US3168460A

US3168460A - Catalytic conversion of heavy oils to gasoline

Info

Publication number: US3168460A
Application number: US149272A
Authority: US
Inventors: Leon M Lehman; Manne Stanley
Original assignee: Sinclair Research Inc
Current assignee: Sinclair Research Inc
Priority date: 1961-11-01
Filing date: 1961-11-01
Publication date: 1965-02-02
Anticipated expiration: 1982-02-02

Description

Feb. 2, 1965 L. M. LEHMAN r-:TAL

CATALYTIC CONVERSION oF HEAVY ons To GAsoLINE Filed Nov. ll.. 1961 United States Patent 3,168,460 CATALYTEC CNVERSHN F HEAVY GILS T GASGLINE Leon M. Lehman, Brooklyn, N.Y., and Stanley Manne,

Park Forest, lll., assignors to Sinclair Research, Inc.,

Wilmington, Dei., a corporation of Delaware Filed Nov. l, 1961, Ser. No. l49,272 1t) Claims. (Cl. 208-67) This invention concerns the catalytic cracking of hydrocarbon oils to obtain lighter components including gasoline of relatively high octane number. In this invention catalytic cracking may be performed upon `an essentially whole crude petroleum or other mineral oil, rather than merely a fraction of the crude. Also, this invention provides for fuller utilization of the values in residual oils, including whole crudes, by providing for benefici-ation of the heavy recycle oils by hydrogenation. The invention employs catalyst demetallization to overcome the effect of catalyst poisons in the residual feed. It can thus be seen that this invention provides a process in which mineral oil hydrocarbon feedstocks not customarily employed for catalytic cracking are subjected to such conversion.

The catalytic cracking of various heavier mineral hydrocarbons, for instance petroleum or other mineral oil products is now practiced to a considerable extent on gas oils. These may be any one of a variety of distillate stocks. The term gas oil includes any fraction distilled from petroleum or other mineral oil which has an initial boiling point of at least about 400 F., say, up to about 850 F., and an end boiling point of at least about 600 F., and generally boiling substantially continuously between the initial boiling point and the end boiling point. Usually the boiling range extends over at least about 100 F. The portion which is not distilled before the end point is reached is considered residual stock. The exact boiling range of a gas oil, therefore, will be determined by the initial distillation temperature (initial boiling point) and by the temperature at which distillation is cut off (end boiling point). In practice, petroleum distillations have been made under vacuum up to temperatures as high as about 1l00-l200 F. (corrected to atmospheric pressure). Accordingly, in the broad sense, a gas oil is a petroleum fraction which boils substantially continuously between two temperatures that establish a range falling Within from about 400 F. to about l1001200 F. Thus, a gas oil could boil over the entire range of about 400- 1200 F. or it could boil over a narrower range, e.g., about 500-900 F. The gas oils can be further roughly classified by boiling ranges. Thus, a gas oil boiling between about 400-500 F. and about 60G-650 F. is termed a light gas oil; a medium gas oil distills between about 60G-650 F. and about 80G-900 F.; a gas oil boiling between about 800-850 F. and about 1l001200 F. is sometimes designated as a vacuum gas oil. It must be understood, however, that a particular stock may bridge two boiling ranges, or even span several ranges, i.e., include, for example, light and medium gas oils.

As can be seen, a gas oil is the product of a distillation, usually performed on a crude oil before the oil is subjected to any processing other than desalting. The forerun of the distillation, that is, the normally liquid fractions of the crude boiling below the gas oil range is generally termed straight-run gasoline or light ends. Such fractions are conventionally used directly as motor fuels, or may be further processed, as by reforming, to increase their octane value.

A residual stock is in general any mineral oil fraction which includes all the heavy bottoms, such as tars, asphalts, etc.; which are not distilled at a selected temperature and pressure. Thus, a whole crude is a residual oil.

3,158,460 Patented Feb. 2, 1965 ICC Also, a whole topped crude, which is the entire portion of the crude remaining after the light ends (the portion boiling up to about 400 F have been removed by distillation, is a residual. Such a fraction includes the entire gas oil fraction (400 F. to T100-1200 F.) land the undistilled portion of the crude petroleum boiling above about 1l00-l200 F. Alternatively, the residual may be conned to those portions of the crude remaining undistilled at about ll00-l200 F., or it can be made up of a gas oil fraction plus the portion undistilled at about 1100- l200 F. Reduced crudes also are residuals. By a whole crude is meant a petroleum or other mineral hydrocarbon oil which has constituents in the full range of liquid hydrocarbons and which has not been processed except perhaps for desalting. Metal poisons in a crude oil stock tend to remain in the undistilled, residual, portion, although metal contaminants frequently also Iare entrained in distillates.

The behavior of a hydrocarbon feedstock in the cracking reactions depends upon various factors including its boiling point, carbon-forming tendencies, content of catalyst contaminating metals, etc., and these characteristics may affect the operation to an extent which makes a given feedstock uneconomical to employ. A large coke deposit on a catalyst seriously limits the cracking activity of a catalyst. Metal poisons, especially nickel and vanadium, deposited on a catalyst seriously affect cracking operations by increasing the hydrogen content of the gases produced. There is also a significant increase in coke production at a given conversion level and a corresponding decrease in yield of liquid products. The adverse economic eiects of these factors may be extremely high as a result of loss in unit capacity, and decrease in liquid yield. Although the cracking catalyst employed can be discarded more often to prevent a high accumulation of poisoning metals in the cracking system, this type of operation represents a substantial cost factor. The feedstock characteristics become even more important as the cost of the catalyst rises and thus the effects of low feedstock quality are particularly burdensome in systems employing cracking catalysts containing relatively expensive synthetic components. In addition, it has generally not been thought useful to submit the light ends to catalytic cracking.

Frequently the reiiner may take special cuts of metalscontaminated stocks usually gas oil fractions, and pretreat them prior to cracking in order that the catalytic cracking operation becomes more desirable overall, even though by reducing the amount of cracking feed per barrel of crude oil the yield of gasoline is thereby reduced. For example, it has been proposed heretofore to hydrotreat various heavy metal-containing hydrocarbon oils prior to charging them, or a fraction thereof, to a catalytic cracking operation. By so doing the hydrocarbon may be given an improved hydrogen-to-carbon ratio and the amount of contaminants, such as cokeformers, sulfur and nitrogen may be reduced. The content of metals which poison cracking catalysts is also reduced; these metals deposit on the hydrotreating catalyst. Removal of any substantial amount of these contaminants from the cracking feed tends to enhance efficiency of the catalytic cracking operation. The degree of feedstock improvement from hydrogenation is dependent, however, upon several factors which include the severity of reaction, hydrogen'consumption and the activity decline of the hydrogenation catalyst during use. A fresh hydrogenation catalyst, for example, may remove the bulk of the metal contaminants even in low severity operations, but as the catalyst activity decreases metals removal will decline at a given severity. Increased severities can then be employed to maintain the desired extent of metals removal as the operation progresses. Increased severity,

however, may involve a greater consumption of hydrogen, a larger capital investment for high-pressure equipment, and a reduction in the yield of cracking feedstock. Alternatively, the effect of metal contaminants in reducing the activity of the hydrogenation catalyst can be overcome by discarding this catalyst more often. Either of these alternatives increases the cost of hydrogenation substantially.

In the process of this invention the crude oil or other residual oil containing a large amount of catalyst poisoning metals is subjected to catalytic cracking and the poisoned catalyst is demetallized. No special theoretical limit exists on the amount of metal contaminants which may be fed to the cracking, it being possible to keep the metals level on the catalyst at a desirably low level by a high demetallization rate and by increasing the amount of catalyst in the reactor (the inventory) per barrel of feed processed, over the conventional inventory. In operation, however, the overall feedstock to the cracking operation usually contains no more than about 300 p.p.m. of the poisons nickel and vanadium, measured as their common oxides, preferably no more than about l to 30 p.p.m. NiO and V205. The preferable maximum content of NiO is about 10 p.p.m. and of V205 about 20 p.p.m. The desired maximum of metals content may be brought about by blending a crude or other residual oil with relatively metals free stocks, including cycle oil, to reduce the proportion of metal in the entire feed. Generally, the feed to the catalytic cracking will contain more than about 0.3 p.p.m. NiO and/or 0.5 p.p.m. V205, preferably more than about 1 p.p.m. NiO and/or 2 p.p.m. V205 to economically justify the provisions made for cracking catalyst demetallization.

Gas oil components in the catalytic cracker effluent are excellent feeds to cracking and although these recycle stocks are substantially free of metal contaminants, hydrogenation of the recycle stocks before return to the cracking operation has been found highly advantageous in improving the cracking operation as a whole. Preferably, therefore, the catalytic cracking feedstock will contain at least about 10%, preferably about S0-65% of hydrogenated recycle stock. By making it possible to feed even the entire desalted crude oil to the catalytic cracking, the process of this invention can eliminate the preliminary fractionation or topping customarily given to crude oil.

The invention will be better understood by reference t0 the accompanying drawing which represents apparatus leaves a residual fraction containing the gas oil components for passage by

lines

22, 25, 27 and 30 to the reactor 33. Steam may be added to the feed by line 35. Alternatively, the desalted crude may by-pass the fractionator, travelling by

lines

10, 36, 25, 27 and 30 to the reactor 33, which preferably is arranged for holding a bed of catalyst in fluidized form. The cracking rfeed carries through line 30 to theV reactor 33 regenerated catalyst from the regenerator standpipe 39. Cracked products from the cracking reactor 33 pass, by liner42, to the fractionator 44 in which fixed gases and gasoline vapors are separated out and from which these products are conducted by line 46 to the condenser 48. In the condenser the gasoline fraction is converted to liquid form and removed by line 50. Gases leave by line 53. 2 ,Y

The cycle oil fraction which may comprise all or part 'of the cracker Yeffluent components which boil primarily over about 400 F., may be conducted back to the reactor 33 by

lines

54 and 30. Alternatively, the cycle oil may be brought to the hydrogenation reactor 55 by line 57 from the fractionator 44. Line A'conducts hydrogen to the hydrogenation reactor, which contains a'hydrogenation catalyst. The hydrogenator effluent is conducted by line 63 to fractionator 66 where gas and gasoline components formed in the hydrogenation are removed. The fractionator 66 also separates hydrogenated light cycle oil from heavy cycle oil. The heavy hydrogenated cycle oil may be removed by line 69 for passage to line 35 by way of line 71 which may be provided with the pump 73. Light cycle oil is drawn from fractionator 66 by line 73 whence it may be removed from the system or recycled to the cracking by line 75.

Cracking catalyst is preferably continuously withdrawn from the reactor 33 for passage to the regenerator 77 through thestandpipe and line 82. Air is supplied to line 82 from source 84, for conduction of the catalyst and combustion of the coke in the regenerator. A sidestream of regenerated catalyst is preferably continuously withdrawn from the regenerator standpipe 39 for passage by line 86 to the demetallization unit 88 which is a series of vessels designed to accomplish the demetallization techniques outlined below. Demetallized catalyst returns to the regenerator by line 90.

The feed to the cracking zone is heated to a temperature sufficient to vaporize most of feed without substantial cracking. The feed in vapor or mixed vapor and liquid form or in the form of vapor having entrained liquid is cat'alytically treated under more or less conventional fluid catalytic cracking conditions. As mentioned, the charge stocks are of the residual type, containing upwards of 0.3 p.p.m. NiO and/or 0.5 p.p.m. V205, which are generally avoided in catalytic cracking. A mixture of vanadium and nickel may be considered as harmful as a single metal even though the individual amounts of each metal are below the values mentioned above because the effect of the total amount of the metallic components is frequently sufficient to give harmful effects during catalytic cracking.

A Cracking conditions generally include a temperature of about 750 to 1000"' F., preferably about 850 to 975 F. and a pressure between atmospheric and p.s.i.g., preferably about 5 and 25 p.s.i.g., advantageously Without substantial addition of free hydrogen to the system, and a weight hourly space velocity from about 0.01 to 10, depending on the poisoning metal content of the feed, to obtain about a 40-70 volume percent, preferably about 50 to 60%, conversion of the 400Y F. to 900 F. Vgas oil components of the feed to gasoline and other desired lighter components. Subjecting the entire crude to catalytic cracking will result in the production of considerably more butylenes and isobutane than obtained from conventional refining procedures. These compounds are `valuable as feed to an alkylation unit where they are used inthe production of high octane gasoline. The gasoline fraction boiling in the range of labout 1GO-400 F. will also undergo a significant amount of cracking. The net result of this cracking will be to increase the octane number of the cut. Therefore, this light ends fraction may not require the usual processing by catalytic reform- Ving to produce gasoline of satisfactory octane number. In the cracking operation a batch, semi-continuous or continuous system is used but most often it is the latter.

VThe .cracking catalyst is of the solid refractory metal oxide type known in the art, for instance silica, alumina, magnesia,titania, etc., or their mixtures. Of most importance are the synthetic gel-containing catalysts, such as the synthetic and the semi-synthetic, i.e., synthetic gel supported on a carrier such as natural clay, cracking catalysts. 'Ihe cracking catalysts which have received the widest acceptance today are usually predominantly silica, `that is silica-based, and may contain solid acidic oxide promoters, e.g., alumina, magnesia, etc., with the promoters usually being less than about 35% of the catalyst, preferablyV about 5 to 25%. These compositions are calcined to astate of very slight hydration. The cracking catalyst can be of macrosize, for instance bead form orfinelyqdivided form, and employed as a fixed, moving or fiuidized bed as noted with respect to the hydrotreating catalyst. In a highly preferred form of this invention finely divided (fiuidized) catalyst, for instance having particles predominantly in the 20 to 150 micron range, is disposed as a fluidized bed in the reaction zone to which the feed is charged continuously and is reacted essentially in the vapor phase.

In cracking, coke yield may be held to a minimum through the use of good steam stripping and a high steam partial pressure, and removal of coke from the catalyst is performed by regeneration. Regeneration of a catalyst to remove carbon is a relatively quick procedure in most commercial catalytic conversion operations. For example, in a typical fluidized cracking unit, a portion of catalyst is continually being removed from the reactor and sent to the regenerator for contact with air at about 950 to 1200 F., more usually about 1000 to 1150 F. Combustion of coke from the catalyst is rapid, and for reasons of economy only enough air is used to supply the needed oxygen. Average residence time for a portion of catalyst in a conventional regenerator may be on `the order of about six minutes and the oxygen content of the eiuent' gases from the regenerator is desirably less than about 1/z%. In this invention, however, coke laydown may be generally somewhat higher and the residence time for catalyst in the regenerator is therefore longer. The regeneration of any particular quantum of catalyst is generally regulated to give a carbon content of less than about 5.0%, generally less than about 0.5%. Regeneration puts the catalyst in a substantially carbon-free state, that is, the state where little, if any, carbon is burned or oxygen consumed even when the catalyst is contacted with oxygen at temperatures conducive to combustion. The regeneration does not remove from the catalyst the metals deposited from the cracking feed, which metals accumulate on the catalyst during the cracking operation.

The ellluent from the cracker conveniently is distilled to isolate the gasoline fraction. Also, products such as fixed gases, boiling below the gasoline range, are removed from the system. The use of residuals as a cracking feed produces a 400 F.+ fraction which is greater in volume than that produced in conventional processes. In addition, the 400 F.} fraction contains a large amount of heavy gas oil, that is, gas oil boiling above about 600 F. which frequently cont'ains a good deal of sulfur and oleiins, although substantially free of metals. In this invention, this heavy gas oil, as well as the light gas oils recycled to the cracking reactor in conventional processing is conveniently recycled to cracking preferably after hydrotreating to improve its cracking characteristics. All or part of the 400 F.-1 fraction may be recycled with or without hydrogenation. Conveniently, the fractionation of the cracker eluent is conducted in such a manner as to leave the entire higher boiling fraction behind as a bottoms which may be hydrotreated and recycled for blending with virgin feed to the cracking Zone.

In the hydrotreating operation the cycle oil is contacted with a catalyst in the presence of free hydrogen under superatrnospheric pressure. The hydrogenation catalysts generally known in the art can be employed. Calcined solid hydrogenation catalyst are preferred and they are usually disposed as a xed bed of macrosized partices, say of about 1/8 to 1A in diameter and about 1/8 to l" or more in length. A moving bed of macrosized catalyst or a iluidized bed of finely divided particles can also be used. The catalyst contains catalytically active amounts of a hydrogenation promoting metal, for instance a heavy metal component such as those of metals having atomic numbers of about 23 to 28, the Group VH1 catalysts of the platinum and iron groups, molybdenum, tungsten and combinations thereof. Frequently the metals are disposed as inorganic components, for instance oxides, sulides or other compounds, supported on a solid carrier exempliiied by alumina, silica, etc. Advantageously, the

catalyst contains a Combination of metals of the iron group with vanadium or a metal of Group Vlb of the periodic chart having atomic numbers from 42 to 74, i.e., molybdenum and tungsten. A commercial catalyst contains cobalt and molybdenum, eg., cobalt molybdate, supported on alumina. The amount of catalytically active metal in the supported catalysts is usually about 1 to 30 weight percent of the catalyst and preferably about 3 to 20 weight percent, with there being at least about 1%, preferably at least about 2% of each catalytically active metal when combinations are used. In the hydrotreating operation there is generally deposition of coke on the catalyst which can be removed by continuous or intermittent regeneration, that is, combustion of the coke by contact with oxygen.

In hydrotreating, the hydrogen has a number of effects on the cycle stock. Hydrotreating removes sulfur and may frequently serve to saturate components of the feed which are susceptible to such. Thus hydrotreating may increase the hydrogen-to-carbon ratio of the cycle oil, reducing the coke-forming tendencies of the feedstock. As mentioned, this effect is highly desirable considering the heaviness of the cycle oils derived from the residuals fed to catalytic cracking in this invention. Also, hydrotreating generally causes a certain amount of the feed to be converted (cracked) to lower boiling materials. Extensive cracking in the hydrotreating operation is usually desirable for the relatively heavy cycle oils of the process of this invention. The extent of cracking which takes place during hydrotreating is determined to a great extent by the temperature employed, higher temperatures in general giving a greater degree of cracking. Effective hydrotreating for saturation of hydrocarbons may be achieved at a relatively low temperature when high pressures are used. If, however, the costs of the process are to be restricted and equipment suitable for withstanding less pressure used, a temperature suiciently high to achieve the main objects of hydrotreating at lower pressure may be dictated. In such an eventuality, gas or gasoline produced by hydrocracking and unsuitable for use as feed to the catalytic cracking may be used as reformer feedstock.

With these factors in mind, the conditions of the hydrotreating operation may be chosen in view of the type of operation contemplated. Conditions are selected to give the desired hydrogen consumption and cycle oil improvement. In general, however, an elevated temperature such as about 600 to 900 F. will be employed and the pressure will be superatmospheric usually falling in the range of about 300 to 3000 p.s.i.g. Free or molecular hydrogen is provided in the operation and generally in an amount of about 50 to 20,000 standard cubic feet per barrel of hydrocarbon oil feedstock, while the space velocity will lie in the area of about 0.1 to 10 or more WHSV (weight of hydrocarbon feedstock per hour per weight of catalyst).

In the hydrotreating operation there is always at least a minor amount of hydrogen which is consumed by chemical combination with a component of the hydrocarbon feed, and the extent of this consumption may depend upon the type of operation effected; that is, saturation or cracking. With the cycle oils hydrogenated in this invention, greater than about 10% of the hydrocarbon charge is usually converted to lower boiling normally liquid materials and quite frequently at least about 25% is converted, and although the hydrogen consumption is high, the catalyst demetallization features of this invention serve to minimize dehydrogenation in the later catalytic cracking, assuring the transmission of the hydrogen into the desired end products. Conversion to lower boiling materials rarely exceeds of the cycle oil charged to hydrotreating. Generally the entire residual fraction of the hydrotreater etlluent is returned to the cracking zone.

Hydrogen is supplied to the hydrotreating operation in amounts of from about 50 to 20,000 standard cubic feet per barrel of feed. Hydrogen consumption is usually at least about 70-300 standard cubic feet of hydrogen per barrel of hydrocarbon oil feed. In this invention, the' hy drogen consumption is often in the range of about 1000- to 2000 or more standard cubic feet per barrel. The conditions of hydrotreating are generally a temperature of about 600-900 F., a superatmospheric pressure of about 300-3000 lbs. and a WHSV of about 0.1 to 10, most often a temperature of about 750 to 900 F., a pressure over about 1000 p.s.i.g., preferably about 1500 to 2500 p.s.i.g., and about 100 to 10,000 standard cubic feet of hydrogen per barrel.

As pointed out, regeneration of the cracking catalyst does not remove poisoning metals deposited on the catalyst, and unless steps are taken to prevent excess accumulation, excessive dehydrogenation and coking take place in the cracking, partially undoing the Work performed in hydrotreating the cycle oil and severely reducing the yield of gasoline in the cracker eluent.

In the practice of this invention, catalyst demetallization is accomplished by the intermittent or continuous Withdrawal of contaminated catalyst from the cracking system, for example, from the regenerator staudpipe. The catalyst is subjected to one or more of the demetallization procedures described hereinafter and then the catalyst, substantially reduced in contaminating metal content, is returned to the cracking system. In the treatment to take poisoning metals from the cracking catalyst the amount of metal is removed which is necessary to keep the average metal content of the catalyst in the cracking system below the units tolerance for poison. The tolerance of the cracker for poison in turn determines to large extent the amount of metals removed in the catalyst demetallization procedure. Where the catalyst contains a greater amount of poisoning metal, a particular treatment may remove a greater amount of metal for example, if the cracker can tolerate an average of 100 p.p.m. Ni and the demetallization process can remove 50% of the nickel content of the catalyst, only 50 p.p.m. of nickel can be removed in a pass through the catalyst demetallization system. However, where the cracker can tolerate 500 p.p.m. of nickel, it may be possible to remove 250 p.p.m. nickel from the catalyst with each pass through the demetallization system. It is advisable, therefore, to operate the cracking and demetallization procedures with a catalyst having a metals content near the limit of tolerance of the cracker for poisoning metals. This tolerance for poisoning metal oxide is seldom greater than about 5000 p.p.m. Catalyst demetallization is not economically justied unless the catalyst contains at least about 50 p.p.m. nickel and/or 50 p.p.m. vanadium. Preferably the equilibrium metals level is allowed to exceed about 200 p.p.m. nickel and/ or 500 p.p.m. vanadium so that total metals removal will be greater per pass through the demetallizer.

In the treatment to take poisoning metals from the cracking catalyst a large or small amount of metal can be removed as desired. The demetallization treatment generally removes about 10 to 90% of one or more poisoning metals from a catalyst portion which passes through the treatment. Advantageously a demetallization system is used which removes about 60 to 90% nickel and 20 to 40% vanadium from the treated portion of catalyst. Preferably at least 50% of the equilibrium nickel content and of the equilibrium vanadium content is removed. The actual time or extent of treating depends on various factors, and is controlled by the operator according to the situation he faces, e.g., the extent of metals content in the feed, the level of conversion unit tolerance for poison, the sensitivity of the particular catalyst toward a particular phase of the demetallization procedure, etc. Also, the thoroughness of treatment of any quantum of catalyst in commercial practice is balanced against the demetallization rate chosen; that is, the amount of catalyst, as compared to the total catalyst in the conversion system proper, which issubjected to the demetallization treatment per unit of time. A high rate of catalyst Withdrawal from the conversion system and quick passage through a mild demetallization procedure maysuice as readily as amore intensive demetallization at a slower rate to keep the total of'poisoning metal in the conversion reactor Within the tolerance of the unit for'poison.V In a continuous operation of the commercial type a satisfactory treating rate may be about- 50 to 150% of the total catalyst inventory in the system, per twenty-four hour day of operation, although othertreating rates may be used. With a continuously circulating catalyst stream, such as in the ordinary tluid system a slip-stream of catalyst, at the equilibriumlevel of poisoning metals may be removed intermittently or continuously from the regenerator standpipe of the cracking system. The catalyst is subjected to one or more of the demetallization procedures described hereinafter and then the catalyst, substantially reduced in contaminating metal content, is returned to the cracking system.

The demetallization of the catalyst will generally include one or more processing steps. Copending patent applications Serial Nos. 758,681, led September 3, 1958, W abandoned; 763,833 and 763,834, led September 29, 1958, now abandoned; 767,794, filed October 17, 1958, now abandoned; 842,618, filed September 28, 1959, now abandoned; 849,119, filed October 28, 1959, now Patent No. 3,094,059; 19,313, led April 1, 1960, now abandoned; 39,810, tiled .lune 30, 1960; 47,598, led August 4, 1960, now abandoned; 53,380, tiled September 1, 1960, now Patent No. 1,122,497; 53,623, iiled September 2, 1960; 54,368, now Patent No. 1,122,512; 54,405, now Patent No. 3,122,510 and 54,532, now abandoned, led September 7, 1960; 55,129; 55,160 and 55,184, led September 12, 1960; 55,703, tiled September 13, 1960; 55,838, filed September 14, 1960, now abandoned; 67,518, led November 7, 1960; 73,917, filed December 2, 1960 and 81,256 and 81,257, led January 9, 1961, now abandoned; all of Which are hereby incorporated by reference, described procedures by which vanadium and other poisoning'm'etal's included in a solid oxide hydrocarbon conversion catalyst are removed by dissolving them from the 'catalyst or subjecting thel catalyst, outside the hydrocarbon conversion system, to elevated temperature conditions and vapors which react with the metal contaminants to put them eventually into the chloride, sulfate or other volatile, water-dispersible or more available form. A significant advantage of these processes lies in the fact that the over- 'all metals removal operation, even if repeated, does not 'unduly deleteriously affect the activity, selectivity, pore tructure and other desirable characteristics of the catayst.

Treatment of the regenerated catalyst with molecular oxygen-containing gas is employed to improve the removal of vanadium from the poisoned catalyst. This treatment is described in Copending application Serial No. 19,313 'and is preferably performed at a temperature at least about 50 VF. higher than'the regeneration temperature, that is, Vthe average temperature at which the major portion of carbon is` removed from the catalyst. The temperature of 'treatment With molecular oxygen-containing gas will generally be in the range of about 1000 to 1800o F. but below a temperature Where the, catalyst undergoes any substantial deleterious change in its physical or chemical characteristics, preferably a temperature of about 1150 to v1350 F. or even as high as 1600 F. The duration of the oxygen treatment and the amount of vanadium prepared by the treatment for subsequent removal is dependent upon the temperature and the characteristics of the equipment used. If any significant amount of carbon is present in the catalyst at the start of this high-temperature treatment, the essential oxygen contact is that continued after carbon removal, which may vary from the short time necessary to produce an observable effect in the later treatment, say, a quarter of an hour, to a time just long enough not to damage the catalyst. In any event,

after carbon removal, the oxygen treatment of the essentially carbon-free catalyst is at least long enough to stacarbon or sulfur.

atea-160 bilize a substantial amount of vanadium in its highest valence state at the surface of the catalyst, as evidenced by a significant increase, say at least about preferably at least about 100%, in the vanadium removal in subsequent stages of the process. This increase is over and above that which would have been obtained by the other metals removal steps without the oxygen treatment. The maximum practical time of treatment Will vary from about 4 to 24 hours, depending on the type of equipment used. The oxygen-containing gas used in the treatment contains molecular oxygen as the essential active ingredient and there is little significant consumption of oxygen in the treatment. The gas may be oxygen, or a mixture of oxygen with inert gas, such as air or oxygen-enriched air, containing at least about 1%, preferably at least about 10% O2.V The partial pressure of oxygen in the treating gas may range widely, for example, from about 0.1 to 30 atmospheres, but rarely will the total gas pressure exceed about 25 atmospheres.

The catalyst may pass directly from the oxygen treatment to a vanadium removal treatment especially Where this is the only important contaminant, as may be the case when a feed is derived, for example, from Venezuelan crude. Such treatment may be a basic aqueous Wash such as described in copending patent applications Serial No. 767,794 and Serial No. 39,810. Alternatively vanadium may be removed by a chlorination procedure as described in copending application Serial No. 849,199.

Vanadium may be removed from the catalyst after the high temperature treatment with molecular oxygen-containing gas by Washing it with a basic aqueous solution. The .pl-I is frequently greater than about 7.5 and preferably the solution contains ammonium ions which may be inthe form of NH4+ ions or organic-substituted NH4+ ions such as methyl ammonium and quaternary hydrocarbon radical ammoniums. The amount of ammonium ion in the solution is suiiicient to give the desired vanadium removal and will often be in the range of about 1 to 25 or more pounds per ton of catalyst treated. The temperature of the Wash solution may vary within Wide limits: room temperature or below, or higher. Temperatures above 215 F. require pressurized equipment, the cost of which does not appear to be justified. Very short contact times, for example, about a minute, are satisfactory, while the time of washing may last 2 to 5 hours or longer. After the ammonium wash the catalyst slurry can be liltered to give a cake which may be reslurried with Water or rinsed in other ways, such as, for example, by a Water wash on the iilter, and the rinsing may be repeated, if desired, several times.

Alternatively, after the high temperature treatment With oxygen-containing gas, treatment of a metals contaminated catalyst with a chlorinating agent at a moderately elevated temperature up to about 1000 F. is of value in removing vanadium contaminants from the catalyst as volatile chlorides. This treatment is described in copending application Serial No. 849,199. The chlorination takes place at a temperature of at least about 300 F., preferably about 550 to 650 F. with optimum results usually being obtained near 600 F. The chlorinating agent is essentially anhydrous, that is, if changed to the liquid state no separate aqueous phase would be observed in the reagent.

The chlorinating reagent is a vapor which contains chlorine or sometimes HCl, preferably in combination with Such reagents include molecular chlorine but preferably are mixtures of chlorine With, for example, a chlorine-substituted light hydrocarbon, such as carbon tetrachloride, which may be used as such or formed in situ by the use of, for example, a vaporous mixture of chlorine gas with low molecular Weight hydrocarbons such as methane, n-pentane, etc. About 1-40 percent active chlorinating agent based on the Weight of the catalyst is generally used. The carbon or sulfur compound promoter is generally used in the `amount of about 1-5 or 10 percent or more, preferably about 2-3 percent, based on the weight of the catalyst for good metals removal; however, even if less than this amount is used, a Considerable improvement in metals conversion is obtained over that which is possible at the same temperature using chlorine alone. The chlorine and promoter may be supplied individually or as a mixture to a poisoned catalyst. Such a mixture may contain about 0.1 to 50 parts chlorine per part of promoter, preferably about 1-10 parts per part of promoter. A chlorinating gas comprising about 1-30 weight percent chlorine, based on the catalyst together with one percent or more S2Cl2 gives good results. Preferably, such a gas provides 1-10 percent Cl2 and about 1.5 percent S2Cl2, based on the catalyst. A saturated mixture of CCL, and C12 or HC1 can be made by bubbling chlorine or hydrogen chloride gas at room temperature through a vessel containing CC14; such a mixture generally contains about 1 part CC14/ 5-10 parts C12 or HCl. Conveniently, a pressure of about 0-100 or more p.s.i.g., preferably about 0-15 p.s.i.g. may be maintained in chlorination. The chlorination may take about 5 to minutes, more usually about 20 to 60 minutes, but shorter or longer reaction periods may be possible or needed, for instance, depending on the linear velocity of the chlorinating and purging vapors.

The demetallization procedure employed in this invention may be directed toward nickel removal from the catalyst, generally in conjunction with vanadium removal. Nickel removal may be accomplished by removing nickel compounds directly from the catalyst for example, by dissolving, and/or by converting the nickel compounds to volatile materials such as nickel carbonyl and/ or materials soluble or dispersible in an aqueous medium, e.g., water or dilute acid. The Water-dispersible form may be one which decomposes in Water to produce Water-soluble products. The removal procedure for the converted metal may be based on the form to which the metal is converted. The mechanism of the washing steps may be one of simultaneous conversion of nickel and/ or vanadium to removable form and removal by the aqueous Wash; however, this invention is not to be limited by such a theory.

Conversion of some of the metal poisons, especially nickel, to the sulfate or other Water-dispersible form as described in copending application Serial No. 758,681 comprises subjecting the catalyst to a sulfating gas, that is SO2, S03 or a mixture of SO2 and O2, at an elevated temperature. Sulfur oxide contact is usually performed at a temperature of about 500 to 1200 F. and frequently it is advantageous to include some free oxygen in the treating gas. Another procedure described in copending applications Serial No. 763,834 and Serial No. 842,618 includes suliiding the catalyst and performing an oxidation process, after which metal contaminants in Water-dispersible form, preferably prior to an ammonium Wash may be removed from the catalyst by an aqueous medium.

The sulding step can be performed by contacting the catalyst with elemental sulfur vapors, or more conveniently by contacting the poisoned catalyst with a volatile sulfide, such as H28, CS2, or a mercaptan. The contact with the sulfur-containing vapor can be performed at an elevated temperature generally in the range of about 500 to 1500 F., preferably about 800 to l300 F. Other treating conditions can include a sulfur-containing vapor parital pressure of about 0.1 to 30 atmospheres or more, preferably about 0.5 to 25 atmospheres. Hydrogen sulde is the preferred sulding agent. Pressures below atmospheric can be obtained either by using a partial vacuum or by diluting the vapor with gas such as nitrogen or hydrogen. The time of contact may vary on the basis of the temperature and pressure chosen and other factors such as the amount of metal to be removed. The sulfiding may run for, say, up to about 20 hours or more depending on these conditions and the severity of the poisoning. Temperatures of about 900 to 1200 F. and pressures approximating 1 atmosphere or less seem near optimum for sulding and this treatment often continues for at least 1 or 2 hours but the time, of course, can depend upon the manner of contacting the catalyst and sulding agent and the nature of the treating system, eg., batch or continuous, as well as the rate of diffusion within the catalyst matrix. The sulfiding step performs the functions not only of supplying a sulfur-containing metal compound which may be easily converted to a water-dispersible form but also appears to concentrate some metal poisons, especially nickel, at the surface of the catalyst particle.

Oxidation after sulfiding may be performed by a gaseous oxidizing agent to provide metal poisons in a dispersi- 'ble form. Gaseous oxygen, or mixtures of gaseous oxygen with inert gases such as nitrogen, may be brought into contact with the sulfided catalyst -at an oxygen partial pressure of about 0.2 atmospheres and upward, temperatures upward of room temperature and usually not above about 1300 F., and times dependent on temperature and oxygen partial pressure. Gaseous oxidation is best carried out near 900 F., about one atmosphere O2 and at very brief contact times.

The metal sulfide may be rendered water-dispersible by a liquid aqueous oxidizing agent such as a dilute hydrogen peroxide or hypochlorous acid water solution, as described in copending application Serial No. 842,618. The inclusion inthe liquid aqueous oxidizing solution of sulfuric acid or nitric acid has been found greatly to reduce the consumption of peroxide. In addition, the inclusion of nitric acid in the oxidizing solution provides for increased vanadium removal. Useful proportions of acid to peroxide to catalyst generally include about 2 to 25 pounds acid (on `a 100% basis) to about 1 to 30 pounds or more H2O2 (also on a 100% basis) in a very dilute aqueous solution, to about one ton of catalyst. A 30% H2O2 solution in water seems to be an advantageous raw material for preparing the aqueous oxidizing solution. Sodium peroxide or potassium peroxide may be used in place of hydrogen peroxide and in such circumstances, extra sulfuric or nitric acid may be used.

Another highly advantageous oxidizing medium is an aerated dilute nitric acid solution in water. Such a solution may be provided by continuously bubbling air into a slurry of the catalyst in very dilute nitric acid. Other .oxygen-containing gases may be substituted for air. Varying oxygen partial pressure in the range of about 0.2 to 1.0 atmosphere appears to have no effect in time required for oxidation, which is generally at least about 7 to 8 minutes. The oxidizing slurry may contain about 20% solids and provide about five pounds of nitric acid per ton of catalyst. Studies have shown a greater concentration of HNO3 to be of no significant advantage. Other oxidizing agents, such as chromic acid where a small residual Cr2O3 content in the catalyst is not significant, and similar aqueous oxidizing solutions such as water solutions of manganates and permanganates, chlorites, chlorates and perchlorates, bromites, bromates and perbromates, iodites, iodates and periodates, are also useful. Bromine or iodine water, or aerated, ozonated or oxygenated water, with or without acid, also will provide a dispersible form. The conditions or oxidation can be selected as desired. The temperature can conveniently range up to about 220 F. with temperatures of above about 150F. being preferred. Temperaures above about 220 F. necessitated the use of superatmospheric pressures and no need for such has been found.

After conversion of nickel sulfide to a dspersible form, the catalyst is Washed with an aqueous medium to remove the metal compound. This aqueous medium, for best removal of nickel, is generally somewhat acidic, and this condition may be brought about, at least initially, by the presence of an acid-acting salt or some entrained acidic oxidizing agent on the catalyst. The aqueous medium can contain extraneous ingredients in trace amounts, so long as the medium is essentially water and the extraneous ingredientsdo not interfere with demetallization or adversely affect the properties of the catalyst. Ambient temperatures can be used in the wash but temperatures of about 150 F. to the boiling point of water are sometimes helpful. Pressures above atmospheric may be used but the results usually do not justify the additional equipment. Where an yaqueous oxidizing solution is used, the solution may perform part or all of the metal compound removal simultaneously with the oxidation. In order to avoid undue solution of alumina from a chlorinated catalyst, contact time in this stage is preferably held to about 3 to 5 minutes which is sufficient for nickel removal. Also, since a slightly acidic solution is desirable for nickel removal, this wash preferably takes place before the arnmonium wash.

Alternative to the removal of poisoning met-als by procedures involving contact of the sulfided or sulfated catalyst with aqueous media, nickel poison may be removed through conversion of the nickel sulfide to the volatile nickel carbonyl by treatment with carbon monoxide, as described in copending application Serial No. 47,598. In such a procedure the catalyst is treated with hydrogen at an elevated temperature during which nickel contaminant is reduced to the elemental state, then treated, preferably under elevated pressure and at a lower temperature, with carbon monoxide, during which nickel carbonyl is formed and flushed off the catalyst surface. Hydroygenation takes place at a temperature of about 800 to 1600 F., at a pressure from atmospheric or les-s up to labout 1000 p.s.i.g. with a vapor containing 10 to 100% hydrogen. Preferred conditions are a pressure up to about 15 p.s.i.g. and a temperature of about 1100 to l300 F. and a hydrogen content greater than about 8O mole percent. The hydrogenation is continued until surface accumulations of poisoning metals, particularly nickel, are substantially reduced to the elemental state. Carbonylation takes place at a temperature substantially lower Vthan the hydrogenation, from about ambient temperature to `300" F. maximum and at a pressure up to about 2000 p.s.i.g., with fa gas containing about 50-100 mole percent CO. Preferred conditions include greater than about mole percent CO, a pressure of up to `about 800 p.s.i.g. and a temperature of about 1GO-180 F. The CO treat- `ment serves generally both to convert the elemental metals,

especially nickel, to volatile carbonyl and to remove the carbonyl. v

After the -ammonium wash, yor after the final treatment which may be used in the catalyst demetallization procedure, the catalyst is conducted back to the cracking system. Where a small amount of the catalyst inventory is demetallized, the catalyst may be returned to the cracking system, preferably to the regenerator standpipe, als a slurry in its final aqueous treating medium. Where a large amount of catalyst inventory is treated, lest the water put out the fire or unduly lower the temperature in the regenerator, it may be desirable first to dry a Wet -catalyst filter cake or filter cake slurry at say about 250 to 450 F. and also, prior to reusing the catalyst in the cracking operation it can be calcined, say at temperatures usually in the range of about 700 to 1300 F. Prolonged calcination of the catalyst at above about 1100 F. may sometimes be disadvantageous. Calcination removes free water, if any is present, land perhaps some but not all of the combined Water, and leaves the catalyst in an active state without undue sintering of its surface. Inert gases such as nitrogen frequently may be employed after contact with reactive vapors to remove any of these vapors Yentrained in the catalyst or to purge the catalyst of reaction products.

The demetallization procedure of this invention has been found to be highly successful whenrused in conjunction with fluidized catalyticcracking systems to controlY the amount of metal poisons on the catalyst. When such catalysts are processed, a fluidized solids technique is recommended for these vapor contact demetallization procedures as a way to shorten the time requirements. Any given step in the demetallization treatment is usually continued for a time sufficient to effect a substantial conversion or removal of poisoning metal and ultimately results in a substantial increase in metals removal compared with that which would have been removed if the particular step had not been performed. After the available catalytically active poisoning metal has been removed, in any removal procedure, further reaction time may have relatively little effect on the catalytic activity of the depoisoned catalyst, although further metals content may be removed by repeated or other treatments.

The following examples are illustrative of the process of this invention but should not be considered limiting.

A South Texas crude petroleum containing about 3 ppm. nickel, measured as NiO, and 5 ppm. vanadium measured as V205, is processed in a catalytic cracking unit having a catalyst inventory of 300 tons and providing for a feed of 30,000 barrels/day. The feedstock analyzes 41% materials boiling over 400 F. `and this residual fraction analyzes as 7.2 p.p.m. Ni() and 12.2 ppm. V205. This whole crude feedstock is sent to the catalytic cracking unit having a synthetic silica-alumina gel catalyst containing about 25% alumina and the remainder essentially silica, wherein a temperature of about S75-975 F. is maintained under a pressure of about 5 to 35 p.s.i.g., at a weight hourly space Velocity of about 0.25 to 25, and a catalyst-to-oil ratio of about 5 to 15. The resulting cracked products are fractionated to recover a fixed gas fraction, distillate gasoline of 430 F. end boiling point, and cycle oils having a boiling point of over 430 F.

The catalyst is continually sent to a regenerator where it is contacted With air at a temperature of about 1000- 1200 F. to burn off the carbon. The catalyst, whose carbon content is reduced from 1.5% to 0.36%, analyzes about 440 p.p.m. nickel oxide and 1690 ppm. vanadium pentoxide. Catalyst is continuously removed from the regeneratorrat a `daily inventory rate of about 125 tons and sent to a zone where it is held for about an hour in contact with air at about l300 F. and then sent to a sulliding zone where it is fluidized with HZS gas at a temperature of about 1100 F. for about an hour. The catalyst is then cooled and purged with inert gas and chlorinated in a chlorination zone with an equimolar mixture of C12 and CCI., at about 600 F. After about an hour no trace of vanadium chloride can be found in the chlorination effluent and the catalyst is quickly washed with water. A pH of about 2.5 is imparted to this wash medium by chlorine entrained in the catalyst and the wash serves to remove nickel chloride.

The catalyst, substantially reduced in nickel and vanadium content, is filtered from the wash slurry, dried at about 350 F. Iand returned to the regenerator. The treated catalyst is analyzed and shows a reduction of about 60% in nickel and about 25% vanadium. This metals reduction, along with about three tons per day of catalyst lost as fines and replaced with fresh catalyst, is sufficient to maintain the metals level at about 440 ppm. NiO on the catalyst. It is estimated that without demetallization 36 tons per day of catalyst would have to be replaced to maintain a level of 940 ppm. nickel.

In another run, a 40% reduced West Texas petroleum crude having an API gravity of about 15.1, a Conradson carbon of about 8.8 Weight percent, and an initial boiling point above about 650 F at atmospheric pressure, containing 32 ppm. of nickel and 75 ppm. of vanadium is preheated and introduced at the rate of 10,000 bbl./ day into a catalytic cracker, mixed with a finely divided cracking catalyst, about 100 lbs. of steam per barrel of residual feed for dispersion and strippiugand about 9,000 bbl./ day of a heavy cycle oil. The cycle oil boiling above about 400 F. is obtained from the cracker effluent, is substantially free of metals and has an API Igravity of 11.2, an 'aniline point of 100 F. and contains about 0.6% sulfur. The catalyst introduced into the feed line is a Nalcat synthetic gel cracking catalyst containing 25% A1203, the balance silica, and having fiuidizable particle size. The catalyst inventory is 158 tons. Cracking temperature is 900 F. The cracking system is provided with a demetallization unit designed to hold a slipstream of regenerated catalyst for about 2 hours in a zone where it is contacted with air at about 1300 F. and then in a sulfiding zone to be fluidized with HZS gas at a temperature of about 1050 F. for about 2 hours. Water containing dilute hydrogen peroxide mixed with nitric acid is designed to be brought in contact with the sulfided catalyst for 10 minutes Iat a temperature of 200 F. The catalyst is then to be washed with an ammonium hydroxide solution having a pH of about 8 to 11, to remove the available vanadium. By treating 165 tons of catalyst per day in this demetallization system and replacing 1.6 tons of catalyst per day lost from the system as fines, it is found that lan equilibrium level of 500 ppm. NiOon the catalyst may be maintained. It is estimated that without catalyst `demetallization tons of catalyst per day would have to be replaced to maintain this metals level.

The cracked products from the cracking zone are introduced into a fractionator where the products are separated into gas fractions and a gasoline fraction having an end boiling point of about 400 F. which is recovered, and the cycle oil fraction boiling above about-400 F. The yields are as follows:

Gasoline, vol, percent 42.6 Butane, vol. percent 13.1 Dry gas, wt. percent 11.6 Coke, wt. percent 8.5 Gas oil, vol. percent 37.0

This operation is shut down and the recycle line is connected to a hydrogenation unit provided with a chromium-molybdenum-on-alumina catalyst. The cycle oil is sent into contact with the catalyst at 750 F. and 1500 `p.s.i.g. A space velocity of 1.3 is maintained and hydro- Gasoline, vol. percent 44.8 Butane, vol. percent 14.5 Dry gas, wt. percent 10.4 Coke, wt. percent 5.7 Gas oil, vol. percent 37.0

Thus it is seen that hydrogenation of the recycle oil stream results in increased gasoline and butane yields at the expense of less desirable products such as coke and dry gas.

It is claimed: i

l. A process for catalytically cracking a residual petroleum stock which contains about 1 to 300 p pm. of metallic impurities selected from the group consisting of vanadium and nickel in an amount sufficient to cause deterioration in selectivity of a silica-based cracking catalyst and at least about 10% hydrogenated recycle stock,-Which comprises subjecting the petroleum oil to catalytic cracking in the presence of a synthetic gel, silica-based cracking catalyst under conditions sufficient to deposit said metallic impurities on the said catalyst and to produce a conversion to lower boiling materials, passing catalyst from the cracking zone to a regeneration zone wherein carbon is burned from the catalyst, passing regenerated catalyst to the cracking zone, withdrawing from the said cracking regeneration system a portion of the contaminated catalyst containing at least about 200 ppm. nickel and at least about 500 p.p.m. vanadium, removing at least about 60 to 90% of the nickel and about 20 to 40% of the vanadium from the catalyst and returning the demetallized catalyst to the said catalytic cracking system, fractionating the products from cracking to separate a substantially metals-free fraction boiling essentially above about 400 F. and a gasoline fraction, and subjecting the fraction boiling essentially above 400 F. to catalytic hydrogenation at a temperature of about 600 to 900 F., a pressure from about 300 to 3000 p.s.i.g. and with about 50 to 20,000 standard cubic feet of hydrogen per barrel of hydrocarbon oil and sending hydrogenator eluent to catalytic cracking.

2. The process of claim 1 in which the cracking catalyst is a synthetic gel silica-based cracking catalyst.

3. A process for catalytically cracking a residual petroleum stock which contains about 1 to 300 p.p.m. of metallic impurities selected from the group consisting of vanadium and nickel and at least about hydrogenated recycle stock, which comprises subjecting Ithe petroleum oil to catalytic cracking in the presence of a synthetic gel silica-based cracking catalyst at a temperature of about 750 to 1000 F. and a pressure between about atmospheric and 100 p.s.i.g., passing catalyst from the cracking Yzone to a regeneration zone wherein carbon is burned from the catalyst, passing regenerated catalyst vto the Vcracking zone, 'withdrawing from the said cracking regeneration system a portion of the contaminated catalyst containing at least about 200 p.p.m. nickel and at least about 500 p.p.m. vanadium, removing at least about 60 to 90% of the nickel and about 20 to 40% of the vanadium from the catalyst and returning the demetallized catalyst to the said catalytic cracking system, fractionating the products from cracking to separate a substantially metalsfree fraction boiling`"'essentially above 400 F. and a gasoline fraction, and subjecting the fraction boiling essentially above 400 F. to catalytic hydrogenation at a temperature of about 600 to 900 F., a pressure from about 300 to 3000 p.s.i.g. and about 100 to 10,000 stand- Vard cubic feet of hydrogen per barrel, and sending hy- 4. A process for catalytically cracking a residual petroleum stock which contains about 1 to 300 p.p.m. of

metallic impurities selected from the group consisting of vanadium and nickel in an amount suicient to cause deterioration in selectivity of a synthetic gel silica-based cracking catalyst and at least about 10% hydrogenated recycle stock, which comprises subjecting the petroleum oil to catalytic cracking under conditions suiiicient to deposit said metallic impurities on the said catalyst and to produce a conversion to lower boiling materials, passing catalyst from the cracking zone to a regeneration zone wherein carbonvis burned from the catalyst, passing regenerated catalyst to the cracking zone, withdrawing from the said cracking regeneration system a portion of the catalyst containing at least about 200 p.p.m. nickel and at least about 500 p.p.m. vanadium, demetallizing the withdrawn catalyst to remove about 10 to 90% of the said poisoning metal content on the catalyst by contacting the catalyst with a molecular oxygen-containing gas at a v temperature of about 1000 to about 1800 F. to enhance subsequent vanadium removal, suliiding the poisoning metal containing component on the catalyst by contact with a suliiding agent at a temperature of about 500 to 1500 F. to enhance sub-sequent nickel removal, chlorinating poisoning metal containing component on the sullided catalyst by contact with'an essentially anhydrous chlorinating agent at a temperature of about 300 to 1000 F., contacting the chlorinating agent-treated catalyst with a liquid essentially aqueous medium to remove dispersible nickel and vanadium components from the catalyst, and returning the demetallized catalyst to said catalytic cracking system, fractionating the products from cracking to :separate a substantially metals-free fraction boiling essentially'abov'er400. F. and a gasoline fraction, and subjecting the fraction boiling essentially above about 400 F. to catalytic hydrogenation at a temperature of about 600 to 900 F., a pressure from about 300 to 3000 p.s.i.g. and with about 50 to 20,000 standard cubic feet of hydrogen per barrel of hydrocarbon oil and sending hydrogenator eflluent to catalytic cracking.

5. A process for catalytically cracking a residual petroleum stock which contains about l to 300 ppm. of metallic impurities selected from the group consisting of vanadium and nickel and at least about 10% hydrogenated recycle stock, which comprises subjecting the petroleum oil to catalytic cracking in the presence of a synthetic gel silica-based cracking catalyst under conditions sufticient to deposit said metallic impurities on the said catalyst and to produce a conversion to lower boiling materials, passing catalyst from the cracking zone to a regeneration zone wherein carbon is burned from the catalyst, passing regenerated catalyst to the cracking zone, withdrawing from the said cracking-regeneration system a portion of the catalyst containing at least about 50 p.p.m. nickel and at least about 50 p.p.m. vanadium, removing about 10 to 90% of said selected metallic impurities from the catalyst and returning the demetallized catalyst to the said catalytic cracking system, fractionating the products from cracking to separate a substantially metals-free fraction boiling essentially above 400 F. and a gasoline fraction, and subjecting the fraction boiling essentially above 4009 F. to catalytic hydrogenation at a temperature of abouty 600 to 900 F., a pressure from about 300 to 3000 p.s.i.g. and with about 50 to 20,000 standard cubic feet of hydrogen per barrel of hydrocarbon oil and sending hydrogenator eiuent to catalytic cracking.

6. The process of claim 4 wherein the sulding agent is H28 and the chlorinating agent is an equimolar mixture of C12 and C014.

7. A process for catalytically cracking a residual petroleum stock boiling above the gasoline range and containing up to about 10 p.p.m. nickel and up to about 20 p.p.m. vanadium and at least about 10% hydrogenated recycle stock, which comprises subjecting the residual oil to catalytic cracking in the presence of a synthetic gel, silica-alumina cracking catalyst` at a .temperature of about ,750 to l000 F. suiicient to deposit said metallic impurities on the said catalyst and to convert about 40 to 70% of said residual oil to lower boiling materials, pass- Ving catalyst from the cracking zone to a regeneration zone wherein carbon is burned from the catalyst by contact with air at a temperature of about 900 to 1200'J F., passing Vregenerated catalyst to the cracking zone, withdrawing metal contaminated catalyst from the cracking-regeneration system at a rate of about 50 to 150% per day ofthe catalyst in the cracking-regeneration system, the withdrawn, contaminated catalyst containing at least about 200 p.p.m. nickel and at least about 500 p.p.m. vanadium, demetallizing withdrawn catalyst to remove about 60 to of the nickel and about 20 to 40% of the vanadium, and returning the demetallized catalyst to the cracking system, fractionating the products from cracking to separate a substantially metals-free fraction boiling essentially above 400 F. and a gasoline fraction, and subjecting the fraction boiling essentially above 400 F. to catalytic hydrogenation at a temperature of about 600 to 900 F., and a pressure from about 300 to 3000 p.s.i.g. and about to 10,000 standard cubic feet of hydrogen per barrel, and sending hydrogenator eiuent to catalytic cracking as said recycle stock.

8. A process for catalytically cracking a residual pe- Vtroleum stock boiling above the gasoline range and conto catalytic cracking in the presence of a synthetic gel,

silica-alumina catalyst at a temperature of a'bout 750 to the said catalyst and to convert about 40 to 70% of said residual oil to lower boiling materials, passing catalyst from the cracking zone to a regeneration zone wherein carbon is burned from the catalyst by contact with air at a temperature of about 950 to 1200 F., passing regenera-ted catalyst to the cracking zone, withdrawing metal contaminated catalyst from the cracking-regeneration system at a rate of about 50 to 150% per day of the catalyst in the cracking-regeneration system, the withdrawn, contaminated catalyst containing at least about 200 p.p.m. nickel and at least about 500 p.p.m. vanadium, demetallizing withdrawn catalyst to remove about 60 to 90% of the nickel and about 20 to 40% of the vanadium by contacting the catalyst with molecular oxygen-containing gas at a temperature of about 1150 to about 1350" F., sultiding the poisoning metal containing component on the catalyst by contact with H2S at a temperature of about 800 to 1300 F., chlorinating poisoning metal-containing component on the catalyst by contact with an equimolar mixture of C12 and CCL, at a temperature of about 300 to 1000 F., removing poisoning metal chloride in vapor form from the catalyst, contacting the catalyst with a liquid, essentially aqueous medium to remove soluble poisoning metal chloride from the catalyst and returning the demetallized catalyst to the cracking system, fractionating the products from cracking to separate a substantially metals-free fraction boiling essentially above 400 F. and a gasoline fraction, and subjecting the fraction boiling essentially above 400 F. to catalytic hydrogenation at a temperature of about 600 to 900 F., and a pressure from about 300 to 3000 p.s.i.g. and about 100 to 10,000 standard cubic feet of hydrogen per barrel, and sending vliydrogenator ellluent to catalytic cracking as said recycle stock.

9. A process for catalytically cracking a residual petroleum stock which contains more than about 0.3 p.p.m. NiO and more than about 0.5 p.p.m. V205 and at least about hydrogenated recycle stock, which comprises subjecting the petroleum oil to catalytic cracking in the presence of a synthetic gel, silica-alumina cracking catalyst at a temperature of about 750 to 1000 F., sutlcient to deposit said metallic impurities on the said catalyst and to produce a conversion to lower boiling materials, passing catalyst from the cracking zone to a regeneration zone wherein carbon is burned from the catalyst passing regenerated catalyst to the cracking zone, withdrawing from said cracking-regeneration system a portion of the catalyst containing at least about 200 p.p.m. nickel and at least about 500 p.p.m. vanadium, removing at least about of the nickel and at least about 15% of the vanadium from the catalyst and returning the demetallized catalyst to the said catalytic cracking system, fractionating the products from cracking to separate a substantially metals-free fraction boiling essentially above 400 F. and a gasoline fraction, and subjecting the fraction boiling essentially above 400 F. to catalytic hydrogenation at a temperature of about 600 to 900 F. and a pressure of from about 300 to 3000 p.s.i.g. and with about to 10,000 standard cubic feet of hydrogen per barrel, and sending hydrogenator etiluent to catalytic cracking as said recycle stock.

10. The process of Claim 9 wherein the residual petroleum stock contains more than about l p.p.m. NiO and more than about 0.5 p.p.m. V205.

References Cited bythe Examiner UNITED STATES PATENTS 2,414,736 1/47 Gray 208-113 2,488,744 11/49 Snyder 208-113 2,575,258 11/51 Cornell et al 252-417 2,671,754 3/54 De Rosset et al. 208-67 3,008,895 11/61 Hansford et al 208-112 ALPHONSO D. SULLIVAN, Primary Examiner.

Claims

9. A PROCESS FOR CATALYTICALLY CRACKING A RESIDUAL PETROLEUM STOCK WHICH CONTAINS MORE THAN ABOUT 0.3 P.P.M. NIO AND MORE THAN ABOUT 0.5 P.P.M. V205 AND AT LEAST ABOUT 10% HYDROGENATED RECYCLE STOCK, WHICH COMPRISES SUBJECTING THE PETROLEUM OIL TO CATALYTIC CRACKING IN THE PRESENCE OF A SYNTHETIC GEL, SILICA-ALUMINA CRACKING CATALYST AT A TEMPERATURE OF ABOUT 750 TO 1000*F., SUFFICIENT TO DEPOSIT SAID METALLIC IMPURITIES ON THE SAID CATALYST AND TO PRODUCE A CONVERSION TO LOWER BOILING MATERIALS, PASSING CATALYST FROM THE CRACKING ZONE TO A REGENERATION ZONE WHEREIN CARBON IS BURNED FROM THE CATALYST PASSING REGENERATED CATALYST TO THE CRACKING ZONE, WITHDRAWING FROM SAID CRACKING-REGENERATION SYSTEM A PORTION OF THE CATALYST CONTAINING AT LEAST ABOUT 200 P.P.M. NICKEL AND AT LEAST ABOUT 500 P.P.M. VANADIUM, REMOVING AT LEAST ABOUT 50% OF THE NICKEL AND AT LEAST ABOUT OF THE VANADIUM FROM THE CATALYST AND RETURNING THE DEMETALLIZED CATALYST TO THE SAID CATALYTIC CRACKING SYSTEM, FRACTIONATING THE PRODUCTS FROM CRACKING TO SEPARATE A SUBSTANTIALLY METALS-FREE FRACTION BOILING ESSENTIALLY ABOVE 400*F. AND A GASOLINE FRACTION, AND SUBJECTING THE FRACTION BOILING ESSENTIALLY ABOVE 400*F. TO CATALYTIC HYDROGENATION AT A TEMPERATURE OF ABOUT 600 TO 900*F. AND A PRESSURE OF FROM ABOUT 300 TO 3000 P.S.I.G. AND WITH ABOUT 100 TO 10,000 STANDARD CUBIC FEET OF HYDROGEN PER BARREL, AND SENDING HYDEROGENATOR EFFLUENT TO CATALYTIC CRACKING AS SAID RECYCLE STOCK.