23. The process of claim 12in which said hydrogel slurry further contains sufficient clay to give a solid combust catalyst with up to 75 percent clay by weight. 24. A composition suitable for use as a catalyst in a cata¬ lytic cracking process which includes contacting the catalys with a hydrocarbon feedstock containing an appreciable amoun of a high molecular weight hydrocarbon component under reac¬ tion conditions depositing coke on said catalyst, and regen¬ erating said catalyst under oxidizing conditions to remove said coke therefrom, said catalyst composition comprising a silica material in an amount within the range of about 15 to 99 percent by weight of the total composition on a volatil free basis and carbon black particles in an amount within the range of about 1 to 35 percent by weight of the total composi tion on a volatile free basis, said carbon black being remov¬ able from said catalyst composition under the oxidizing con¬ ditions of said catalyst regeneration to provide large feeder pores in said silica material. \ 25. The catalyst composition of claim 24 which further includ a quantity of clay in an amount within the range of 0 to 75 percent by weight of said catalyst composition after removal of said carbon black. 26. A process for the catalytic conversion of hydrocarbon feedstock containing an appreciable amount of heavy hydrocar¬ bon components boiling above 1,025°F comprising contacting said feedstock with the catalyst of claim 1 in a reaction zon maintained under hydrocarbon conversion conditions providing hydrocarbon conversion products and depositing coke on said catalyst, stripping said hydrocarbon conversion products from said catalyst, regenerating said catalyst under oxidizing con ditions sufficient to remove said deposited coke from said catalyst, and returning said regenerated catalyst to said re¬ action zone. 27. A process for converting to lighter products a hydrocar¬ bon feedstock containing an appreciable amount of high molecu lar weight hydrocarbons boiling above 1,025°F which comprij - 77 - which has not been hydrotreated and includes an appreciable amount of a hydrocarbon fraction boiling above 1,025°F; contacting said converter feed with the catalyst composi¬ tion of claim 1 in the presence of added gaseous material and under conditions sufficient to form a suspension of said cata¬ lyst in said feed and said gaseous material wherein the ratio of the partial pressure of the added gaseous material relative to that of said feed is in the range of about 0.25 to about 2. causing said suspension to flow through an elongated vert cally extending reaction chamber of a progressive flow type re actor for a predetermined vapor residence time in said reactio chamber in the range of about 0.1 to about 3.0 seconds at a temperature in the range of about 900 to about 1,400°F and un¬ der a total pressure in the range of about 10 to about 50 poun per square inch absolute, said time, temperature and pressure and the weight ratio of catalyst to oil in said suspension be¬ ing sufficient for causing a conversion per pass in the range of about 50 to about 90 volume percent and for laying down cok on the catalyst in amounts in the range of about 0.3 to about 3% by weight; separating said catalyst from the resultant conversion products; stripping absorbed hydrocarbons from said separated cata¬ lyst; regenerating said stripped catalyst with an oxygen con¬ taining combustion gas under combustion conditions sufficient to reduce the carbon on the catalyst to less than about 0.25% by weight while forming combustion product gases comprising CO and/or C0 2 ; and, recycling said regenerated catalyst to said reactor for contact with fresh feed. - 28. The conversion process of claim 27 in which said 650°F+ material is further characterized by a carbon residue upon pyrolysis of at least about 1.0 and a heavy metals content of at least about 4.0 parts per million of Nickel Equivalents, an in which at least a portion of the feeder pores provided in sa conversion catalyst by the removal of said carbon black have effective pore diameters in the range of 400 to 6,000 Angstrom 29. The conversion process of claim 28 in which said 650°F+ material is further characterized by a carbon residue upon pyro sis of at least about 2.0, a heavy metals content of at least about 10 ppm of Nickel Equivalents, andanitrogen content of at least 0.1 weight percent. 30. The conversion process of claim 28 in which said 650°F+ material is further characterized by a carbon residue upon pyrol sis of at least about 4.0, a heavy metals content of at least about 20 ppm of Nickel Equivalents, and a nitrogen content of at least 0.2 weight percent.

Inventors: BECK WAYNE H, HETTINGER WILLIAM P JR, LEWIS JAMES E
Applicants: ASHLAND OIL INC, BECK WAYNE H, HETTINGER WILLIAM P JR, LEWIS JAMES E
Application year: 1981
Main classification: B01J29
IPC codes: B01J29, B01J37, C10G11