EP0065627B1 - Immobilisation de vanadium déposé sur des matériaux catalytiques pendant la conversion d'huiles contenant des métaux lourds et des précurseurs de coke - Google Patents

Immobilisation de vanadium déposé sur des matériaux catalytiques pendant la conversion d'huiles contenant des métaux lourds et des précurseurs de coke Download PDF

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Publication number
EP0065627B1
EP0065627B1 EP82101794A EP82101794A EP0065627B1 EP 0065627 B1 EP0065627 B1 EP 0065627B1 EP 82101794 A EP82101794 A EP 82101794A EP 82101794 A EP82101794 A EP 82101794A EP 0065627 B1 EP0065627 B1 EP 0065627B1
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EP
European Patent Office
Prior art keywords
catalyst
feed
coke
vanadium
conversion
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EP82101794A
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German (de)
English (en)
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EP0065627A2 (fr
EP0065627A3 (en
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William P. Hettinger, Jr.
James D. Carruthers
William D. Watkins
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Ashland LLC
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Ashland Oil Inc
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Priority claimed from US06/258,265 external-priority patent/US4377470A/en
Application filed by Ashland Oil Inc filed Critical Ashland Oil Inc
Priority to AT82101794T priority Critical patent/ATE19411T1/de
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Publication of EP0065627A3 publication Critical patent/EP0065627A3/en
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    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G11/00—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils
    • C10G11/14—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils with preheated moving solid catalysts
    • C10G11/18—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils with preheated moving solid catalysts according to the "fluidised-bed" technique
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/10—Feedstock materials
    • C10G2300/107—Atmospheric residues having a boiling point of at least about 538 °C

Definitions

  • This invention relates to processes for converting heavy hydrocarbon oils into lighter fractions, and especially to processes for converting heavy hydrocarbons containing high concentrations of coke precursors and heavy metals into gasoline and other hydrocarbon fuels.
  • VGO vacuum gas oils
  • the catalysts employed in early homogenous fluid dense beds were of an amorphous siliceous material, prepared synthetically or from naturally occurring materials activated by acid leaching.
  • Tremendous strides were made in the 1950's in FCC technology in the areas of metallurgy, processing equipment, regeneration and new more- active and more stable amorphous catalysts.
  • increasing demand with respect to quantity of gasoline and increased octane number requirements to satisfy the new high horsepower-high compression engines being promoted by the auto industry put extreme pressure on the petroleum industry to increase FCC capacity and severity of operation.
  • the new catalyst developments revolved around the development of various zeolites such as synthetic types X and Y and naturally occurring faujasites; increased thermal-steam (hydro- thermal) stability of zeolites through the inclusion of rare earth ions or ammonium ions via ionexchange techniques; and the development of more attrition resistant matrices for supporting the zeolites.
  • These zeolitic catalyst developments gave the petroleum industry the capability of greatly increasing throughput of feedstock with increased conversion and selectivity while employing the same units without expansion and without requiring new unit construction.
  • the effect of increased Conradson carbon is to increase that portion of the feedstock converted to coke deposited on the catalyst.
  • the amount of coke .deposited on the catalyst averages around about 4-5 wt.% of the feed.
  • This coke production has been attributed to four different coking mechanisms, namely, contaminant coke from adverse reactions caused by metal deposits, catalytic coke caused by acid site cracking, entrained hydrocarbons resulting from pore structure adsorption and/or poor stripping, and Conradson carbon resulting from pyrolytic distillation of hydrocarbons in the conversion zone.
  • the coked catalyst is brought back to equilibrium activity by burning off the deactivating coke in a regeneration zone in the presence of air, and the regenerated catalyst is recycled back to the reaction zone.
  • the heat generated during regeneration is removed by the catalyst and carried to the reaction zone for vaporization of the feed and to provide heat for the endothermic cracking reaction.
  • the temperature in the regenerator is normally limited because of metallurgical limitations and the hydrothermal stability of the catalyst.
  • the hydrothermal stability of the zeolite-containing catalyst is determined by the temperature and steam partial pressure at which the zeolite begins to rapidly lose its crystalline structure to yield a low-activity amorphous material.
  • the presence of steam is highly critical and is generated by the burning of adsorbed and absorbed (sorbed) carbonaceous material which has a significant hydrogen content (hydrogen to carbon atomic ratios generally greater than about 0.5).
  • This carbonaceous material is principally the high-boiling sorbed hydrocarbons with boiling points as high as 1500-1700°F (815-927°C) or above that have a modest hydrogen content and the high boiling nitrogen containing hydrocarbons, as well as related porphyrins and asphaltenes.
  • the high molecular weight nitrogen compounds usually boil above 1025°F (552°C) and may be either basic or acidic in nature.
  • the basic nitrogen compounds may neutralize acid sites while those that are more acidic may be attracted to metal sites on the catalyst.
  • the porphyrins and asphaltenes also generally boil above 1025°F (552°C) and may contain elements other than carbon and hydrogen.
  • the term "heavy hydrocarbons" includes all carbon and hydrogen compounds that do not boil below about 1025°F (552°C), regardless of the presence of other elements in the compound.
  • the heavy metals in the feed are generally present as porphyrins and/or asphaltenes.
  • certain of these metals, particularly iron and copper, may be present as the free metal or as inorganic compounds resulting from either corrosion of process equipment or contaminants from other refining processes.
  • the metal-containing fractions of reduced crudes contain Ni-V-Fe-Cu in the form of porphyrins and asphaltenes. These metal-containing hydrocarbons are deposited on the catalyst during processing and are cracked in the riser to deposit the metal or are carried over by the coked catalyst as the metallo-porphyrin or asphaltene and converted to the metal oxide during regeneration.
  • the adverse effects of these metals as taught in the literature are to cause non-selective or degradative cracking and dehydrogenation to produce increased amounts of coke and light gases such as hydrogen, methane and ethane. These mechanisms adversely affect selectivity, resulting in poor yields and quality of gasoline and light cycle oil.
  • vanadium is especially detrimental to catalyst life.
  • the vanadium deposited on the catalyst under the reducing conditions in the riser is in an oxidation state less than +5.
  • free hydrogen and reactive hydrogen-compounds e.g. CH4, C2H6 and C3H8 are present which give rise to such reducing conditions.
  • the vanadium on the catalyst is converted to vanadium oxides, in particular vanadium pentoxide.
  • the vanadium pentoxide has a melting point lower than temperatures encountered in the regeneration zone, and it can become a mobile liquid, flowing across the catalyst surface and plugging pores.
  • This vanadia may also enter the zeolite structure, neutralizing the acid sites and, more significantly, irreversibly destroying the crystalline aluminosilicate structure and forming amorphous material.
  • this molten vanadia can, at high vanadia levels, especially for catalyst materials having a low surface area, coat the catalyst microspheres and thereby coalesce particles which adversely affects their fluidization.
  • a process for converting a vanadium- containing hydrocarbon oil feed to ' lighter products comprising the steps of contacting said oil feed under conversion conditions with a cracking catalyst to form lighter products and coke, whereby vanadium in an oxidation state less than +5 is deposited on said catalyst together with coke.
  • the lighter products are separated from the spent catalyst and the catalyst is regenerated by contacting it with an oxygen-containing gas under conditions whereby said coke is burned forming CO and C0 2 and said vanadium is maintained in an oxidation state less than +5.
  • This invention by retaining vanadium in an oxidation state wherein the vanadium has a high melting point, permits the recycle of catalyst to levels of vanadium as high as 10,000 ppm, or even 20,000 ppm or 50,000 ppm.
  • the adverse effects such as clumping of the catalyst and pore closings brought about by molten pentavalent vanadium, are thus avoided.
  • the catalyst can withstand a much higher vanadium loading than previously experienced the amount of make-up catalyst is reduced.
  • Figs. 1 and 2 are schematic designs of catalyst regeneration and associated cracking apparatus which may be used in carrying out this invention.
  • the invention may be carried out by controlling the regeneration of the spent, vanadium- containing catalyst using several methods, alone or in combination.
  • the objective of these methods is to retain vanadium in a low oxidation state, either by not exposing the vanadium to oxidizing conditions, or by exposing vanadium to oxidizing conditions for too short a time to oxidize a significant amount of vanadium to the +5 state.
  • the concentration of vanadium on the catalyst particles increases as the catalyst is recycled, and the vanadium on the catalyst introduced into the reactor becomes coated with coke formed in the reactor.
  • the generator conditions are selected to ensure that the concentration of coke is retained at least a minimum level on the catalyst. This coke may serve either to ensure a reducing environment for the vanadium or to provide a barrier to the movement of oxidizing gas to underlying vanadium.
  • the concentration of coke on the catalyst particles is at least about 0.05 percent and the preferred coke concentration is at least about 0.15 percent.
  • the regeneration is carried out in an environment which is non- oxidizing for the vanadium in an oxidation state less than +5. This may be accomplished by adding reducing gases such as, for example, CO or ammonia to the regenerator, or by regenerating under oxygen-deficient conditions. Oxygen-deficient regeneration increases the ratio of CO to C0 2 and in this method of providing a non- oxidizing atmosphere the CO/C0 2 ratio is at least about 0.25, preferably is at least about 0.3, and most preferably is at least about 0.4.
  • the CO/CO 2 ratio may be controlled by controlling the extent of oxygen deficiency within the regenerator.
  • the CO/C0 2 ratio may be increased by providing chlorine in an oxidizing atmosphere within the regenerator, the concentration of chlorine preferably being from about 100 to 400 ppm.
  • This method of increasing the CO/CO 2 ratio is disclosed in U.S. Patent 4,376,696 for "Addition of MgCI 2 to Catalyst" and U.S. Patent 4,375,404 for "Addition of Chlorine to Regenerator", both in the name of George D. Myers.
  • the use of a reducing atmosphere within the regenerator is especially useful in combusting coke in zones where the coke level approaches or is reduced below about 0.05 percent, and it is preferred to have a CO/CO 2 ratio of at least about 0.25 in zones where the coke loading is less than about 0.05 percent by weight.
  • a riser regenerator is used as one stage in a multi-stage regenerator to contact the catalyst with an oxidizing atmosphere for a short period of time, such as for example less than about two seconds and preferably less than about one second.
  • the riser stage of the regenerator has the advantage in reducing the carbon concentration to a level less than about 0.15 percent or less than about 0.05 percent, that vanadium, which is no longer protected by a coating of carbon, may not be in an oxidizing atmosphere for a long enough time to form molten +5 vanadium.
  • the low density of the particles in a riser-regenerator minimizes coalescence of those particles which may have liquid pentavalent vanadia on their surfaces.
  • the particles are contacted with a reducing atmosphere, such as one containing CO or other reducing gas, after leaving the riser.
  • a reducing atmosphere such as one containing CO or other reducing gas
  • the particles may then be accumulated, as for example, in a settled bed, before being recycled to contact additional fresh feed.
  • the catalyst particles to be accumulated are contacted with a reducing atmosphere, preferably immediately after leaving the riser and before accumulating in a dense bed of regenerated particles, and in the preferred method of carrying out this process the particles are retained in a reducing atmosphere within such dense bed, and in the most preferred method a reducing atmosphere is provided for the particles until about the time they are contacted with fresh feed.
  • the preferred riser regenerator is similar to the vented riser reactor as is disclosed in U.S. Patents 4,066,533 and 4,070,159 to Myers et al which achieves ballistic separation of gaseous products from catalyst.
  • This apparatus has the advantages of achieving virtually instantaneous separation of the regenerated catalyst, now containing some vanadia to which any oxygen present would have access, from the oxidizing atmosphere.
  • the catalyst is contacted with a reducing atmosphere, preferably immediately after its separation from the oxidizing atmosphere and most preferably also in collection zones for the regenerated catalyst.
  • This invention may be used in processing any hydrocarbon feed containing a significant concentration of vanadium, and FCC as well as RCC processes are contemplated. It is, however, especially useful in processing reduced crudes having high metal and high Conradson carbon values, and the invention will be described in detail with respect to its use in processing an RCC feed.
  • the carbo-metallic feed comprises or is composed of oil which boils above about 343°C.
  • oil or at least the 343°C+ portion thereof, is characterized by a heavy metal content of at least about 4, preferably more than about 5, and most preferably at least about 5.5 ppm of Nickel Equivalents by weight and by a carbon residue on pyrolysis of at least about 1% and more preferably at least about 2% by weight.
  • the carbo-metallic feed in the form of a pumpable liquid, is brought into contact with hot conversion catalyst in a weight ratio of catalyst to feed in the range of 3 to 18 and preferably more than about 6.
  • the feed in said mixture undergoes a conversion step which includes cracking while the mixture of feed and catalyst is flowing through a progressive flow type reactor.
  • the feed, catalyst, and other materials may be introduced at one or more points.
  • the reactor includes an elongated reaction chamber which is at least partly vertical or inclined and in which the feed material, resultant products and catalyst are maintained in contact with one another while flowing as a dilute phase or stream for a predetermined riser residence time in the range of 0.5 to 10 seconds.
  • the reaction is conducted at a temperature of 482°C to 760°C, measured at the reaction chamber exit, under a total pressure of .68 to 3.4 ATM under conditions sufficiently severe to provide a conversion per pass in the range of about 50% or more and to lay down coke on the catalyst in an amount in the range of 0.3 to 3% by weight and preferably at least about 0.5%.
  • the overall rate of coke production, based on weight of fresh feed, is in the range of 4 to 14% by weight.
  • the catalyst is separated from the products, is stripped to remove high boiling components and other entrained or adsorbed hydrocarbons and is then regenerated with oxygen-containing combustion-supporting gas under conditions of time, temperature and atmosphere sufficient to reduce the carbon on the regenerator catalyst to about 0.25% or less.
  • the process may be operated without added hydrogen in the reaction chamber. If desired, and preferably, the process may be operated without prior hydrotreating of the feed and/or without other process of removal of asphaltenes of metals from the feed, and this is true even where the carbo-metallic oil as a whole contains more than about 4, or more than about 5 or even more than about 5.5 ppm Nickel Equivalents by weight of heavy metal and has a carbon residue on pyrolysis greater than about 1%, greater than about 1.4% or greater than about 2% by weight. Moreover, all of the converter feed, as above described, may be cracked in one and the same conversion chamber.
  • the cracking reaction may be carried out with a catalyst which has previously been used (recycled, except for such replacement as required to compensate for normal losses and deactivation) to crack a carbo-metallic feed under the above described conditions.
  • Heavy hydrocarbons not cracked to gasoline in a first pass may be recycled with or without hydrotreating for further cracking in contact with the same kind of feed in which they were first subjected to cracking conditions, and under the same kind of conditions; but operation in a substantially once-through or single pass mode (e.g. less than about 15% by volume of recycle based on volume of fresh feed) is preferred.
  • the catalyst is projected in a direction established by the elongated reaction chamber or an extension thereof, while the products, having lesser momentum, are caused to make an abrupt change of direction, resulting in an abrupt, substantially instantaneous ballistic separation of products from catalyst.
  • the thus separated catalyst is then stripped, regenerated and recycled to the reactor as above described.
  • the converter feed contains 343°C+ material which has not been hydrotreated and is characterized in part by containing at least about 5.5 parts per million of nickel equivalents of heavy metals.
  • the converter feed is brought together not only with the above mentioned cracking catalyst, but also with additional gaseous material including steam whereby the resultant suspension of catalyst and feed also includes gaseous material wherein the ratio of the partial pressure of the added gaseous material relative to that of the feed is in the range of 0.25 to 4.0.
  • the vapor residence time is in the range of 0.5 to 3 seconds when practicing this embodiment or aspect of the invention.
  • This preferred embodiment or aspect and the one referred to in the preceding paragraph may be used in combination with one another or separately.
  • the carbo-metallic feed is not only brought into contact with the catalyst, but also with one or more additional materials including particularly liquid water in a weight ratio relative to feed ranging from 0.04 to 0.25, more preferably 0.04 to 0.2 and still more preferably 0.05 to 0.15.
  • additional materials including the liquid water, may be brought into admixture with the feed prior to, during or after mixing the feed with the aforementioned catalyst, and either after or, preferably, before, vaporization of the feed.
  • the feed, catalyst and water e.g.
  • the progressive flow type reactor which may or may not be a reactor embodying the above described ballistic separation, at one or more points along the reactor. While the mixture of feed, catalyst and steam produced by vaporization of the liquid water flows through the reactor, the feed undergoes the above mentioned conversion step which includes cracking. The feed material, catalyst, steam and resultant products are maintained in contact with one another in the above mentioned elongated reaction chamber while flowing as a dilute phase or stream for the above mentioned predetermined riser residence time which is in the range of 0.5 to 10 seconds.
  • the present invention provides a process for the continuous catalytic conversion of a wide variety of carbo-metallic oils to lower molecular weight products, while maximizing production of highly valuable liquid products, and making it possible, if desired, to avoid vacuum distillation and other expensive treatments such as hydrotreating.
  • the invention may be applied to the processing of such widely diverse materials as heavy bottoms from crude oil, heavy bitumen crude oil, those crude oils known as "heavy crude” which approximate the properties of reduced crude, shale oil, tar sand extract, products from coal liquification and solvated coal, atmospheric and vacuum reduced crude, extracts and/or bottoms (raffinate) from solvent de- asphalting, aromatic extract from lube oil refining, tar bottoms, heavy cycle oil, slop oil, other refinery waste streams and mixtures of the foregoing.
  • Such mixtures can for instance be prepared by mixing available hydrocarbon fractions, including oils, tars, pitches and the like.
  • powdered coal may be suspended in the carbo-metallic oil.
  • demetalated oils may be converted using the invention; but it is an advantage of the invention that it can employ as feedstock carbo-metallic oils that have had no prior demetalation treatment.
  • the invention can be applied to hydrotreated feedstocks; but it is an advantage of the invention that it can successfully convert carbo-metallic oils which have had substantially no prior hydrotreatment.
  • the preferred application of the proces is to reduced crude, i.e., that fraction of crude oil boiling at and above 343°C, alone or in admixture with virgin gas oils.
  • carbo-metallic oil feedstock at least about 70%, more preferably at least about 85% and still more preferably about 100% (by volume) of which boils at and above about 343°C. All boiling temperatures herein are based on standard atmospheric pressure conditions.
  • carbo-metallic oil partly orwholly composed of material which boils at and above about 343°C, such material is referred to herein as 343°C+ material; and 343°C+ material which is part of or has been separated from an oil containing components boiling above and below 343°C may be referred to as a 343°C+ fraction.
  • carbo-metallic oils contemplated by the invention may contain material which may not boil under any conditions; for example, certain asphalts and asphaltenes may crack thermally during distillation, apparently without boiling.
  • the feed comprises at least about 70% by volume of material which boils above about 343°C, it should be understood that the 70% in question may include some material which will not boil or volatilize at any temperature.
  • these non-boilable materials when present may frequently or form the most part be concentrated in portions of the feed which do not boil below about 538°C, 552°C or higher.
  • the contemplated feeds, or at least the 343°C+ material therein have a carbon residue on pyrolysis of at least about 2 or greater.
  • the Conradson carbon content may be in the range of about 2 to about 12 and most frequently at least about 4. A particularly common range is 4 to 8.
  • the feed has an average composition characterized by an atomic hydrogen to carbon ratio in the range of 1.2 to 1.9, and preferably 1.3 to 1.8.
  • the carbo-metallic feeds employed in accordance with the invention, or at least the 343°C+ material therein, may contain at least about 4 parts per million of Nickel Equivalents, as defined above, of which at least about 01. ppm is vanadium.
  • Carbo-metallic oils within the above range can be prepared from mixtures of two or more oils, some of which do and some of which do not contain the quantities of Nickel Equivalents and vanadium set forth above. It should also be noted that the above values for Nickel Equivalents and nickel represent time-weighted averages for a substantial period of operation of the conversion unit, such as one month, for example.
  • the carbo-metallic oils useful in the invention may and usually do contain significant quantities of compounds containing nitrogen, a substantial portion of which may be basic nitrogen.
  • the total nitrogen content of the carbo-metallic oils may be at least about 0.05% by weight. Since cracking catalysts owe their cracking activity to acid sites on the catalyst surface or in its pores, basic nitrogen-containing compounds may temporarily neutralize these sites, poisoning the catalyst. However, the catalyst is not permanently damaged since the nitrogen can be burned off the catalyst during regeneration, as a result of which the acidity of the active sites is restored.
  • the carbo-metallic oils may also include significant quantities of pentane insolubles, for example at least about 0.5% by weight,.and more typically 2% or more or even about 4% or more. These may include for instance asphaltenes and other materials.
  • Alkali and alkaline earth metals generally do not tend to vaporize in large quantities under the distillation conditions employed in distilling crude oil to prepare the vacuum gas oils normally used as FCC feedstocks. Rather, these metals remain for the most part in the "bottoms" fraction (the non-vaporized high boiling portion) which may for instance be used in the production of asphalt or other by-products.
  • bottoms the non-vaporized high boiling portion
  • reduced crude and other carbo-metallic oils are in many cases bottoms products, and therefore may contain significant quantities of alkali and alkaline earth metals such as sodium. These metals deposit upon the catalyst during cracking.
  • these metals may undergo interactions and reactions with the catalyst (including the catalyst support) which are not normally experienced in processing VGO under conventional FCC processing conditions. If the catalyst characteristics and regeneration conditions so require, one will of course take the necessary precautions to limit the amounts of alkali and alkaline earth metal in the feed, which metals may enter the feed not only as brine associated with the crude oil in its natural state, but also as components of water or steam which are supplied to the cracking unit. Thus, careful desalting of the crude used to prepare the carbo-metallic feed may be important when the catalyst is particularly susceptible to alkali and alkaline earth metals.
  • the content of such metals (hereinafter collectively referred to as "sodium") in the feed can be maintained at about 1 ppm or less, based on the weight of the feedstock.
  • the sodium level of the feed may be keyed to that of the catalyst, so as to maintain the sodium level of the catalyst which is in use substantially the same as or less than that of the replacement catalyst which is charged to the unit.
  • the carbo-metallic oil feedstock constitutes at least about 70% by volume of material which boils above about 343°C, and at least about 10% of the material which boils above about 343°C will not boil below about 552°C.
  • the average composition of this 343°C material may be further characterized by: (a) an atomic hydrogen to carbon ratio in the range of 1.3 to 1.8; (b) a Conradson carbon value of at least about 2; (c) at least about four parts per million of Nickel Equivalents, as defined above, of which at least about two parts per million is nickel (as metal, by weight), at least about 0.1 parts per million vanadium; and (d) at least one of the following: (i) at least about 0.3% by weight of sulfur, (ii) at least about 0.05% by weight of nitrogen; and (iii) at least about 0.5% by weight of pentane insolubles.
  • the preferred feed will include all of (i), (ii) and (iii), the other components found in oils of petroleum and non-petroleum origin may also be present in varying quantities providing they do not prevent operation of the process.
  • the present invention has the definite advantage that it can successfully produce large conversions and very substantial yields of liquid hydrocarbon fuels from carbo-metallic oils which have not been subjected to any substantial amount of cracking.
  • the carbo-metallic feed introduced into the present process is oil which has not previously been contacted with cracking catalyst under cracking conditions.
  • the process of the invention is suitable for operation in a substantially once-through or single pass mode.
  • the volume of recycle, if any, based on the volume of fresh feed is preferably about 15% or less and more preferably about 10% or less.
  • the weight ratio of catalyst to fresh feed (feed which has not previously been exposed to cracking catalyst under cracking conditions) used in the process is in the range of 3 to 18.
  • Preferred and more preferred ratios are 4 to 12, more preferably 5 to 10 and still more preferably 6 to 10, a ratio of about 10 presently being considered most nearly optimum.
  • controlling the catalyst to oil ratio at relatively low levels within the aforesaid ranges tends to reduce the coke yield of the process, based on fresh feed.
  • the ratio between the number of barrels per day of plant through-put and the total number of tons of catalysts undergoing circulation throughout all phases of the process can vary widely.
  • daily plant through- put is defined as the number of M 3 of fresh feed boiling above about 343°C which that plant processes per average day of operation to liquid products boiling below about 221°C.
  • 46 to 68 Kg M 3 of catalyst are under circulation in the process per 1000 M 3 per day of plant through-put.
  • this ratio is in the range of about 2 to 3.
  • While the present invention may be practiced in the range of 11 Kg to 171 Kg and more typically 11 Kg to 68 Kg of catalyst inventory per 1000 M 3 of daily plant through-put, it is preferred to carry out the process of the present invention with a very small ratio of catalyst weight to daily plant through-put. More specifically, it is preferred to carry out the process of the present invention with an inventory of catalyst that is sufficient to contact the feed for the desired residence time in the above indicated catalyst to oil ratio while minimizing the amount of catalyst inventory, relative to plant through- put, which is undergoing circulation or being held for treatment in other phases of the process such as, for example, stripping, regeneration and the like. Thus, more particularly, it is preferred to carry out the process of the present invention with 11 to 29 Kg and more preferably about 11 Kg of catalyst inventory or less per M 3 of daily plant through-put.
  • catalyst may be added continuously or periodically, such as, for example, to make up for normal losses of catalyst from the system.
  • catalyst addition may be conducted in conjunction with withdrawal of catalyst, such as, for example, to maintain or increase the average activity level of the catalyst in the unit.
  • the rate at which virgin catalyst is added to the unit may be in the range of 0.0019 Kg/M 3 to 0.057 Kg/M 3 , more preferably 0.00289 Kg/M 3 to 0.038 Kg/M 3 , and most preferably to 0.0038 Kg/M 3 to 0.289 Kg/M 3 of feed.
  • equilibrium catalyst from FCC operation is to be utilized, replacement rates as high as about 0.095 Kg/M 3 can be practiced.
  • the process may be practised with catalyst bearing high accumulations of heavy metal(s) in the form of elemental metal(s), oxide(s), sulfide(s) or other compounds.
  • catalyst bearing heavy metals accumulations in the range of about 3000 or more ppm Nickel Equivalents, on the average, is contemplated.
  • the concentration of Nickel Equivalents of metals on catalyst can range up to 50,000 ppm or higher. More specifically, the accumulation may be in the range of 3000 to 30,000 ppm, preferably in the range of 3000 to 20,000 ppm, and more particularly 3000 to 12,000 ppm.
  • the equilibrium concentration of heavy metals in the circulating inventory of catalyst can be controlled (including maintained or varied as desired or needed) by manipulation of the rate of catalyst addition discussed above.
  • addition of catalyst may be maintained at a rate which will control the heavy metals accumulation on the catalyst in one of the ranges set forth above.
  • a catalyst having a relatively high level of cracking activity providing high levels of conversion and productivity at low residence times.
  • the conversion capabilities of the catalyst may be expressed in terms of the conversion produced during actual operation of the process and/or in terms of conversion produced in standard catalyst activity tests.
  • conversion is expressed in liquid volume percent, based on fresh feed.
  • the preferred catalyst may be defined as one which in its virgin or equilibrium state, exhibits a specified activity expressed as a percentage in terms of MAT (micro-activity test) conversion.
  • the foregoing percentage is the volume percentage of standard feedstock which a catalyst under evaluation will convert to 221°C end point gasoline, lighter products and coke at 482°C, 16 WHSV (weight hourly space velocity, calculated on a moisture free basis, using clean catalyst which has been dried at 593°C, weighed and then conditioned, for a period of at least 8 hours at about 25°C and 50% relative humidity, until about one hour or less prior to contacting the feed) and 3C/0 (catalyst to oil weight ratio) by ASTM D-32 MAT test D-3907-80, using an appropriate standard feedstock, e.g. sweet light primary gas oil, such as that used by Davison, Division of W. R. Grace, having the following analysis and properties:
  • the gasoline end point and boiling temperature-volume percent relationships of the produce produced in the MAT conversion test may for example be determined by simulated distillation techniques, for example modifications of gas chromate graphic "Sim-D", ASTM D-2887-73. The results of such simulations are in reasonable agreement with the results obtained by subjecting larger samples of material to standard laboratory distillation techniques. Conversion is calculated by subtracting from 100 the volume percent (based on fresh feed) of those products heavier than gasoline which remain in the recovered product.
  • relative activity is a ratio obtained by dividing the weight of a standard or reference catalyst which is or would be required to produce a given level of conversion, as compared to the weight of an operating catalyst (whether proposed or actually used) which is or would be required to produce the same level of conversion in the same or equivalent feedstock under the same or equivalent conditions.
  • Said ratio of catalyst weights may be expressed as a numerical ratio, but preferably is converted to a percentage basis.
  • the standard catalyst is preferably chosen from among catalysts useful for conducting the present invention, such as for example zeolite fluid cracking catalysts, and is chosen for its ability to produce a predetermined level of conversion in a standard feed under the conditions of. temperature, WHSV, catalysts to oil ratio and other conditions set forth in the preceding description of the MAT conversion test and in ASTM D-32 MAT test D-3907-80. Conversion is the volume percentage of feedstock that is converted to 221°C end point gasoline, lighter products and coke. For standard feed, one may employ the above-mentioned light primary gas oil, or equivalent.
  • a "standard catalyst curve" a chart or graph of conversion (as above defined) vs. reciprocal WHSV for the standard catalyst and feedstock.
  • a sufficient number of runs is made under ASTM D-3907-80 conditions (as modified above) using standard feedstock at varying levels of WHSV to prepare an accurate "curve" of conversion vs. WHSV for the standard feedstock.
  • This curve should traverse all or substantially all of the various levels of conversion including the range of conversion within which it is expected that the operating catalyst will be tested. From this curve, one may establish a standard WHSV for test comparisons and a standard value of reciprocal WHSV corresponding to that level of conversion which has been chosen to represent 100% relative activity in the standard catalyst.
  • the aforementioned reciprocal WHSV and level of conversion are, respectively, 0.0625 and 75%.
  • the relative activity may then be calculated by dividing the hypothetical reiprocal WHSV by the reciprocal standard WHSV, which is 1/16, or .0625.
  • the result is relative activity expressed in terms of a decimal fraction, which may then be multiplied by 100 to convert to percent relative activity.
  • a relative activity of 0.5, or 50% means that it would take twice the amount of the operating catalyst to give the same conversion as the standard catalyst, i.e., the production catalyst is 50% as active as the reference catalyst.
  • the catalyst may be introduced into the process in its virgin form or, as previously indicated, in other than virgin form; e.g. one may use equilibrium catalyst withdrawn from another.unit, such as catalyst that has been employed in the cracking of a different feed.
  • equilibrium catalyst withdrawn from another.unit such as catalyst that has been employed in the cracking of a different feed.
  • the preferred catalysts may be described on the basis of their activity "as introduced” into the process of the present invention, or on the basis of their "as withdrawn” or equilibrium activity in the process of the present invention, or on both of these bases.
  • a preferred activity level of virgin and non-virgin catalyst "as introduced” into the process of the present invention is at least about 60% by MAT conversion, and preferably at least about 20%, more preferably at least about 40% and still more preferably at least about 60% in terms of relative activity.
  • An acceptable "as withdrawn” or equilibrium activity level of catalyst which has been used in the process of the present invention is at least about 20% or more, but about 40% or more and preferably about 60% or more are preferred values on a relative activity basis, and an activity level of 60% or more on a MAT conversion basis is also contemplated. More preferably, it is desired to employ a catalyst which will, under the conditions of use in the unit, establish an equilibrium activity at or above the indicated level. The catalyst activities are determined with catalyst having less than 0.01 coke, e.g. regenerated catalyst.
  • a particularly preferred class of catalysts includes those which have pore structures into which molecules of feed material may enter for adsorption and/or for contact with active catalytic sites within or adjacent the pores.
  • Various types of catalysts are available within this classification, including for example the layered silicates, e.g. smectites. Although the most widely available catalysts within this classification are the well-known zeolite-containing catalysts, non-zeolite catalysts are also contemplated.
  • the preferred zeolite-containing catalysts may include any zeolite, whether natural, semisynthetic or synthetic, alone or in admixture with other materials which do not significantly impair the suitability of the catalyst, provided the resultant catalyst has the activity and pore structure referred to above.
  • the virgin catalyst may include the zeolite component associated with or dispersed in a porous refractory inorganic oxide carrier, in such case the catalyst may for example contain 1% to 60%, more preferably 15 to 50%, and most typically 20 to 45% by weight, based on the total weight of catalyst (water free basis) of the zeolite, the balance of the catalyst being the porous refractory inorganic oxide alone or in combination with any of the known adjuvants for promoting or suppressing various desired and undesired reactions.
  • the zeolite components of the zeolite-containing catalysts will be those which are known to be useful in FCC cracking processes.
  • these are crystalline aluminosilicates, typically made up of tetra coordinated aluminum atoms associated through oxygen atoms with adjacent silicon atoms in the crystal structure.
  • zeolite as used in this disclosure contemplates not only aluminosilicates, but also substances in which the aluminum has been partly or wholly replaced, such as for instance by gallium and/or other metal atoms, and further includes substances in which all or part of the silicon has been replaced, such as for instance by germanium. Titanium and zirconium substitution may also be practiced.
  • the zeolite may be ion exchanged, and where the zeolite is a component of a catalyst composition, such ion exchanging may occur before or after incorporation of the zeolite as a component of the composition.
  • Suitable cations for replacement of sodium in the zeolite crystal structure include ammonium (decomposable to hydrogen), hydrogen, rare earth metals, alkaline earth metals, etc.
  • suitable ion exchange pro-* cedures and cations which may be exchanged into the zeolite crystal structure are well known to those skilled in the art.
  • Examples of the naturally occurring crystalline aluminosilicate zeolites which may be used as or included in the catalyst for the present invention are faujasite, mordenite, clinoptilote, chabazite, analcite, crionite, as well as levynite, dachiardite, paulingite, noselite, ferriorite, heulandite, scolccite, stibite, harmotome, phillipsite, brew- sterite, flarite, datolite, gmelinite, caumnite, leucite, lazurite, scaplite, mesolite, ptolite, nephline, matrolite, offretite and sodalite.
  • Examples of the synthetic crystalline aluminosilicate zeolites which are useful as or in the catalyst for carrying out the present invention are Zeolite X, U.S. Patent No. 2,882,244, Zeolite Y, U.S. Patent No. 3,130,007; and Zeolite A, U.S. Patent No. 2,882,243; as well as Zeolite B, U.S. Patent No. 3,008,803; Zeolite D, Canada Patent No. 661,981; Zeolite E, Canada Patent No. 614,495; Zeolite F, U.S. Patent No. 2,996,358; Zeolite H, U.S. Patent No. 3,010,789; Zeolite J, U.S. Patent No.
  • the crystalline aluminosilicate zeolites having a faujasite-type crystal structure are particularly preferred for use in the present invention. This includes particularly natural faujasite and Zeolite X and Zeolite Y.
  • the crystalline aluminosilicate zeolites such as synthetic faujasite, will under normal conditions crystallize as regularly shaped, discrete particles of about one to about ten microns in size, and, accordingly, this is the size range frequently found in commercial catalysts which can be used in the invention.
  • the particle size of the zeolites is from 0.1 to 10 microns and more preferably is from 0.1 to 2 microns or less.
  • zeolites prepared in situ from calcined kaolin may be characterized by even smaller crystallites. Crystalline zeolites exhibit both an interior and an exterior surface area, which we have defined as "portal" surface area, with the largest portion of the total surface area being internal.
  • portal surface area we refer to the outer surface of the zeolite crystal through which reactants are considered to pass in order to convert to lower boiling products.
  • Blockages of the internal channels by, for example, coke formation, blockages of entrance to the internal channels by deposition of coke in the portal surface area, and contamination by metals poisoning, will greatly reduce the total zeolite surface area. Therefore, to minimize the effect of contamination and pore blockage, crystals larger than the normal size cited above are preferably not used in the catalysts of this invention.
  • zeolite-containing catalysts are available with carriers containing a variety of metal oxides and combination thereof, including for example silica, alumina, magnesia, and mixtures thereof and mixtures of such oxides with clays as e.g. described in U.S. Patent No. 3,034,948.
  • metal oxides and combination thereof including for example silica, alumina, magnesia, and mixtures thereof and mixtures of such oxides with clays as e.g. described in U.S. Patent No. 3,034,948.
  • One may for example select any of the zeolite-containing molecular sieve fluid cracking catalysts which are suitable for production of gasoline from vacuum gas oils.
  • certain advantages may be attained by judicious selection of catalysts having marked resistance to metals.
  • a metal resistant zeolite catalyst is, for instance, described in U.S. Patent No.
  • the catalyst contains 1-40 weight percent of a rare earth-exchanged zeolite, the balance being a refractory metal oxide having specified pore volume and size distribution.
  • Other catalysts described as "metals-tolerant" are described in the above mentioned Cimbalo et al article.
  • a useful catalyst may have a skeletal density of about 2400 Kg/M 3 and an average particle size of about 60-70 microns, with less than 10% of the particles having a size less than about 40 microns and less than 80% having a size less than about 50-60 microns.
  • the AGZ-290, GRZ-1, CCZ-220 and Super DX catalysts referred to above are products of W. R. Grace and Co.
  • F-87 and FOC-90 are product of Filtrol, while HFZ-20 and HEZ-55 are products of Engelhard/Houdry.
  • the above are properties of virgin catalyst and, except in the case of zeolite content, are adjusted to a water free basis, i.e. based on material ignited at 954°C.
  • the zeolite content is derived by comparison of the x-ray intensities of a catalyst sample and of a standard material composed of high purity sodium Y zeolite in accordance with draft #6, dated January 9, 1978, of proposed ASTM Standard Method entitled "determination of the Faujasite Content of a Catalyst".
  • the Super D family and especially a catalyst designated GRZ-1 are particularly preferred.
  • Super DX has given particularly good results with Arabian light crude.
  • the GRZ-1 although substantially more expensive than the Super DX at present, appears somewhat more metals-tolerant.
  • catalysts for carrying out the present invention will be those which, according to proposals advanced by Dr. William P. Hettinger, Jr. and Dr. James E. Lewis, are characterized by matrices with feeder pores having large minimum diameters and large mouths to facilitate diffusion of high molecular weight molecules through the matrix to the portal surface area of molecular sieve particles within the matrix.
  • Such matrices preferably also have a relatively large pore volume in order to soak up unvaporized portions of the carbo-metallic oil feed.
  • significant numbers of liquid hydrocarbon molecules can diffuse to active catalytic sites both in the matrix and in sieve particles on the surface of the matrix.
  • the process of the present invention can be conducted in the substantial absence of tin and/or antimony or at least in the presence of a catalyst which is substantially free of either or both of these metals.
  • the process of the present invention may be operated with the above described carbo-metallic oil and catalyst as substantially the sole materials charged to the reaction zone. But the charging of additional materials is not excluded. The charging of recycled oil to the reaction zone has already been mentioned. As described in greater detail below, still other materials fulfilling a variety of functions may also be charged. In such case, the carbo-metallic oil and catalyst usually represent the major proportion by weight of the total of all materials charged to the reaction zone.
  • Certain of the additional materials which may be used perform functions which offer significant advantages over the process as performed with only the carbo-metallic oil and catalyst. Among these functions are: controlling the effects of heavy metals and other catalyst contaminants: enhancing catalyst activity; absorbing excess heat in the catalyst as received from the regenerator; disposal of pollutants or conversion thereof to a form or forms in which they may be more readily separated from products and/or disposed of; controlling catalyst temperature; diluting the carbo-metallic oil vapors to reduce their partial pressure and increase the yield of desired products; adjusting feed/catalyst contact time; donation of hydrogen to a hydrogen deficient carbo-metallic oil feedstock, for example as disclosed in U.S.
  • Certain of the metals in the heavy metals accumulation on the catalyst are more active in promoting undesired reactions when they are in the form of elemental metal, than they are when in the oxidized form produced by contact with oxygen in the catalyst regenerator.
  • the time of contact between catalyst and vapors of feed and product in past conventional catalytic cracking was sufficient so that hydrogen released in the cracking reaction was able to reconvert a significant portion of the less harmful oxides back to the more harmful elemental heavy metals.
  • additional materials which are in gaseous (including vaporous) form in the reaction zone in admixture with the catalyst and vapors of feed and products.
  • the increased volume of material in the reaction zone resulting from the presence of such additional materials tends to increase the velocity of flow through the reaction zone with a corresponding decrease in the residence time of the catalyst and oxidized heavy metals borne thereby. Because of this reduced residence time, there is less opportunity for reduction of the oxidized heavy metals to elemental form and therefore less of the harmful elemental metals are available for contacting the feed and products.
  • Added materials may be introduced into the process in any suitable fashion, some examples of which follow. For instance, they may be admixed with the carbo-metallic oil feedstock prior to contact of the latter with the catalyst. Alternatively, the added materials may, if desired, be admixed with the catalyst prior to contact of the latter with the feedstock. Separate portions of the added materials may be separately admixed with both catalyst and carbo-metallic oil. Moreover, the feedstock, catalyst and additional materials may, if desired, be brought together substantially simultaneously. A portion of the added materials may be mixed with catalyst and/or carbo-metallic oil in any of the above described ways, while additional portions are subsequently brought into admixture.
  • a portion of the added materials may be added to the carbo-metallic oil and/or to the catalyst before they reach the reaction zone, while another portion of the added materials is introduced directly into the reaction zone.
  • the added materials may be introduced at a plurality of spaced locations in the reaction zone or along the length thereof, if elongated.
  • the amount of additional materials which may be present in the feed, catalyst or reaction zone for carrying out the above functions, and others, may be varied as desired; but said amount will preferably be sufficient to substantially heat balance the process.
  • These materials may for example be introduced into the reaction zone in a weight ratio relative to feed of up to 0.4, preferably in the range of 0.02 to 0.4, more preferably 0.03 to 0.3 and most preferably 0.05 to 0.25.
  • H 2 0 may or all of the above desirable functions may be attained by introducing H 2 0 to the reaction zone in the form of steam or of liquid water or a combination thereof in a weight ratio relative to feed in the range of 0.04 or more, or more preferably 0.05 to 0.1 or more.
  • H 2 0 tends to inhibit reduction of catalyst-borne oxides, sulfites and sulfides to the free metallic form which is believed to promote condensation-dehydrogenation with consequent promotion of coke and hydrogen yield and accompanying loss of product.
  • H 2 0 may also, to some extent, reduce deposition of metals onto the catalyst surface.
  • H 2 0 tends to increase the acidity of the catalyst by Bronsted acid formation which in turn enhances the activity of the catalyst. Assuming the H 2 0 as supplied is cooler than the regenerated catalyst and/or the temperature of the reaction zone, the sensible heat involved in raising the temperature of the H 2 0 upon contacting the catalyst in the reaction zone or elsewhere can absorb excess heat from the catalyst. Where the H 2 0 is or includes recycled water that contains for example 500 to 5000 ppm of H 2 S dissolved therein, a number of additional advantages may accrue.
  • H 2 S need not be vented to the atmosphere, the recycled water does not require further treatment to remove H 2 S and the H 2 S may be of assistance in reducing coking of the catalyst by passivation of the heavy metals, i.e. by conversion thereof to the sulfide form which has a lesser tendency than the free metals to enhance coke and hydrogen production.
  • the presence of H 2 0 can dilute the carbo-metallic oil vapors, thus reducing their partial pressure and tending to increase the yield of the desired products. It has been reported that H 2 0 is useful in combination with other materials in generating hydrogen during cracking; thus it may be able to act as a hydrogen donor for hydrogen deficient carbo-metallic oil feedstocks.
  • the H 2 0 may also serve certain purely mechanical functions such as: assisting in the atomizing or dispersion of the feed; competing with high molecular weight molecules for adsorption on the surface of the catalyst, thus interrupting coke formation; steam distillation of vaporizable product from unvaporized feed material; and disengagement of product from catalyst upon conclusion of the cracking reaction. It is particularly preferred to bring together H 2 0, catalyst and carbo-metallic oil substantially simultaneously. For example, one may admix H 2 0 and feedstock in an atomizing nozzle and immediately direct the resultant spray into contact with the catalyst at the downstream and of the reaction zone.
  • liquid water be brought into intimate admixture with the carbo-metallic oil in a weight ratio of 0.04 to 0.25 at or prior to the time of introduction of the oil into the reaction zone, whereby the water (e.g., in the form of liquid water or in the form of steam produced by vaporization of liquid water in contact with the oil) enters the reaction zone as part of the flow of feedstock which enters such zone.
  • the foregoing is advantageous in promoting dispersion of the feedstock.
  • the heat of vaporization of the water which heat is absorbed from the catalyst, from the feedstock, or from both, causes the water to be a more efficient heat sink than steam alone.
  • the weight ratio of liquid water to feed is 0.04 to 0.2 more preferably 0.05 to 0.15.
  • the liquid water may be introduced into the process in the above described manner or in other ways, and in either event the introduction of liquid water may be accompanied by the introduction of additional amounts of water as steam into the same or different portions of the reaction zone or into the catalyst and/or feedstock.
  • the amount of additional steam' may be in a weight ratio relative to feed in the range of 0.01 to 0.25, with the weight ratio of total H 2 0 (as steam and liquid water) to feedstock being 0.3 or less.
  • the charging weight ratio of liquid water relative to steam in such combined use of liquid water and steam may for example range from 15 which is presently preferred, to 0.2. Such ratio may be maintained at a predetermined level within such range or varied as necessary or desired to adjust or maintain heat balance.
  • the hydrogenation-condensation activity of heavy metals may be inhibited by introducing hydrogen sulfide gas into the reaction zone.
  • Hydrogen may be made available for hydrogen deficient carbo-metallic oil feedstocks by introducing into the reaction zone either a conventional hydrogen donor diluent such as a heavy naphtha or relatively low molecular weight carbon-hydrogen fragment contributors, including for example: light paraffins; low molecular weight alcohols and other compounds which permit or favor intermolecular hydrogen transfer; and compounds that chemically combine to generate hydrogen in the reaction zone such as by reaction of carbon monoxide with water, or with alcohols, or with olefins, or with other materials or mixtures of the foregoing.
  • a conventional hydrogen donor diluent such as a heavy naphtha or relatively low molecular weight carbon-hydrogen fragment contributors, including for example: light paraffins; low molecular weight alcohols and other compounds which permit or favor intermolecular hydrogen transfer; and compounds that chemically combine to generate hydrogen in
  • the invention may be practiced in a wide variety of apparatus.
  • the preferred apparatus includes means for rapidly vaporizing as much feed as possible and efficiently admixing feed and catalyst (although not necessarily in that order), for causing the resultant mixture to flow as a dilute suspension in a progressive flow mode, and for separating the catalyst from cracked products and any uncracked or only partially cracked feed at the end -of a predetermined residence time or times, it being preferred that all or at least a substantial portion of the products should be abruptly separated from at least a portion of the catalyst.
  • the apparatus may include, along its elongated reaction chamber, one or more points for introduction of carbo-metallic feed, one or more points for introduction of catalyst, one or more points for introduction of additional material, one or more points for withdrawal of products and one or more points for withdrawal of catalyst.
  • the means for introducing feed, catalyst and other material may range from open pipes to sophisticated jets or spray nozzles, it being preferred to use means capable of breaking up the liquid feed into fine droplets.
  • the catalyst, liquid water (when used) and fresh feed are brought together in an apparatus similar to that disclosed in U.S. 4,432,864.
  • the liquid water and carbo-metallic oil, prior to their introduction into the riser are caused to pass through a propeller, apertured disc, or any appropriate high shear agitating means for forming a "homogenized mixture" containing finely divided droplets of oil and/or water with oil and/or water present as a continuous phase.
  • the reaction chamber or at least the major portion thereof, be more nearly vertical than horizontal and have a length to diameter ratio of at least about 10, more preferably about 20 or 25 or more.
  • a vertical riser type reactor is preferred. If tubular, the reactor can be of uniform diameter throughout or may be provided with a continuous or step-wise increase in diameter along the reaction path to maintain or vary the velocity along the flow path.
  • the charging means (for catalyst and feed) and the reactor configuration are such as to provide a relatively high velocity of flow and dilute suspension of catalyst.
  • the vapor or catalyst velocity in the riser will be usually at least 7.62 and more typically at least 10.7 meters per second. This velocity may range up to about 16.8 meters or 22.9 meters or 30.5 meters per second or higher.
  • the vapor velocity at the top of the reactor may be higher than that at the bottom and may for example be 24.4 meters per second at the top and 12.2 meters per second at the bottom.
  • the velocity capabilities of the reactor will in general be sufficient to prevent substantial build-up of catalyst bed in the bottom or other portions of the riser, whereby the catalyst loading in the riser can be maintained below 1.8 to 2.3 Kg, as for example about 0.23 Kg, and below about 0.9 Kg, as for example 0.36 Kg, per cubic foot, respectively, at the upstream (e.g. bottom) and downstream (e.g. top) ends of the riser.
  • the progressive flow mode involves, for example, flowing of catalyst, feed and products as a stream in a positively controlled and maintained direction established by the elongated nature of the reaction zone. This is not to suggest however that there must be strictly linear flow. As is well known, turbulent flow and "slippage" of catalyst may occur to some extent especially in certain ranges of vapor velocity and some catalyst loadings, although it has been reported advisable to employ sufficiently low catalyst loadings to restrict slippage and back-mixing.
  • the reactor is one which abruptly separates a substantial portion or all of the vaporized cracked products from the catalyst at one or more points along the riser, and preferably separates substantially all of the vaporized cracked products from the catalyst at the downstream end of the riser.
  • a preferred type of reactor embodies ballistic separation of catalyst and products; that is, catalyst is projected in a direction established by the riser tube, and is caused to continue its motion in the general direction so established, while the products, having lesser momentum, are caused to make an abrupt change of direction, resulting in an abrupt, substantially instantaneous separation of product from catalyst.
  • the riser tube is provided with a substantially unobstructed discharge opening at its downstream end for discharge of catalyst.
  • An exit port in the side of the tube adjacent the downstream end receives the products.
  • the discharge opening communicates with a catalyst flow path which extends to the usual stripper and regenerator, while the exit port communicates with a product flow path which is substantially or entirely separated from the catalyst flow path and leads to separation means for separating the products from the relatively small portion of catalyst, if any, which manages to gain entry to the product exit port.
  • a ballistic separation apparatus and technique as above described are found in U.S. Patents 4,066,533 and 4,070,159 to Myers et al, the disclosures of which patents are hereby incorporated herein by reference in their entireties. According to a particularly preferred embodiment, based on a suggestion understood to have emanated from Paul W. Walters, Roger W.
  • the ballistic separation step includes at least a partial reversal of direction by the product vapors upon discharge from the risertube; that is, the product vapors make a turn or change of direction which exceeds 32,2°C at the riser tube outlet.
  • This may be accomplished for example by providing a cup-like member surrounding the riser tube at its upper end, the ratio of cross-sectional area of the cup-like member relative to the cross-sectional area of the riser tube outlet being low i.e. less than 1 and preferably less than about 0.6.
  • the lip of the cup is slightly downstream of, or above the downstream end or top of the riser tube, and the cup is preferably concentric with the riser tube.
  • the feedstock is customarily preheated, often to temperatures significantly higher than are required to make the feed sufficiently fluid for pumping and for introduction into the reactor.
  • preheat temperatures as high as about 371°C or 427°C have been reported.
  • it may be fed at ambient temperature. Heavier stocks may be fed at preheat temperatures of up to about 316°C, typically 93°C to 360°C, but higher preheat temperatures are not necessarily excluded.
  • the catalyst fed to the reactor may vary widely in temperature, for example from 593°C to 871°C, more preferably 593°C to 816°C and most preferably 704°C to 760°C, with 718°C to 746°C being considered optimum at present.
  • the conversion of the carbo-metallic oil to lower molecular weight products may be conducted at a temperature of 482°C to 760°C measured at the reaction chamber outlet.
  • the reaction temperature as measured at said outlet is more preferably maintained in the range of 518°C to 704°C, still more preferably about 526°C to about 621°C, and most preferably 527°C to 621°C.
  • all of the feed may or may not vaporize in the riser.
  • the pressure in the reactor may, as indicated above, range from 0.68 to 3.4 atm, preferred and more preferred pressure ranges are 1.0 to 2.4 and 1.4 to 2.4.
  • the partial (or total) pressure of the feed may be in the range of 0.2 to 0.68 to 1.2 atm.
  • the feed partial pressure may be controlled or suppressed by the introduction of gaseous (including vaporous) materials into the reactor, such as for instance the steam, water and other additional materials described above.
  • the process has for example been operated with the ratio of feed partial pressure relative to total pressure in the riser in the range of 0.2 to 0.8, more typically 0.3 to 0.7 and still more typically 0.4 to 0.6, with the ratio of added gaseous material (which may include recycled gases and/or steam resulting from introduction of H20 to the riser in the form of steam and/or liquid water) relative to total pressure in the riser correspondingly ranging from 0.8 to 0.2, more typically 0.7 to 0.3 and still more typically 0.6 fo 0.4.
  • the ratio of the partial pressure of the added gaseous material relative to the partial pressure of the feed has been in the range of 0.25 to 4.0, more typically 0.4 to 2.3 and still more typically 0.7 to 1.7.
  • the residence time of feed and product vapors in the riser may be in the range of 0.5 to 10 seconds, as described above, preferred and more preferred values are 0.5 to 6 and 1 to 4 seconds, with 1.5 to 3.0 seconds currently being considered about optimum.
  • the process has been operated with a riser vapor residence time of 2.5 seconds or less by introduction of copious amounts of gaseous materials into the riser, such amounts being sufficient to provide for example a partial pressure ratio of added gaseous materials relative to hydrocarbon feed of 0.8 or more.
  • the process has been operated with said residence time being about two seconds or less, with the aforesaid ratio being in the range of 1 to 2.
  • the catalyst riser residence time may or may not be the same as that of the vapors.
  • the ratio of average catalyst reactor residence time versus vapor reactor residence time i.e. slippage, may be in the range of 1 to 5, more preferably 1 to 4 and most preferably 1 to 3, with 1 to 2 currently being considered optimum.
  • slippage there will usually be a small amount of slippage, e.g., at least 1.05 or 1.2. In an operating unit there may for example be a slippage of 1.1 at the bottom of the riser and 1.05 at the top.
  • the vapor riser residence time and vapor catalyst contact time are substantially the same for at least about 80%, more preferably at least about 90% and most preferably at least about 95% by volume of the total feed and product vapors passing through the riser.
  • the combination of catalyst to oil ratio, temperatures, pressures and residence times should be such as to effect a substantial conversion of the carbo-metallic oil feedstock. It is an advantage of the process that very high levels of conversion can be attained in a single pass; for example the conversion may be in excess of 50% and may range to about 90% or higher. Preferably, the aforementioned conditions are maintained at levels sufficient to maintain conversion levels in the range of about 60 to 90% and more preferably 70 to 85%. The foregoing conversion levels are calculated by subtracting from 100% the percentage obtained by dividing the liquid volume of fresh feed into 100 times the volume of liquid product boiling at and above 221°C (tbp, standard atmospheric pressure).
  • coke laydown may be in excess of about 0.3, more commonly in excess of about 0.5 and very frequently in excess of about 1% of coke by weight, based on the weight of moisture free regenerated catalyst. Such coke laydown may range as high as about 2%, or about 3%, or even higher.
  • the present process includes stripping of spent catalyst after disengagement of the catalyst from product vapors.
  • Persons skilled in the art are acquainted with appropriate stripping agents and conditions for stripping spent catalyst, but in some cases the present process may require somewhat more severe conditions than are commonly employed. This may. result, for example, from the use of a carbo-metallic oil having constituents which do not volatilize under the conditions prevailing in the reactor, which constituents deposit themselves at least in part on the catalyst. Such adsorbed, unvaporized material can be troublesome from at least two standpoints.
  • the gases (including vapors) used to strip the catalyst can gain admission to a catalyst disengagement or collection chamber connected to the downstream end of the riser, and if there is an accumulation of catalyst in such chamber, vaporization of these unvaporized hydrocarbons in the stripper can be followed by adsorption on the bed of catalyst in the chamber. More particularly, as the catalyst in the stripper is stripped of adsorbed feed material, the resultant feed material vapors pass through the bed of catalyst accumulated in the catalyst collection and/or disengagement chamber and may deposit coke and/or condensed material on the catalyst in said bed.
  • the condensed products can create a demand for more stripping capacity, while the coke can tend to increase regeneration temperatures and/or demand greater regeneration capacity.
  • stripping may for example include reheating of the catalyst, extensive stripping with steam, the use of gases having a temperature considered higher than normal for FCCNGO operations, such as for instance flue gas from the regenerator, as well as other refiner stream gases such as hydro- treater off-gas (H2S containing), hydrogen and others.
  • the stripper may be operated at a temperature of about 177°C using steam at a pressure of about 11.2 atm and a weight ratio of steam to catalyst of 0.002 to 0.003.
  • the stripper may be operated at a temperature of about 552°C or higher.
  • Regeneration of catalyst burning away of coke deposited on the catalyst during the conversion of the feed, may be performed at any suitable temperature in the range of 593°C to 871°C, measured at the regenerator catalyst outlet. This temperature is preferably in the range of 749°C to 816°C, more preferably 691°C to 724°C and optimally 718°C to 746°C.
  • the process has been operated, for example, with a fluidized regenerator with the temperature of the catalyst dense phase in the range of 704°C to 760°C.
  • regeneration is conducted while maintaining the catalyst in one or more fluidized beds in one or more fluidization chambers.
  • Such fluidized bed operations are characterized, for instance, by one or more fluidized dense beds of ebulliating particles having a bed density of, for example, 400 to 800 Kg/m 3 per cubic foot.
  • Fluidization is maintained by passing gases, including combustion supporting gases, through the bed at a sufficient velocity to maintain the particles in a fluidized state but at a velocity which is sufficiently small to prevent substantial entrainment of particles in the gases.
  • the lineal velocity of the fluidizing gases may be in the range of 0.06 to 1.2 meters per second and preferably 0.06 to 0.9 meters per second.
  • the average total residence time of the particles in the one or more beds is substantial, ranging for example from 5 to 30, more preferably 5 to 20 and still more preferably 5 to 10 minutes.
  • the amount of regenerator heat which is transmitted to fresh feed by way of recycling regenerated catalyst can substantially exceed the level of heat input which is appropriate in the riser for heating and vaporizing the feed and other materials, for supplying the endothermic heat of reaction for cracking, for making up the heat losses of the unit and so forth.
  • the amount of regenerator heat transmitted to fresh feed may be controlled, or restricted where necessary, within certain approximate ranges.
  • the amount of heat so transmitted may for example be in the range of 277.8 to 666.7, more particularly 333.3 to 499.9 and more particularly 305.5 to 472.2 cal/g of fresh feed.
  • the aforesaid ranges refer to the combined heat, in cal/grams of fresh feed, which is transmitted by the catalyst to the feed and reaction products (between the contacting of feed with catalyst and the separation of product from catalyst) for supplying the heat of reaction (e.g. for cracking) and the difference in enthalpy between the products and the fresh feed.
  • One or a combination of techniques may be utilized for controlling or restricting the amount of regeneration heat transmitted via catalyst to fresh feed. For example, one may add a combustion modifier to the cracking catalyst in order to reduce the temperature of combustion of coke to carbon dioxide and/or carbon monoxide in the regenerator. Moreover, one may remove heat from the catalyst through heat exchange means, including for example heat exchangers (e.g. steam coils) built into the regenerator itself, whereby one may extract heat from the catalyst during regeneration. Heat exchangers can be built into catalyst transfer lines, such as for instance the catalyst return line from the regenerator to the reactor, whereby heat may be removed from the catalyst after it is regenerated.
  • heat exchangers e.g. steam coils
  • the amount of heat imparted to the catalyst in the regenerator may be restricted by reducing the amount of insulation on the regenerator to permit some heat loss to the surrounding atmosphere, especially if feeds of exceedingly high coking potential are planned for processing; in general, such loss of heat to the atmosphere is considered economically less desirable than certain of the other alternatives set forth herein.
  • One may also inject cooling fluids into portions of the regenerator other than those occupied by the dense bed, for example water and/or steam, whereby the amount of inert gas available in the regenerator for heat absorption and removal is increased.
  • Another suitable and preferred technique for controlling or restricting the heat transmitted to fresh feed via recycled regenerated catalyst involves maintaining a specified ratio between the carbon dioxide and carbon monoxide formed in the regenerator while such gases are in heat exchange contact or relationship with catalyst undergoing regeneration.
  • Still another particularly preferred technique for controlling or restricting the regeneration heat imparted to fresh feed via recycled catalyst involves the diversion of a portion of the heat borne by recycled catalyst to added materials introduced into the reactor, such as the water, steam naphtha, other hydrogen donors, flue gases, inert gases and other gaseous or vaporizable materials which may be introduced into the reactor.
  • the catalyst discharged from the regenerator may be stripped with appropriate stripping gases to remove oxygen containing gases.
  • stripping may for instance be conducted at relatively high temperatures, for example 732°C to 743°C, using steam nitrogen or other inert gas as the stripping gas(es).
  • nitrogen and other inert gases is beneficial from the standpoint of avoiding a tendency toward hydro-thermal catalyst deactivation which may result from the use of steam.
  • a unit design in which system components and residence times are selected to reduce the ratio of catalyst reactor residence time relative to catalyst regenerator residence time will tend to reduce the ratio of the times during which the catalyst is respectively under reduction conditions and oxidation conditions. This too can assist in maintaining desired levels of selectivity.
  • the metals content of the catalyst is being managed successfully may be observed by monitoring the total hydrogen plus methane produced in the reactor and/or the ratio of hydrogen to methane thus produced.
  • the hydrogen to methane mole ratio should be less than about 1 and preferably about 0.6 or less, with about 0.4 or less being considered about optimum.
  • the hydrogen to methane ratio may range from 0.5 to 1.5 and average 0.8 to 1.
  • Careful carbon management can improve both selectivity (the ability to maximize production of valuable products), and heat productivity.
  • the techniques of metals control described above are also of assistance in carbon management.
  • the usefulness of water addition in respect to carbon management has already been spelled out in considerable detail in that part of the specification which relates to added materials for introduction into the reaction zone.
  • those techniques which improve dispersion of the feed in the reaction zone should also prove helpful, these include for instance the use of fogging or misting devices to assist in dispersing the feed.
  • Catalyst to oil ratio is also a factor in heat management.
  • the reactor temperature may be controlled in the practice of the present invention by respectively increasing or decreasing the flow of hot regenerated catalyst to the reactor in response to decreases and increases in reactor temperature, typically the outlet temperature in . the case of a riser type reactor.
  • the automatic controller for catalyst introduction is set to maintain an excessive catalyst to oil ratio, one can expect unnecessarily large rates of carbon production and heat release, relative to the weight of fresh feed charged to the reaction zone.
  • Relatively high reactor temperatures are also beneficial from the standpoint of carbon management. Such higher temperatures foster more complete vaporization of feed and disengagement of product from catalyst.
  • Carbon management can also be facilitated by suitable restriction of the total pressure in the reactor and the partial pressure of the feed.
  • relatively small decreases in the aforementioned pressures can substantially reduce coke production. This may be due to the fact that restricting total pressure tends to enhance vaporization of high boiling components of the feed, encourage cracking and facilitate disengagement of both unconverted feed and higher boiling cracked products from the catalyst. It may be of assistance in this regard to restrict the pressure drop of equipment downstream of and in communication with the reactor. But if it is desired or necessary to operate the system at higher total pressure, such as for instance because of operating limitations (e.g. pressure drop in downstream equipment) the above described benefits may be obtained by restricting the feed partial pressure. Suitable ranges for total reactor pressure and feed partial pressure have been set forth above, and in general it is desirable to attempt to minimize the pressures within these ranges.
  • a particularly desirable mode of operation from the standpoint of carbon management is to operate the process in the vented riser using a hydrogen donor if necessary, while maintaining the feed partial pressure and total reactor pressure as low as possible, and incorporating relatively large amounts of water, steam and if desired, other diluents, which provide the numerous benefits discussed in greater detail above.
  • a hydrogen donor if necessary, while maintaining the feed partial pressure and total reactor pressure as low as possible, and incorporating relatively large amounts of water, steam and if desired, other diluents, which provide the numerous benefits discussed in greater detail above.
  • liquid water, steam, hydrogen donors, hydrogen and other gaseous or vaporizable materials are fed to the reaction zone, the feeding of these materials provides an opportunity for exercising additional control over catalyst to oil ratio.
  • the practice of increasing or decreasing the catalyst to oil ratio for a given amount of decrease or increase in reactor temperature may be reduced or eliminated by substituting either appropriate reduction or increase in the charging ratios of the water, steam and other gaseous or vaporizable material, or an appropriate reduction or increase in the ratio of water to steam and/or other gaseous materials introduced into the reaction zone.
  • Heat management includes measures taken to control the amount of heat released in various parts of the process and/or for dealing successfully with such heat as may be released. Unlike conventional FCC practice using VGO, wherein it is usually a problem to generate sufficient heat during regeneration to heat balance the reactor, the processing of carbo-metallic oils generally produces so much heat as to require careful management thereof.
  • Heat management can be facilitated by various techniques associated with the materials introduced into the reactor.
  • heat absorption by feed can be maximized by minimum preheating of feed, it being necessary only that the feed temperature be high enough so that it is sufficiently fluid for successful pumping and dispersion in the reactor.
  • the catalyst is maintained in a highly active state with the suppression of coking (metals control), so as to achieve higher conversion, the resultant higher conversion and greater selectivity can increase the heat absorption of the reaction.
  • higher reactor temperatures promote catalyst conversion activity in the face of more refractory and higher boiling constituents with high coking potentials. While the rate of catalyst deactivation may thus be increased, the higher temperature of operation tends to offset this loss in activity.
  • Fig. 1 petroleum feedstock is introduced into the lower end of riser reactor 2 through inlet line 1 at which point it is mixed with hot regenerated catalyst coming from regenerator 9 through line 3.
  • the feedstock is catalytically cracked in passing up riser 2 and the product vapors are separated from spent catalyst in vessel 8.
  • the catalyst particles move upwardly from riser 2 into the space within vessel 8 and fall downwardly into dense bed 16.
  • the cracking products together with some catalyst fines pass through horizontal line 4 into cyclone 5.
  • the gases are separated from the catalyst and pass out through line 6.
  • the catalyst fines drop into bed 16 through dipleg 19.
  • the spent catalyst coated with coke and vanadium in a reduced state, passes through line 7 into upper dense fluidized bed 18 within regenerator 9.
  • the spent catalyst is fluidized with a mixture of air, CO and C0 2 passing through porous plate 21 from lower zone 20.
  • the spent catalyst is partially regenerated in bed 18 and is passed into the lower portion of vented riser 13 through line 11. Air is introduced into riser 13 through line 12 where it is mixed with partially regenerated catalyst.
  • the catalyst is forced rapidly upwards through the riser and it falls into dense settled bed 17.
  • Line 14 provides a source of reducing gas such as CO for bed 17 to keep the regenerated catalyst in a reducing atmosphere and thus keep vanadium present in a reduced oxidation state.
  • Regenerated catalyst is returned to the riser reactor 2 through line 3, which is provided with a source of a reducing gas such as CO through line 22.
  • spent catalyst coated with coke and vanadium in a reduced state flows into dense fluidized bed 32 of regenerator 31 through inlet line 33.
  • Air to combust the coke and fluidize the catalyst is introduced through line 34 into air distributor 35.
  • Coke is burned and passes upwardly into riser regenerator 36.
  • the partially regenerated catalyst which reaches the riser 35 is contacted with air from line 37 which completes the regeneration.
  • the regenerated catalyst passes upwardly from the top of the riser 36 and falls down into dense settled bed 37.
  • Dense bed 37 and the zone above 37 through which the regenerated catalyst falls are supplied with a reducing gas such as CO through lines 40 and 41.
  • the regenerated catalyst is returned to the cracking reactor through line 38.
  • the CO-rich flue gases leave the regenerator through line 39.
  • a carbo-metallic feed at a temperature of about 204°C is fed at a rate of about 907 Kg per hour into the bottom of a vented riser reactor where it is mixed with a zeolite catalyst at a temperature of about 691°C and a catalyst to oil ratio by weight of about 11.
  • the carbo-metallic feed has a heavy metal content of about 5 ppm Nickel Equivalents, including 3 ppm vanadium, and has a Conradson carbon content of about 7 percent. About 85 percent of the feed boils above 343°C and about 20 percent of the feed boils above 552°C.
  • the temperature within the reactor is about 552°C and the pressure is about 1.84 atm.
  • About 75 percent of the feed is converted to fractions boiling at a temperature less than 221°C and about 53 percent of the feed is converted to gasoline. During the conversion, about 11 percent of the feed is converted to coke.
  • the catalyst containing about one percent by weight of coke contains about 20,000 ppm Nickel Equivalents including about 12,000 ppm vanadium.
  • the catalyst is stripped with steam at a temperature of about 552°C to remove volatiles and the stripped catalyst is introduced into the upper zone of the regenerator as shown in Figure 1 at a rate of about 10,433 Kg per hour, and is partially regenerated to a coke concentration of about 0.2 percent by a mixture of air, CO and C0 2 .
  • the CO/CO 2 ratio in the fluidized bed in the upper zone is about 0.3.
  • the partially regenerated catalyst is passed to the bottom of a riser reaction where it is contacted with air in an amount sufficient to force the catalyst up the riser with a residence time of about 1 second.
  • the regenerated catalyst having a coke loading of about 0.05 percent exits from the top of the riser and falls into a dense bed having a reducing atmosphere comprising CO.
  • the regenerated catalyst is recycled to the riser reactor for contact with additional feed.

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  • Chemical & Material Sciences (AREA)
  • Oil, Petroleum & Natural Gas (AREA)
  • Engineering & Computer Science (AREA)
  • Chemical Kinetics & Catalysis (AREA)
  • General Chemical & Material Sciences (AREA)
  • Organic Chemistry (AREA)
  • Production Of Liquid Hydrocarbon Mixture For Refining Petroleum (AREA)
  • Catalysts (AREA)

Claims (10)

1. Un procédé pour convertir une alimentation d'huile d'hydrocarbures contenant du vanadium en produits plus légers, consistant:
à mettre en contact l'alimentation d'huile, dans des conditions de conversion par réduction, avec un catalyseur de craquage pour former des produits plus légers et du coke, de sorte que, par suite des conditions de réduction provoquées dans cette conversion, du vanadium à un état d'oxydation inférieur à +5 et du coke se déposent sur le catalyseur; et
à séparer ces produits plus légers du catalyseur usé portant du vanadium à un état d'oxydation inférieur à +5 et du coke, caractérisé par:
a. le fait que l'on régénère le catalyseur usé en le mettant en contact avec un gaz contenant de l'oxygène, dans des conditions par suite desquelles le coke brûle sur le catalyseur usé, formatn des produits gazeux comprenant CO et CO2, la régénération s'effectuant dans des conditions par suite desquelles le vanadium est maintenu à un état d'oxydation inférieur à +5, et
b. le fait que l'on recycle le catalyseur régénéré vers le régénérateur pour qu'il entre en contact avec de l'alimentation fraîche.
2. Un procédé selon la revendication 1, dans lequel l'alimentation contient de la matière 343°C+, caractérisé par un résidu de carbone à la pyrolyse d'au moins environ 1 et une teneur en Equivalent Nickel de métaux lourds d'au moins environ 4.
3. Un procédé selon la revendication 2, dans lequel la matière 343°C+ représente au moins environ 70% en volume de l'alimentation et comprend au moins environ 10% en volume de matière qui ne bout pas en dessous d'environ 552°C.
4. Le procédé de la revendication 1, dans lequel l'alimentation contient au moins environ 1 ppm de vanadium.
5. Le procédé de la revendication 1, dans lequel le catalyseur de craquage comprend un catalyseur à tamis moléculaire de zéolite contenant de 1 à 60% en poids de tamis.
6. Le procédé de la revendication 1, dans lequel on régénère le catalyseur en au moins deux opérations, dans la première opération desquelles on met en contact le catalyseur usé, en une couche fluidisée dense, avec un gaz contenant moins qu'une quantité stoechiométrique d'oxygène pour convertir l'hydrogène contenu dans le coke en H20 et le carbone contenu dans le coke en C02 et dans l'opération de régénération finale desquelles on met en contact du catalyseur partiellement régénéré avec un excès stoechiométrique d'oxygène pendant un laps de temps inférieur à environ 2 secondes.
7. Le procédé de la revendication 6, dans lequel le catalyseur de l'opération finale comprend une phase dispersée ayant une masse volumique inférieure à 64 kg/m3 (4 livres par pied cubique), le temps de séjour du catalyseur dans la couche fluidisée dense est d'au moins 5 minutes et la couche fluidisée a une masse volumique de 400 à 800 kg/m3 (25 à 50 livres par pied cubique).
8. Le procédé de la revendication 6, dans lequel on met le catalyseur partiellement régénéré en contact avec au moins une quantité stoechiométrique d'oxygène dans un régénérateur à colonne montante, le temps de séjour du catalyseur dans le régénérateur à colonne montante est inférieur à environ 2 secondes et on sépare le catalyseur régénéré des produits gazeux.
9. Le procédé de la revendication 8, dans lequel on met immédiatement en contact le catalyseur, régénéré et séparé, avec un gaz réducteur et on le recueille alors dans une couche dense maintenue sous une atmosphère réductrice.
10. Le procédé de la revendication 8, dans lequel on sépare le catalyseur régénéré des produits gazeux en le projetant dans une direction établie par le régénérateur à colonne montante ou un prolongement de celui-ci, tandis que l'on amène les produits gazeux à effectuer un changement brusque de direction aboutissant à une séparation ballistique brusque pratiquement instantanée des produits gazeux de catalyseur régénéré.
EP82101794A 1981-04-28 1982-03-06 Immobilisation de vanadium déposé sur des matériaux catalytiques pendant la conversion d'huiles contenant des métaux lourds et des précurseurs de coke Expired EP0065627B1 (fr)

Priority Applications (1)

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AT82101794T ATE19411T1 (de) 1981-04-28 1982-03-06 Immobilisierung von vanadin, das bei der umwandlung von schwermetalle und koksvorlaeufer enthaltenden oelen auf katalysatoren abgelagert worden ist.

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US06/258,265 US4377470A (en) 1981-04-20 1981-04-28 Immobilization of vanadia deposited on catalytic materials during carbo-metallic oil conversion
US258265 1994-06-10

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EP0065627A2 EP0065627A2 (fr) 1982-12-01
EP0065627A3 EP0065627A3 (en) 1984-01-11
EP0065627B1 true EP0065627B1 (fr) 1986-04-23

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JP (1) JPS5946994B2 (fr)
AT (1) ATE19411T1 (fr)
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ATE19647T1 (de) * 1981-10-13 1986-05-15 Ashland Oil Inc Verfahren zum passivieren von waehrend der verarbeitung von toprueckstaenden auf krackkatalysatoren abgesetzten metallen.
JP6387501B1 (ja) 2017-11-16 2018-09-12 アイ・エヌ製薬株式会社 屋内動物忌避装置および屋内動物忌避方法
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US4162213A (en) * 1976-04-29 1979-07-24 Mobil Oil Corporation Catalytic cracking of metal-contaminated oils
GB2001545B (en) * 1977-07-28 1982-04-15 Ici Ltd Hydrocarbon processing
EP0009842B1 (fr) * 1978-10-02 1982-11-10 THE PROCTER & GAMBLE COMPANY Liposomes pour la libération d'agents pharmaceutiques et composition contenant un système d'agent pharmaceutique à liposome
US4268416A (en) * 1979-06-15 1981-05-19 Uop Inc. Gaseous passivation of metal contaminants on cracking catalyst
US4341624A (en) * 1979-11-14 1982-07-27 Ashland Oil, Inc. Carbo-metallic oil conversion
US4295961A (en) * 1979-11-23 1981-10-20 Standard Oil Company (Indiana) Method and apparatus for improved fluid catalytic riser reactor cracking of hydrocarbon feedstocks
US4280895A (en) * 1979-12-31 1981-07-28 Exxon Research & Engineering Co. Passivation of cracking catalysts

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ATE19411T1 (de) 1986-05-15
CA1174192A (fr) 1984-09-11
MX161136A (es) 1990-08-07
EP0065627A2 (fr) 1982-12-01
EP0065627A3 (en) 1984-01-11
JPS5946994B2 (ja) 1984-11-16
DE3270718D1 (en) 1986-05-28
JPS57202382A (en) 1982-12-11

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