|Publication number||US3168459 A|
|Publication date||Feb 2, 1965|
|Filing date||May 4, 1961|
|Priority date||May 4, 1961|
|Publication number||US 3168459 A, US 3168459A, US-A-3168459, US3168459 A, US3168459A|
|Inventors||Anderson Arvin D, Sanford Robert A|
|Original Assignee||Sinclair Research Inc|
|Export Citation||BiBTeX, EndNote, RefMan|
|Patent Citations (8), Referenced by (18), Classifications (9)|
|External Links: USPTO, USPTO Assignment, Espacenet|
Feb. 2, 1965 A. D. ANDERSON ETAL 3,168,459
CRACKING A METAL-CONTAMINTED RESDUAL OIL Filed May 4. 1961 ROBERT A. SANFORD MZ@ WW2/@16A ATTORNEYS.
United States Pafeftf@ 3,168,459 h1Patented Feb. 2, 1965 3,168,459, v CRACKING A METAL-CQNTAMENATED V RESEDUALXL p Ari/in D.Y Anderson, Anaheim, Calif., and Robert A. Sanford, Homewood, Ill., assigno'r's, by mesne assignments, to Sinclairy Research, Inc., New York, N.Y., a corporation of Delaware Filed May 4, 1961, Ser. No. 107,679 13 Claims. (Cl. 26S-'57) This invention relates to the upgrading of hydrocarbon oils, particularly high boiling petroleum oils, to obtain lighter components including gasoline of realtively high octane number. More specifically, this invention comprises an integrated process for hydrogenating and solvent deasphalting residual oils to obtain feedstocks capable of being catalytically converted to lighter boiling material. The cracking system is coupled with a catalyst demetallization system.
The catalytic cracking of various heavier mineral hydrocarbons, for instance, petroleum or other mineral oil distillates such as straight run and cracked gas oils; shale oils; petroleum residues, etc., has been proposed for many years and the catalytic cracking of gas oils is practiced commercially to a considerable extent. The behavior of a hydrocarbon feedstock in the cracking reactions depends upon various factors including its boiling point, carbonforming tendencies, content of catalyst contaminating metals, etc., and these characteristics may affect the operav tion to an extent which makes a given feedstock uneconomical to employ. By and large, residual stocks have not been catalytically cracked on a commercial scale as their carbon-forming tendencies and catalyst poisoning d ing catalyst employed can be discarded often to prevent i' a high accumulation of poisoning metals in the cracking system, this type of operation represents a substantial cost factor. Improvements in the feedstock characteristics become even morel important as the cost of the catalyst rises and thus the effects of low feedstock quality are particularly burdensome in systems' employing c racking catalysts containing relatively expensive synthetic components.
Desirable petroleum cuts are frequently obtained from a crude by solvent extraction. In solvent extraction processes, such as solvent deasphalting, the coke-formers and metal contaminants tend to segregate into the asphalt railinate fraction. Refiners, therefore, are provided with an extract fraction of the crude which is relatively low in metals and coke-formers to serve as feed to catalytic cracking. The asphalt raffinate fraction is generally put to lovv value uses, such as heating oils or road surfacing materials.
It has been proposed heretofore to hydrogenate the various heavy metal-containing hydrocarbon oils prior tochargingY them, 6r a fraction thereof, to a catalytic cracking operation. By so doing the hydiocarbon maybe given an improved hydfgeii-'to-carb'on ratio and the amount of 'contaminants, such as coke-f'rmers, sulfur, 'and nitrogen may be reduced. The c'ntnt of metals which poison cracking catalysts is also reduced arid removal of any vsiil'nstanti'al 'amount'of these contaminants from the cracking feed tends to :enhance efficiency of the catalytic cracking operation. ,The degree of feedstock improvement from hydrogenatio'nis dependent, however, upon several factors which include ythe severity of 'reaction andthe amount of hydrogen consumed. High severity can increase the extent o'f metals removali however, it may involveV a greater consumption of hydrogen, a larger capital investment for high-pressure equipment, and a reduction in the yield of cracking feedstock. Petroleum fractions containing large amounts of coke-forming and/ or metal components, such as the asphaltic and residual materialsv described above, frequently require such severe hydrot'reating to make them trouble-free cracking feeds that the expense of such hydrotreating is not practical.
This invention employs partial demetallization ofra cracking feed by hydrogenating and deasphalting, along with demetallization of the cracking catalyst by procedures to be described, to achieve greater economy than would be obtained by employing only one of the demetallizing techniques: hydrogenation and/orrdeasphalting on the one hand, or poison removal from the cracking catalyst, on the other hand, in the attempt to obviate poisoning effects. In the operation of this invention, hydrogenation, solvent deasphalting and catalytic cracking of heavier mineral hydrocarbon oil feed stocks to produce gasoline Vcan be combined and balanced with a procedure for reducing poisoning metals on the cracking catalyst to present a Vmuch more attractive alternative to the individual operations described above for overcoming an overallmetal's problem. Under these conditions all four of the hydrogenation, solvent deasphalting, catalyticcracking and cracking catalyst denietalli'zation can Vbe' operated to make a relatively low consumption of hydrogen during hydrogenatio'n more attractive, with the hydrogen being better utilized through minimization of dehydrogenation in the catalytic cracking operation; In this method the metal-containing hydrocarbon feedstock is hydrogenated under conditions giving a desired hydrogenation effect with partial, but not complete, removal of Apoisoning metals. The hydrogenated o'il ora selected portion thereof boiling above about 400 F. and containing a significant amount of metal contaminant is contacted with a solvent to further remove rntal contaminants and the extract, containing substantially less contaminants than the hydrogenated product, is then catalytically cracked by itself or after blending with conventional cracking feedstocks to produce in good yield a gasoline fraction of relatively high octane rating. As the cracking operation proceeds the catalyst is treated to remove accumulated metal poisons and is then reused in the cracking operation. Besides preserving the effects of hydrogenatio'ngby minimizing the dehydrogention effects of a poisoned catalyst, catalyst demetallizati'on provides a cracking operation in which there is relatively less carbon laydown on the catalyst. This further increases4 gasoline yieldin a system' having a gig/en' carbon-burning capacity. The cracking aspect of t 's invention with its de'me'talliiation features is economically attractive when a cracking feedstock is obtained containing as little as about 1 p.p.m. nickel and/or about 1 p.p.m. vanadium.
The feeds to the present process comprise petroleum residua which may be exemplified by vacuum residua, atmospheric residua, tars, pitches, etc., boiling primarily above about 600 F. or even above about 900 F. The residual feed often has an API gravity in the range of about to 25, a Conradson carbon content 4in the range of about 3 to 35 weight percent and a viscosity often above about 200 seconds Saybolt Furol at 210 F. These charge stocks contain metals which are poisonous to the cracking catalyst to be used subsequently. The residual feedstock Will usually include at least about 5 or 10 parts per million of one or both of vanadium and nickel. These are materials generally avoided for use as catalytic cracking feeds. In the pro-cess of this invention metal contents above these ranges may be present; it will be apparent that oils having metal contents in these generally undesirable ranges are the oils which this invention salvages. In most cases, the totalrof one or both of these metals in the residual will be at least about 5 p.p.m. and may often contain about 25 or 50 p.p.m. nickel and about 50 or 100 p.p.m. vanadium. The maximum amount of metals in the residuals can vary widely. The maximum amount of these poisoning metals in the residual stock will usually not exceed about 500 p.p.m. nickel, and/or about 1000 p.p.m. vanadium, to be economically processed. Residuals containing a low level of metals contaminants profit by the hydrogenation and deasphalting step of this invention by improvements in their general cracking characteristics. Although referred to as metals the contaminants may be in the form of free metals or metal compounds and it is to be understood that the term metal used herein refers to either form.
In hydrogenation the hydrogen has a number of effects on the feedstock. One use is to dissociate the heavy poisoning metals from their compounds in the feedstock. Hydrogenation also serves to saturate components of the feed which are susceptible to such. Thus hydrogenation increases the hydrogen-to-carbon ratio ofthe hydrogenated effluent reducing the coke-forming tendencies of the feedstock. Also, hydrogenation may cause a certain amount of the feed to be converted (cracked) to lower boiling materials. This'factor may sometimes make lresidual oils a more desirable feedstock for this invention than lower boiling materials of the same metal content. Since more cracking is not only permissible but indeed sometimes desirable when hydrogenation is performed on a residual to prepare it for catalytic cracking, the hydrogenation may be performed at higher temperatures. At these temperatures, hydrogenation can perform its demetallization and saturation functions at a lower pressure, thereby lowering equipment costs. Any gas or gasoline produced in the hydrogenation and unsuitable for use as feed to the catalytic cracking may be used as reformer feedstock, although hydrogenation of residuals usually does not produce overly large Aamounts of these materials. Usually greater than about of the residual hydrocarbon charge may be converted to lower boiling normally liquid materials by hydrogenation and quite frequently at least about 25% is converted. Conversion to lower boiling materials rarely exceeds 80% of a residual charge.
In the hydrogenation operation hydrogen is consumed by chemical combination with a component of the hydrocarbon feed. Where hydrocracking is performed the hydrogen consumption Will be high, as in the case of treating most residual hydrocarbon oils. Hydrogenation operations are subject to various modifications from which the petroleum rener may choose, depending upon the residual stock to be treated and the results desired. Hydrogenation may be performed by free or molecular hydrogen in the presence of a catalyst or by a hydrogen donating chemicalwith or without a catalyst. Elevated temperature (about 400-1200 F.) and pressure (atmospheric to 3000 p.s.i.g.) conditions usually prevail in hydrogenation, and Within these ranges conditions can be chosen to give the desired degree of oil demetallization, saturation and/or cracking.
. Hydrogen donor diluent cracking (HDDC) is Widely known in the art and is illustrated in abandoned application Serial No. 365,335, filed July l, 1953, by Arthur W. Langner, Jr., as disclosed in U.S. Patent 2,772,718. The donor diluent is a material, generally a hydrocarbon, which has the ability to take up hydrogen in a hydrogenation zone and readily release it in a thermal treating zone. It is believed that the donor diluent operates by yielding hydrogen atoms to the radicals that have been created from the residuum by the thermal treatment, thereby upgrading the residuum and preventing condensation and/ or polymerization of the radicals. In donor diluent operations, the donor diluent material is substantially unaltered as it passes through the process, and it is usually customary to recycle the material so that it is used over and over again as a hydrogen carrier. Donor materials may be added as a relatively pure chemical such as tetralin or decalin or in admixture with other materials, particularly hydrocarbons, or the donor diluent may be a partially hydrogenated catalytic cycle oil, a partially hydrogenated lubricating oil extract or other partially hydrogenated aromatic. Hydrogen donors usually contain condensed ring aromatics in sulicient quantities to serve as a hydrogen carrier. These aromatics are partially hydrogenated; there is added to them some easily removable hydrogen atoms but not enough to convert the aromatics substantially to naphthenes. This material after being partially hydrfogenated, can be admixed with the feedstock to this process .and the mixture thermally treated, whereby the hydrogen is transferred from the partially hydrogenated material to the hydrogen-deficient residuum.
After blending of hydrogen donor diluent and residuum, -the blend is then taken to a thermal or catalytic cracking zone where recycle oil from the cracker, fractionator or deasphalting zone of this pnocess may also be added. The mixture of residual and diluent may be thermally cracked by heating to a temperature of about 700 to 1200 F., preferably about 800 to 1000 F. at pressures within the range of from about atmospheric to about 2000 p.s.i.g., preferably from about atmospheric to about 200 p.s.i.g. with a holding time of about 5 to 30 minutes. Aconventional coil or coil and drum heater may be used. A Wide range of conversion may be obtained by varying the temperature or feed rate in the thermal cracking operation depending upon the feed to be processed, to produce from about 40 to 80 percent conversion to lower boiling materials per pass. The Weight ratio of diluent to residual plus recycle bottoms is usually between about 0.1/ l to about 10/ 1, preferably about 0.5 to 2/ 1. A high rate of diluent to feed and a moderately high hydrogen content tend to reduce coke formation and remove metal contaminants at any severity. However, the severity of thermal cracking is primarily a function of cracking temperature and feed rate. The nature of the residual and its prior processing, if any, may affect the cracking severity and the amount of metal removal.
In the catalyzed donor cracking method (CDC) a petroleum residual can be contacted at elevated temperatures with an essentially anhydrous catalyst comprising the hydride of a halogen having an atomic number of 35 to 53 and in the presence of a hydroaromatic material. As a result, valuable -low molecular weight hydrocarbon gases primarily boiling below C3 and gasoline having an end boiling point of about 430 F. are produced along with a goodyield of gas oils which provide a high quality catalytic cracking stock, and a residual oil boiling primarily above 950 F. which is reduced in contaminating metals. The catalyst in this process is hydrogen iodide or hydrogen bromide. These can be added as such to the reaction zone or the corresponding elemental halogen, that is, bromine or iodine, or other material which gives the halogen hydride may be charged. The halogen hydride is apparently in equilibrium in the reaction zone with elemental halogen. The halogen hydride may predominate in the equilibrium mixture. The catalyst selected is employed in a substantially anhydrous form although it may be used in solution with alcohol or other solvents. The amount of catalyst utilized normally depends on the characteristics of the residual treated, for instance, the type and amount of metal contaminants, and the amount of nitrogen, sulfur, etc. present. The amount of catalyst employed is generally from about 0.01 to 5 percent by weight of the residual treated with a preferred amount being about 0.1 to 2 percent.
The hydrogen donor can -be contacted with the residual oil to be treated in any suitable manner. The donor may be added to the oil prior or subsequent to the addition of the catalyst to the oil. The amount of hydroaromatic compound employed is generally at `least about 'of the residual feed and usually is in a range of from about 50 to 200 percent by weight of the oil treated. The hydroaromatic material is a liquid at the conditions of the process and acts as a hydrogen donor as described above. When added in admixture with other hydrocarbons the hydroaromatic is usually at least about 40 or 50%, preferably at least about 75% of the mixture. During the CDC treatment the temperature is usually in the range of about 700 to 1200 F. with a preferred temperature being about 750-850 F. The pressure may vary widely depending on the particular feedstock undergoing treatment and the temperature employed, but it is essentially that substantial cracking and conversion to lower boiling oils occurs. The pressures will generally be elevated and vary from about atmospheric to 2000 p.s.i.g. with a preferred range being about 500 to 1500 p.s.i.g. In this process it is essential that the conditions of temperature and pressure be such that the particular feedstock undergoes substantial hydrocracking during treatment.
The addition of free hydrogen in the CDC process is normally advantageous as it can increase the liquid product yield and aid in the hydrocracking. Under normal operating conditions it may be desirable to employ hydrogen, preferably at a partial pressure of at least about 100 p.s.i.g. There does not appear to be any particular benet in providing a hydrogen partial pressure in excess of about 1500 p.s.i.g., the preferred pressure being from about 300 to 1000 p.s.i.g. Hydrogen consumption will usually be no more than about 1000 standard cubic feet per barrel of residual treated. This process allows for a conversion of at least about 20% of the feedstock to a liquid material boiling below about 950 F.
Hydrogenation of the residual using free hydrogen and a catalyst may be conducted by contacting the petroleum feed with the catalyst in the presence of free hydrogen under superatmospheric pressure. The hydrogenation catalysts generally known in the art can be employed. Calcined solid hydrogenation catalysts are preferred and they are usually disposed as a fixed bed of macrosized particles, say of about 1/a to 1A" in diameter and about l" to 1 or more in length. A moving bed of macrosized catalyst or a uidized 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 VIII catalysts of the iron group, molybdenum, tungsten and combinations thereof. Frequently the metals are disposed as inorganic components, for instance oxides, suliides Ior other compounds, supported on a solid carrier exemplied by alumina, silica, etc. Advantageously, the catalyst contains a combination of metals of the iron group with vanadium or a metal of Group VIfz of the periodic chart having atomic numbers from 42 to 74, i.e., molybdenum and tungsten. A commercial catalyst contains cobalt and molybdenum, e.g., cobalt molybdate, supported on alumina. The amount of catalytically active `(i 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 cataly-tically active metal when combinations are used.
Catalytic hydrogenation conditions are selected to give the desired hydrogen consumption and poisoning metals reduction. In general, however, an elevated temperature such as about 600 to 900 F. may be employed and the pressures are generally superatmospheric usually falling in the range of about 300 to 3000 p.s.i. Free or molecular hydrogen may be 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 l0 or more WHSV (weight of hydrocarbon feedstock per hour per weight of catalyst). Hydrogen consumption is usually at least about 70-300 standard cubic feet of hydrogen per barrel of hydrocarbon oil feed. Where hydro'cracking is desired, such asin the treatment of residuals, the hydrogen consumption is often in the range of about 1000 to 2000 or more standard cubic -feet per barrel. Residual oils are often treated at about 750 to 900 F., at a pressure over about 1000 p.s.i.g., preferably about 1500 to 2500 p.s.i.g., and about to 10,000 standard cubic feet of hydrogen per barrel.
The treatment of residual material under non-.cracking conditions in the presence of catalytic material and hydrogen or a hydrogen donor is conducted under suitable temperature and pr-essure conditions so that there is not a substantial amount of cracking, i.e., less than about 15 weight percent of the petroleum feedstock, preferably less than about 5% is cracked. The treating temperature Will usually be in the range of about 400 to 700 F., with about 600 to 700 F. being ino'st suitable. The total fpressure in the reactor is usually at least about 100 p.s.i.g., .more often at least about 300 p.s.i.g., and rio reason has been seen yfor going above about 3000 p.s.i.g. Preferably the pressure lwill not be above about 1500 to 2000 p.s.i.g. t is preferred not to introduce `free hydrogen into the system when treating residuals with .a hydrogen donor diluent because the `added hydrogen is not consumed efficiently. However, when free Ihydrogen is introduced into the system the hydrogen partial pressure is about 100-2000 or 3000 p.s.i.g., preferably `about 200-500 p.s.i.g. The length of time of the treatment may vary widely so long as conversion of the' petroleum feed is limited as noted before. The treatment may take from about 0.1 to 5 hours lor more and seems of little benefit after -10 hours. The preferred time is about 0.5 to 3 hours, and, of course, lower temperatures may require longer contact times to obtain a given result.
As in HDDC, the hydroaroma-tic compound can be contacted with the hydrocarbon oil to be treated in any suitable manner and in the proportions recited. lWhen treating residuals the hydrogen donor material is preferably low boiling so that it may be separa-ted overhead by distillation `from the product. Since substantial cracking is avoided in this hydrogenation operation the lighter donor is readily recovered since it will not be unduly contaminated with light gas oil.
The catalyst in the non-cracking process can be essentially hydrogen iodide which can be added as such to the reaction zone or iodine or another hydrogen iodide-producing material may be added. yIn any event, the hydro gen iodide is apparently in equilibrium with elemental iodine in the reaction zone although the catalyst may be predominantly hydrogen iodide. The catalyst can be .contacted with the petroleum feedstock in any convenient manner and the catalyst is essentially in anhydrous form although it may be used in solution with alcohol or .other solvents. The amount of catalyst used can depend upon the reaction conditions `a-nd the amount of feedstock demetallizaltion required, but is generally from about 0.1 to
7 210% of the oil to be hydrotnea-ted, preferably about 1 to 10%.
After the hydrogenation the hydrogenated product material may be fractionated to obtain gaseous hydrocarbons, gasoline, a diluent cut and heavy bottoms comprising essentially the entire gas oil fraction and material boiling above about 950 F. The gasoline fraction may be hydroformed, thereby increasing its -octane rating. The diluent cut is separated from other portions of the hydrotreater effluent and hydnotreated to increase its hydrogen-tocarbon ratio sufficient for reuse as a hydrogen donor. Such hydnogenation may be by the use of free hydrogen in the presence of a solid catalyst as described above.
The entire hydrogenated product may be charged |to the deasphalter, but it is preferable in the process of this invention to remove the substantially metals-free gas oil fraction boiling essentially in the range fof about 400 to 11100 F., which is amenable to direct catalytic cracking fnom the bottoms before deasphalting and conduct it directly to the cracking unit. The bottoms .are then treated in the deasphalting unit to further reduce the metal contaminant content.
Hydnogenation gives a partial reduction in metals content Iof the residual feed. The metals remaining in the hydrogenated produ-ct would accumulate on the catalyst during the cracking operation and unless steps are taken to prevent excess accumulation, excessive dehydrogenation takes place in the cracking, partially undoing the work performed in the hydrogenating step and severely reducing the yield of gasoline in the cracker effluent. In this invention hydrogenation may remove only about 10% of the poisoning metal in the residual feed, but preferably much more of the poison. Thus the hydrogenated prodnot perhaps contains about 50 to 90 or more Weight percent less of one or both of nickel and vanadium than the hydrocarbon -charged .to the hydrogenating reaction; preferably there is this much reduction in nickel and vanadium or in each of these metals. Frequently the reduction of one or all of nickel, vanadium and iron will be about 65 to 90 weight percent. The hydrogenated product contains at least about 2 p.p.rn. nickel and/ or about 3 p.p.rn. vanadiurn, more usually about to 50 p.p.rn. total nickel and vanadium, but rarely more than 150 p.p.rn.
Hydrogenation usually Vdoes not remove metal contaminants to a point that is insignificant in subsequent catalytic cracking. The hydrotreated product also may contain some hydrocarbon constituents which are unsuitable for inclusion in a catalytic cracking feedstock. Therefore, a bottoms product Yfrom the hydrogenation operation is deaspl'nalted by solvent extraction to obtain further reduction in metal content and a product of low Conradson carbon and thus of relatively low coke-forming tendency in `catalytic cracking. The hydrogenated product, or fraction thereof, may be treated, lfor example, with a liquid hydrocarbon solvent having from I'about 3 to 7 carbon atoms. The metal poisons in this hydrogenated material gather primarily in the asphalti-c raffinate phase. The deasphalted oil phase, upon removal of the solvent, is a good feed for catalytic cracking; the :asphalti-c phase, containing the contaminants and heavy asphaltene constituents @of the hydrogenated product or portion thereof may be recycled to the hydrogenation zone for further treatment. The amount of gas oil component recoverable from the hydrogenated product by any particular solvent varies with the volume natio of solvent-to-oil used in the extraction. However, the amount recover-ed is not a linear function of the solvent-to-oil ratio, i.e., for each solvent a certain ratio exists where gas oil recovery is at a minimum, When a greater solvent-to-oil ratio is employed, a greater amount of gas oil having a low metals cont-ent is obtained. Also when a solvent-to-oil ratio below Ithis minimum is used high yi-elds of gas oil are obtained but the gas oil produced has a high content of the metal contaminants.
The solvent may be chosen from a number of hydrocarbon types, as well as other extracting solvents well known in the art. Suitable hydrocarbon solvents are liqueed propane, propylene, butylene, butane, including isoand normal butane, n-pentane, etc., or any combination thereof, which are normally employed in a solventto-oil ratio of from about l to 12/1, the preferred solventto-oil ratio being from about 2/1 to 7/1. The solvent extraction deasphalting step may be conducted within a Wide range of temperatures. The minimum will be the softening point of the asphalt feed, while the maximum will be the boiling temperature of the solvent at the pressure used. The extraction step will frequently be performed at a temperature of about to 300 F. There may be a temperature gradient in the deasphalting step with the highest temperature found at the deasphalted oil (extract phase) outlet. The temperature gradient may be within the range of about 0 to 60 F. Preferably about a Ztl-45 F. temperature gradient should be maintained, for example a 30 gradient may be used where the top of the tower is the point where the deasphalted oil is withdrawn. Thus, the temperature at the top of the extraction tower may be maintained at 180 F. while the bottom of the tower, where the asphalt is removed, may be at about F. Normal operating pressures should be higher than the vapor pressure of the solvent system used at the temperature of operation. The extraction may be carried out in a plurality of stages in one vessel or in a plurality of vessels in series. The separate stages may be conducted with a temperature gradient and pressure gradient between the stages.
After contacting, the mixture of residual and solvent separates into two phases, an extract phase containing solvent and gas oil components and a raffinate or asphalt phase. The two phases are separately withdrawn, the hydrocarbon is removed from the extract phase as desired, and the deasphalted gas oil amounting to about 40 to 95% by weight and preferably about 50 to 90% by weight of the hydrogenated material fed to the extraction unit is then used as a feedstock to a catalytic cracking operation. Deasphalting may remove as much as about 95 preferably about 30-80%, of the metal contaminants, such as nickel and vanadium, from the hydrogenated product or portion thereof. In the practice of the present invention the deasphalting generally provides an extract containing about 1-20 p.p.rn. nickel and/or about 1 to 30 ppm. vanadium calculated as the common oxide. Deasphalting may also remove about 30 to 80 weight percent of the carbon content of the hydrogenated oil and preferably deasphalting reduces the carbon content by as much as about 40 to 80% by Weight. The use of a catalyst demetallization unit with the catalytic cracker counteracts the remaining metals content, thereby enabling much deeper cuts of the residual feed to be brought to the hydrogenation, much deeper cuts of hydrogenated effluent distillate to be sent to catalytic cracking without solvent treatment and a much deeper solvent treating to be performed, giving a larger extract phase and less raffinate phase; a slightly higher metals content in the extract phase being rendered acceptable by catalyst demetallization.
The recovered deasphalted extract oil is subjected to catalytic cracking. Contaminating metals in greater quantities than are acceptable to the art generally are present in the cracker feedstock. The cracking feedstock boils above the gasoline range, preferably in the range of about 400 to 1200 F. and contains a significant amount of the solvent treated extract. The amount of this product in the cracking feed will be at least about 5-10%, preferably about 20-70%. The remaining portion of the cracker feed may comprise cracking feeds of more or less conventional types, that is, the cracking feed will be adjusted to provide a feed containing more than about l ppm. nickel and/ or 1 p.p.rn. vanadium and preferably the total feed to cracking will contain less than about 10 p.p.rn. nickel and/or 20 p.p.rn. vanadium. At least about 1 p.p.rn. and/or about 1 p.p.rn. vanadium is contributed to 9 the cracker feed by the extracted oil boiling above about 400 F.
Catalytic cracking is ordinarily effected to produce gasoline as the most valuable product and is generally conducted at temperatures of about 750 to 1050 F., preferably about 850 to 975 F., at pressures up to about 100 p.s.i.g., preferably about atmospheric to -15 p.s.i.g., and advantageously without substantial addition of free hydrogen to the system. In the cracking operation a batch, semi-continuous or continuous system may be used but most often is the latter.
The 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. The 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, preferably about 5 to 25%. These compositions are calcined to a state of very slight hydration. The cracking catalyst can be of macrosize, for instance bead form or linely divided form, and employed as a fixed, moving or fiuidiZe-d bed. In a highly preferred form of this invention finely divided (fluidized) catalyst, for instance having particles predominantly in the 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.
Vaporous products are taken overhead and a portion of the catalyst is continuously withdrawn and passed to a regeneration zone where coke or carbon is burned from the catalyst in a fluidized bed by contact with a free oxygen-containing gas before its return to the reaction zone. In a typical operation the catalytic cracking of the hydrocarbon feed would normally result in the conversion of about 40 to 70%, preferably about 50 to 60%, of the feedstock into a product boiling in the gasoline range. The eluent 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. Bottoms, that is, products boiling above the gasoline range conveniently are recycled to the hydrotreating or catalytic cracking zones by blend ing them with virgin residual feed and/or hydrogenated product. These bottoms, or cycle oil, are substantially free of metal poisons.
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 lto 1200 F., more usually about 1000 to ll50 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 the regenerator may be on the order of about six minutes and the oxygen content of the efliuent gases from the regenerator is desirably less than about 1/2%. 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 ac- 10 cumulate on the catalyst during the cracking operation. Unless steps are taken to prevent excess accumulation, excessive dehydrogenation takes place in the cracking, partially undoing the work performed in the hydrogenation step and severely reducing the yield of gasoline in the cracker efiiuent.
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 limit of the units tolerance for poison. The tolerance of the cracker for poison in turn determines .to a large extent the amount of metals removed in the catalyst demetallization procedure. Where the catalyst contains a greater amountof poisoning metal, a particular treatment will remove a greater amount of metal, for example, if ,the cracker can tolerate an average of l0() p.p.m. Ni and demetalliza tion 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 is 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 500G-10,000 p.pm. Catalyst demetallization is not economically justified 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 metal-s 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 Ztl-40% vanadium from the treated portion of catalyst. Preferably at least 50% of the equilibrium nickel content and 15% 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 is subjected 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 may sufhce as readily as a more 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. In a continuous operation of the commercial type a satisfactory treating rate may be about 5 to 50% of the total catalyst inventory in the system, per twenty-four hour day of operation although other treating rates may be used. With a continuously circulating catalyst stream, such as in the ordinary fluid system a slip-stream of catalyst, at the equilibrium level 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 in- 'clude one or more processing steps.
' Copending patent applications Serial Nos. 758,681, tiled September 3, 195.8, now abandoned; 763,833 and 763,834, filed September 29, 1958, now abandoned; 767,794, led October 17, 1958, now abandoned; 842,618, led September 28, V1959, now abandoned; 849,119, led October 28, 1959, now Patent No. 3,094,059; 19,313,'iled April 1, 1960, now abandoned; 39,810, filed June 30, 1960; 47,598, filed August 4, 1960; 53,380, now Patent No. 3,122,497, led September 1, 1960; 53,623, filed September 2, 1960; 54,368, now Patent No. 3,122,512, 54,405, now 'Patent No. 3,122,510 and 54,532 now abandoned, led September 7, 1960; 55,129, 55,160 and 55,184, filed September 12, 1960; 55,703, led September 13, 1960; 55,838, filed September 14, 1960, now abandoned; 67,518, led November 7, 1960; 73,199, led December 2, 1960; and 81,256 and 81,257, led January 9, 1961, now abandoned; all of which are hereby incorporated by reference, describe procedures by which vanadium and other poisoning metals included in a solid oxide hydrocarbon conversion catalyst are removed by dissolving them from the catalyst or subjecting the catalyst, outside the hydrocarbon conversion system, to elevated temperature conditions which put the metal contaminants into the chloride, sulfate or other volatile, water-dispersible or more available form. A signicant advantage of these processes lies in the fact that the overall metals removal operation, even if repeated, does not unduly deleteriously alfect the activity, selectivity, pore structure and other desirable characteristics of the catalyst.
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 F. higher than the regeneration temperature, that is, the Vaverage temperature at which the major portion of carbon is removed from the catalyst. The temperature of treatment with molecular oxygencontaining gas will generally be in the range of about 1000 to 18007 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 l350 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 signiicant amount of carbon is present in the catalyst at the start of this high-temperature treatment, the essential oxygen contact is that continued aftervcarbon removal, which may vary from the short time necessary to produce an observable effect in the later treatment, say, a quarter of an hourV 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 stabilize a substantial amount of vanadium to its highest valence state, as evidenced by a signincant increase, say at least-about 10%, 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- Aenriched air, containing at least about 1%, preferably at least about 10% O2. The partial pressure of oxygen in the treating gas may range Widely, for example', from about 0.1 to 30 atmospheres, but usually the total gas pressure will not exceed about 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 Nos. 767,794, and Serial N0. 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 pH is frequently greater than about 7.5 and preferably the solution contains ammonium ions which may be NH4+ ions or organic-substituted NH4-l' ions such as methyl ammonium and quaternary hydrocarbon radical ammoniums. The amount of ammonium ion in the solution is suicient 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. vrequire pressurized equipment, the cost of which does not appear to be justified. Very short contact times, for example, about a minute, are satisfac tory, While the time of Washing may last 2 to 5 hours or longer. After the ammonium wash the catalyst slurry can be ltered 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 filter, 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 1000o 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 carbon or sulfur. 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 improvementin 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 weghtrpercent chlorine, based on the catalyst, together with one percent or more S2Cl2 gives good results.` Preferably, such a gas provides 1-10 percent C12 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 CCl4:5-10 parts C12` or HCl. Conveniently, a pressure of about 0-100 or more p.s.i.g.,
i3 preferably about -15 p.s.i.g. may be maintained in chlorination. The chlorination may take about to 120 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. Nicked removal may be accomplished by dissolving nickel compounds directly from the catalyst and/or by converting the nickel compounds to volatile materials 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 watersoluble 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 salt or other dispersible 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 nlckel, to Va water-dispersible form is described in copending application Serial No. 758,681 and may be accomplished, for instance, by 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 blo. 842,618, includes sulding 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.
v The suliiding step can be performed by contacting the poisoned catalyst with elemental sulfur vapors, or more vconveniently by contacting the poisoned catalyst with a volatile sulfide, such as H28, CS2 0r a mercaptan. The Acontact 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 partial pressure of about 0.1 to 30 atmospheres or more, preferably about 0.5 to 25 atmospheres. Hydrogen sulfide is the preferred suliiding 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 suliiding 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 approximately 1 atmosphere `or less seem near optimum for sulfiding 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 suliiding agent and the nature of the treating system, e.g., batch or continuous, as Well as the rate of diffusion within the catalyst matrix. The suliiding step performs the function 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 suliiding may be performed by a gaseous oxidizing agent to provide metal poisons in a dispersible form. Gaseous oxygen, or mixtures of gaseous oxygen with inert gases such as nitrogen, may be brought into contact with the sulded catalyst at an oxygen partial pressure of about 0.2 atmosphere 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 in the liquid aqueous oxidizing solution of sul-` furic 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 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, enough extra sulfuric or nitric acid may be used to provide one mole of sulfate or two moles of nitrate for each two moles of sodium or potassium.
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 to ten 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 cliromic acid where a small residual CrZOs 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 of oxidation can be selected as desired. The temperature can conveniently range up to about 220 F. with temperatures of above about F. being preferred. Temperatures above about 220 F. necessitate the use of superatmospheric pressures and no need for such has been found.
After provision of nickel sulfide in a dispersible form, the catalyst is washed with an aqueous medium to remove the metal poisons. This aqueous medium, for best removal of nickel is generally somewhat acidic. The aqueous medium can contain extraneous ingredients in trace amounts, so long as the medium is essentially water and the extraneous ingredients do not interfere with demetallization or adversely affect lthe 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 aqueous 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 sufiicient for nickel removal. Also, since a slightly acidic solution is desirable for nickel removal, this wash preferably takes place before the ammoni-um wash.
Alternative to the removal of poisoning metals =by procedures involving contact of the suliided or sulfated catalyst With aqueous media, nickel poison may be removed through conversion of the nickel sulfide to the volatile nickel carbonyl byy 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 olf the catalyst surface. Hydrogenation takes place at a temperature of about 800 to 1600 FV., at a pressure from atmospheric or less up to about 1000 p.s.i,g. with a vapor containing to 100% hydrogen. Preferred conditions are a pressure up to about p.s.i.g. and a temperature of about 1100 to 1300 F. and a hydrogen content greater than about 80 mole percent. The hydrogenation is continued until surface accumulations of poisoning metals, particularly nickel, are su-bstantially reduced to the elemental state. Carbonylation takes place at a temperature substantially lower than the hydrogenation, from about ambient temperature to 300 F. maximum and at a pressure up torabout 2000 p.s.i.g., with a gas containing about 50-100 mole percent CO. Preferred conditions include greater than about 90 mole percent CO, a pressure of up to about 800 p.s.i.g. and a temperature of about 10D-180 F. The CO treatment serves generally both to convert the elemental metals, especially nickel to volatile carbonyls and to remove the carbonyls.
After the ammonium Wash, or 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, as a slurry in its iinal aqueous treating medium. Where a large amount of catalyst inventory is treated, lest the water put out the iire or unduly lower the temperature in the regenerator, it may be desirable first to dry a Wet catalyst lter 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 ll00 F. may sometimes be disadvantageous. Calcination removes free water, if any is present, and 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 entrained in the catalyst or to purge the catalyst of reaction products.
The demetallization procedure of this invention has been found to be highly successful when used in conjunction with iiuidized catalytic cracking systems to control the amount of metal poisons on the catalyst. When such catalysts are processed, a uidized 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 suicient to effect a substantial con version 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 theV 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.
This invention will be better understood by reference to the accompanying drawing which shows the schematic -of a representative processing system but is not to be construed as limiting.
A residual feed contaminated with poisoning metals, for example, a vacuum asphalt, is fed by line 10 to hydrogia-nation ,unit 12, The structure of the unit will, of
course, depend upon the hydrogenation processes; for instance in the HDDC method, the unit 12 may consist of a coil or Aa coil and drum. In this method the feed is heated in the coil to about 700 to 1200a F. and thencontacted in the drum with a donor diluent, such as tetralin. The extent of conversion and the amount of metals removed will, of course, depend upon the severity of the hydrogenation and the characteristics of the original feed. In the alternative hydrogenation procedures, hydrogenation may be conducted under other conditions previously described. The hydrogen donor enters hydrogenation unit 12 by line 14. Line 1S is provided for the introduction of free hydrogen, if any is employed. Suitable pumps, not shown, may be provided to mix the residual feed with the hydrogen donor. VThe conditions in the unit 12, are adjusted for the results required. The total products are passed from the unit 12 to fractionator 16 by line 18. Fixed gases are removed by line 20 and gasoline and lighter components having an end boiling point of about 430 F. is removed by line 22 to storage or further treatment such as hydroforming. The hydrogen donor diluent may be removed from the fractionator by line 24 and hydrogenated by means, not shown, before being reintroduced into the system by line 14 with fresh hydrogen donor diluent, if needed, from line 25. Gas oils having a boiling point within the range of about 400 to 1505 F. may be removed by line 26 and carried directly to the catalytic cracker 28 by lines 30 and 32 or it may be removed from fractionator 16 along With the bottoms fraction, which usually has a boiling point essentially above about 700 F. by line 34 and conducted to -deasphalting tower 36, where the combined gas oil and bottoms fractions or the bottoms fraction alone may be contacted, preferably countercurrently, by an extraction solvent entering by line 38. If the gas oil fraction has been previously removed to the cracking unit 28, by lines 26 and 30, line 34 will convey substantially the entire bottoms fraction, boiling essentially above about 1050 F. to the deasphalting unit. The raiiinate, containing essentially heavier asphaltic constituents, is removed from the extraction tower by line 40 and the rainate may be conducted back to the hydrotreater by lines 40 and 10 for further processing or may be withdrawn from the system by line 41. The extract, containing essentially cracking components of reduced metal content, is removed from the extraction tower by line 42 and after removal of the solvent is conveyed to the cracker 28 by line 32. The cracking feed may be further diluted with low metals content conventional cracking stock from an outside source by lines 44 and 46 or recycle oil from the cracking procedure by lines 48 and 46. The cracker effluent leaves by line 50 and is brought to fractionator 52, where components of the efliuent are withdrawn by line 54 for fixed gases, line 56 for gasoline, line 58 for gas oil components and line 60 for materials higher boiling than gas oil. The latter components may be withdrawn from the system by line 62, or may be recycled by lines 64, 66, 40 and 10 to the hydrogenation zone 12, for further processing. The gas oil fraction is preferably recycled to the catalytic cracking step by lines 58, 48, 46 and 32 since it is substantially poison-free or it may be withdrawn for distillate fuel.
Contaminated catalyst is continuously removed from the cracker 28 by line 68 which conducts it to the regenerator 70. The regenerator is provided with the exit 72 for exhaust gases and with line 74 for the removal of regenerated catalyst and return to the cracker 28. A small slip stream of catalyst may be removed from line 74 for demetallization by line 76 and conveyed to demetallization unit 78. The demetallization unit 78 may comprise a system which includes apparatus (not shown), for example, for sulfiding, chlorinating, Washing and iiltering the catalyst. Alternatively, instead of chlorinating the sullided catalyst, means foroxidizing the sulded cata- 5 tionator.
alsa-45e 17 lyst may be substituted. The demetallization procedure used will, of course, depend upon the metals present. When, for instance, the feed. to be treated contains predominant amounts of vanadium, the demetallization treatment will be geared mainly toward the removal of vanadium Such as for example, dernetallization apparatus for high temperature treatment with molecular oxygen-containing gas as disclosed in SerialANo. 19,313, followed by y a basic Wash as disclosed in application Serial No. 39,810.
After a substantial portion of the metal contaminants are removed from the cracking catalyst, the catalyst is returned tothe cracking system by line 80.
The treating process of the present invention may be exemplified by the following:
A North Texas reduced crude having an API gravity of about 22, a carbon content of about 5.3 Weight percent, having an initial boiling point above about 650 F. and containing about 25 ppm. nickel and 60 ppm. of vanadium is mixed with a hydrogen donor diiuent comprising a hydrogenated catalytic cycle oil in a 1 to 1 donor diluent-to-oil ratio and thermally cracked. The thermal cracking is operated at a temperature of about 800 to 900 F. and under a pressure of about 100 psig. The feed rate is controlledA so that the blend or reduced crude and donor diluent is held for about 30 minutes at about 820 F. About 40% of the 1050 F .-l-components of the crude are converted to gas oil and lower boiling products. The total products are conducted to a fractionator where C4* gases and gasoline having an end boiling point of about 430 F. are taken olf. The combined gas oil and bottoms fraction having a metals content of about 12 ppm. nickel and' 20 ppm. vanadium is treated counter-currently with liquefied propane in a solvent-to-feed ratio of about 3.5 in an extraction tower. The rainate, containing material not suitable for catalytic cracking, is conveyed back to the hydrogenation zone for further processing. The solvent is removed from the extract, and then the extract, amounting to about 49% of the feed to the extraction tower, containing materials boiling above 400 F. and containing 5.9 ppm. nickel and 7.7 p.p.m. vanadium is diluted with a recycle gas oil from the cracker frac- The feed to the cracking unit contains about 3.9 p ptm. Ni and 5.1 ppm. V205. In the catalytic cracker the feed contacts a synthetic gel silica-alumina catalyst, having an A1203 content about 25%, at a temperature of about 950 to 975 F. and a pressure of about p.s'.i.g. The cracked products are introduced to a fractionator where a 65% yield of gasoline and other components are removed. The gas oil fraction is recycled to the cracker. A portion of the silica-alumina catalyst is continuously removed from the cracking reactor and brought to a regenerator. Average residence time in the regenerator is about 5 minutes at a temperature of about 1100 F. before catalyst return to the reactor at a carbon level of less than about 0.5%.
About per day of the cracking catalyst inventory, poisoned to a metals level of about 390 ppm. N10 and 1020 ppm. V205 is sent as a side stream from the regenerator to demetallization. In the demetallization process the catalyst is held in air for about an hour at about 1300 F. and then sent to a sulfiding zone where it is iluidized with H28 gas at a temperature of about 1175 F. for about 1 hour. Dilute nitric acid is brought in contact with the sulded catalyst and the slurry is aerated for about 10 minutes at a temperature of 200 F. to convert nickel poisons to dispersible form and remove them. The catalyst is then washed with an ammonium hydroxide solution having a pH of about 8 to 11, removing the available vanadium. The catalyst, substantially reduced in nickel and vanadium content is filtered from the wash slurry, dried at about 350 F. and returned to the regenerator. The treated catalyst analyzes a metals content of 155 ppm. nickel and 745 ppm. vanadium.
In another example, an asphaltic residual oil from the vacuum distillation of the bottoms from an atmospheric distillation of petroleum crude oil is contacted with tetralin in a ratio of diluent to feed of 1 to 1,V at about 800 F. and a pressure range of about 1160 to 1260 p.s.i.g. and a hydrogen partial pressure of 500 p.s.i.g. inthe presence of a catalyst comprising about 2% I2. The residual oil has a metals content ofy 82 ppm. Ni0 and 244 ppm. V205. 30% of the feed is converted into products boiling below about 950 F. The tota-l products are conducted to afractionator where about 7% lof gas. and gasoline are removed, and about 23% of `gas oil is recovered and sent to the cracking unit. The bottoms fraction, boiling essentially above about 950 F. and having a vmetals content of 40 ppm. nickel and 81 p.p.m. vanadium, is treated countercurrently with propane in a solvent-to-oil ratio of 7 and propane is removed from the extract phase, resulting in a yield of 52.7 percent -by Weight extract having a metals content of 6.8 p pm. nickel and 15.4 ppm. vanadium. The extract is diluted with a recycle gas oil and a conventional cracking feed to give a cracking feed containing 1.5 p.p.m. nickel and 2.4 ppm. vanadium.
In the catalytic cracker the feed contacts a syntheticgel silica-alumina catalyst, having an A1203 content of about 25%, `at a temperature of about 950 to 975 F. and a pressure of about 5 p.s.i.g. The cracked products are introduced to a fractionator Where a 70% yield of gasoline and other components are removed. TheA gas oil is recycled to the cracker for fur-ther processing. A Vportion of the silica-alumina catalyst is continuously removed from the cracking reactor and brought to a regenerator. Average residence time in theV regenerator is about 5 minutes at a temperature of about 1100 F. before returning to the reactor at a carbon level of less than about 0.4%
About 10% per day of the cracking catalyst inventory poisoned to a metals level of about 215 ppm. .NiO' and 650 p.p.m. vanadium is sent as a side stream from the regenerator to demetallization. In the demetallization process the catalyst is held in air for about an hour at about 1300 F. and then sent to a suliiding zone where it is fluidi'zed with H2S gas at a temperature of about 1175 F. for about 1 hour. The catalyst is then purged with flue gas at a temperature of about 575 F. and chlorinated in a chlorination zone with an equilmolar mixture of C12 and CCL, at about 600 F. After about 1 hour no trace of vanadium chloride can be found 'in the chlorination eluent and the catalyst is quickly washed with Water. A pH of about 2 is imparted to this Wash medium by chlorine entrained in the catalyst and the wash serves to remove nickel chloride. The demetallization procedure removes about 60% ofthe nickel and about 25% of the vanadium on the catalyst.
A third run was conducted with a West Texas asphalt having an initial boiling point of about 950 lF., a specific gravity of 1.002, a Conradson carbon content ofl about 20.5 and a metals level of 76 ppm. nickel-oxide and v110 p.p.m. vanadium oxide. The feed was diluted `with tetralin in a ratio of 1 to 1 and heated to about 700 F. in the presence of free ,hydrogen and under a partial pressure of about 1000 p.s.i.g. and 5 `percent of iodine by weight of asphalt present. Cracking lwas held to a minimum and only about 5% of the total feed was converted to products boiling below about 950 F. Hydrogen consumption was about 437 sci/bbl. of feed. The hydrogenated product, analyzing about 5.5 ppm. nickel oxide and 10.2 ppm. vanadium oxide, is deasphalted in a solvent-to-oil ratio of 5 with propane and the propane-free extract, amounting to about of the hydrogenated product, analyzes 1.5 ppm. nickel oxide and 3.3 p;p.rn. vanadium oxide. The extract is catalytically cracked at a temperature of about 950 F., at 10 p.s.i.g. pressure in the presence of a synthetic-gel silica-alumina catalyst containing about 13% Al203. The cracked products are introduced to a fractionator where a 60% yield of gasoline and other low boiling components are removed. The products boiling above about 950 F., are recycledto the hydrotreater for further processing. A portion of the vvat about 350 F. and returned to the regenerator.
` silica-alumina catalyst is continuously removed from the cracking reactor and brought to a regenerator. Average residence time in the regenerator is about 5 minutes at a temperature of about 1100 F., before returning to the reactor at a carbon level of about 0.5%.
About 15% of the cracking catalyst inventory, poisoned to a metals level of about 600 p.p.m. nickel and 890 p.p.m. vanadium, is each day sent as a side stream from the regenerator `to demetallization. In the demetallization process the catalyst is held in air for about an hour at about l300 F. and then sent to a basic wash with an ammonium hydroxide solution having a pH of about 8 to ll. rlhe catalyst is filtered from the wash slurry, dried The treated catalyst is analyzed and shows a metals content `of 540 p.p.m. nickel and 650 p.p.m. vanadium.
Thus, this invention provides for overcoming poisoning elects by a balanced process which includes hydrogenavtion, solvent deasphalting and cracking catalyst demetallization. v
It is claimed: 1. A process for treating a residual hydrocarbon oil vboiling above the gasolineV range and containing at least to 10 volumes of diluent per volume of residual hydro-4 carbon oil; at a temperature in the range of about 700- 1200 F. and a pressure from about atmospheric to about 200 p.s.i.g. thereby reducing partially the content of said contaminating metals by about 10 to 90% and increasing the hydrogen-to-carbon ratio of the hydrogenated effluent, subjecting resulting hydrogenated bottoms boiling primarilyr above about 950 F. and containing at least about 2 parts per million nickel and at least about 3 parts per million vanadium to solvent deasphalting to remove about 30 to 95% of said metal contaminants and produce an extract and raflinate, said extract containing about 40 to 95% of said hydrogenated bottoms, subjecting a hydrocarbon oil boiling above the gasoline range and containing suicient of said extract to provide at least about v1 part per million of nickel and at least about l part per million vanadium, to contact with a solid cracking catalyst under cracking conditions to produce gasoline, removing metal-contaminated catalyst from the cracking system, said removed catalyst containing at least about 50 parts per million nickel-and at least about 50 parts per million vanadium, demetallizing the removed catalyst to reduce its nickel and vanadium content by about l to 90%, returning demetallized catalyst to said cracking systern and recovering products from said cracking.
.2. The method of claim l in which the residual oil contains at least about l0 parts per million of each of nickel and vanadium and the hydrogenation reduces the amount of contaminating metals by about 65 to 90% and the catalyst is a'synthetic gel, silica-alumina catalyst.
3. The process of claim l wherein hydrogenaton is performed in the presence of about l to 10% of a hydride of a halogen having an atomic number of 35 to 53.
4. The method of claim 1 wherein the catalyst demetallization includes contact of the catalyst with a vapor reactive with a metal contaminant.
5. A process for treating a residual hydrocarbon oil boiling above the gasoline range and containing at least about 5 parts per million of each of nickel and vanadium as contaminating metals, comprising the steps of contacting said hydrocarbon oil in a hydrogenation zone with about 1 to 10 percent by weight of the residual of hydride of a halogen having an atomic number of 35 to 53; and a hydrogen donor diluent in the proportions of about 0.1 to volumes of diluent per volume of residual hy- =-Hrocarbon oil at a temperature of about 400-700 F. and
a pressure of about 0 to 2,000I p.s.i.g., thereby reducing partially the content of said contaminating metals by 'as about 10 to 90% and increasing the hydrogen-to-carbon ratio of the hydrogenated etlluent, without substantial cracking, subjecting resulting hydrogenated bottoms containing at least about 2 parts per million nickel and at least about 3 parts per million vanadium and boiling essentially above about 400 F., to solvent deasphalting to remove about 30 to 95% of said metal contaminants and produce an extract and rainate, said extract contairung about 40 to 95 of said hydrogenated bottoms, subjecting a hydrocarbon oil boiling above the gasoline range and containing suiiicient of said extract to provide at least about 1 part per million of nickel and at least about 1 part per million vanadium, to contact with a solid cracking catalyst under cracking conditions to produce gasoline, removing metal-contaminated catalyst from the cracking system, said removed catalyst containing at least about 50 parts per million nickel and at least about 50 parts per million vanadium, demetallizing the removed catalyst to reduce its nickel and vanadium content by about 10 to 90%, returning demetallized catalyst to said cracking system and recovering products from said crackg6. The method of claim 5 in which the residual oil contains at least about 10 parts per million of each of nickel and vanadium and the hydrogenation reduces the amount ot contaminating metals by about 65 to 90%.
7. The process of claim 5 wherein catalyst demetallization includes contact of the catalyst with a vapor reactive with a metal contaminant.
8. The process of claim 5 wherein the solid cracking catalyst is a synthetic gel silica-based cracking catalyst.
9. A process for treating a residual hydrocarbon oil boiling above the gasoline range and containing at least about 5 parts per million of each of nickel and vanadium as contaminating metals, comprising treating said hydrocarbon oil in a hydrogenation Zone with about l0() to 10,000 standard cubic feet of hydrogen per barrel of oil at a temperature of about 750 to 900 F. and a pressure of about 1500 to 2500 p.s.i.g. in the presence of a solid hydrogenation catalyst, thereby reducing partially the -content of said contaminating metals by about 10 to 90% and increasing the hydrogen-to-carbon ratio of the hydrogenated eluent, subjecting resulting hydrogenated bottoms containing at least about 2 parts per million nickel and at least about 3 parts per million vanadium and boil ing essentially above about 400 F., to solvent deasphalting to remove about 3,0 to of said metal contaminants and produce an extract and rainate, said extract containing about 40 to 95 of said hydrogenated bottoms, subjecting a hydrocarbon oil boiling above the gasoline range and containing suiiicient of said extract to provide at least about 1 part per million of nickel and at least about 1 part per million vanadium, to Contact with a solid crack-A ing catalyst under cracking conditions to produce gasoline, removing metal-contaminated catalyst from the' cracking system, said removed catalyst containing at least'- about 50 parts per million nickel and at least about 50` parts per million vanadium, demetallizing the removed catalyst to reduce its nickel and vanadium content by about 10 to 90%, returning demetallized catalyst to said cracking system and recovering products from said cracking.
l0. The method of claim 9 in which the residual oil contains at least about 10 parts per million of each of nickel and vanadium yand the hydrogenation reduces the amount of contaminating metal by about 65 to 90%.
11. The method of claim 9 wherein the catalyst demetallization includes contact of the catalyst with a vapor reactive with a metal contaminant.
l2. The process of claim 9 wherein the solid cracking catalyst is a synthetic gel, silica-based catalyst.
13. The method of claim l2 wherein the cracking catalyst is regenerated at about 950 F. to l200 F. and demetallized by contacting Ithe regenerated catalyst with a molecular oxygen-containing gas at a temperature of about 1000 to about 1800 F., and at least about 50 F. higher than the regeneration temperature, sulliding the poisoning metal component on the catalyst by contact with a suliding agent at a temperature of about 500 to 1500 F., chlorinating poisoning metal containing component on the Catalyst by Contact with an essentially anhydrous chlorinating agent at a temperature of about 300 to 1000 F., and contacting the catalyst with a liquid, essentially aqueous medium to remove soluble poisoning metal chloride from the catalyst.
References Cited by the Examiner UNITED STATES PATENTS Houdry 208-119 Pier et al 1 208-108 Shabaker et al. 208-120 Snyder 208-120 Snyder 208-113 Blue et al. 208-145 Conneil et al 208-86 Scott et al 208-61 ALPHONSO D, SULLIVAN, Primary Examiner.
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|U.S. Classification||208/57, 208/61, 208/96, 208/120.1, 208/56|
|International Classification||C10G67/00, C10G67/04|