EP0171460A1 - Procédé de craquage d'huile résiduelle en utilisant du gaz sec tel que le gaz d'entraînement dans un réacteur à colonne montante - Google Patents

Procédé de craquage d'huile résiduelle en utilisant du gaz sec tel que le gaz d'entraînement dans un réacteur à colonne montante Download PDF

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EP0171460A1
EP0171460A1 EP84112717A EP84112717A EP0171460A1 EP 0171460 A1 EP0171460 A1 EP 0171460A1 EP 84112717 A EP84112717 A EP 84112717A EP 84112717 A EP84112717 A EP 84112717A EP 0171460 A1 EP0171460 A1 EP 0171460A1
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European Patent Office
Prior art keywords
catalyst
gas
cracking
riser
hydrogen
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EP84112717A
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German (de)
English (en)
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EP0171460B1 (fr
Inventor
Ronald A. Kmecak
William P. Hettinger, Jr.
Stephen M. Kovach
Larry M. Fraley
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Ashland LLC
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Ashland Oil Inc
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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
    • C10G47/00—Cracking of hydrocarbon oils, in the presence of hydrogen or hydrogen- generating compounds, to obtain lower boiling fractions
    • C10G47/24—Cracking of hydrocarbon oils, in the presence of hydrogen or hydrogen- generating compounds, to obtain lower boiling fractions with moving solid particles
    • C10G47/30—Cracking of hydrocarbon oils, in the presence of hydrogen or hydrogen- generating compounds, to obtain lower boiling fractions with moving solid particles 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
    • 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

Definitions

  • the invention relates to a novel and improved method for effecting the catalytic cracking of fractions of crude oil and particularly residual portions of crude oils comprising gas oils which may or may not comprise vacuum resid including asphaltenes, asphalt, substantial metal contaminants of nickel and vanadium, sulfur and nitrogen compounds.
  • the hydrocarbon feed may be a 343°C (650°F) plus carbometallic containing heavy oil feed providing very little or substantial amounts of Ramsbottom carbon or Conradson carbon contributing materials.
  • the crude oil fraction employed in the cracking operation of this invention may be a reduced crude, a topped crude, vacuum residues, heavy oil extracts of tar sands, a coal liquefaction product, oil product of oil shale pyrolysis and mixtures thereof.
  • a typical oil feed may boil from 343°C (650°F) up to 566°C (1050°F) or as high as 816°C (1500°F) when including vacuum bottoms.
  • the invention relates to improving the product selectivity obtained and maintaining desired equilibrium catalyst activity during the cracking of such heavy oil feeds with particularly an active crystalline zeolite containing catalyst.
  • the essence of the invention is achieved by initially forming an upflowing suspension of regenerated catalyst particles of desired high elevated temperature in a lift gas comprising a dry gas composition contributing- little, if any, coke to the catalyst suspension before effecting contact with the oil feed charged to a riser catalytic cracking operations.
  • the present invention is concerned with and relates to the field of hydrocarbon conversion disclosed in (docket 6034AUS) USSN 288952, filed May 13, 1981, now U.S. Patent 4432863 Myers et al; (docket 6049JUS) USSN 373599 filed April 30, 1982, now U.S. Patent 4419223; and USSN 411719 filed August 19, 1982, now U.S. Patent 4435279.
  • the invention also particularly relates to the combination operation of residual oil feed treatment to partially decarbonize and demetallize before effecting catalytic cracking, thereof, disclosed in USSN 550985, filed Nov.
  • US 2,888,395 to Henny contacts a heavy hydrocarbon with a catalyst in the presence of substantially pure hydrogen in a riser.
  • the hydrogen is produced outside the cracking unit.
  • the use of hydrogen with the oil feed is said to reduce coke make and to reduce the production of unsaturated products.
  • US 4,361,496 to Castillo treats regenerated-metal contaminated catalyst with a hydrocarbon gas comprising three carbon atoms or less, or a mixture thereof, to achieve complete reduction of contaminant metals which are carbonized in the dipleg located between the regenerator and the riser reaction zone.
  • This dipleg is used to convey catalyst from the regenerator to the riser cracking zone.
  • US 4,364,848 to Castillo is similar to US 4,361,496 above in that the reducing gas used is a mixture of one, two, three carbon atoms to passivate the metals to the metallic state before carbonization thereof.
  • Polack 2,937,988 discloses a riser reactor system for cracking an oil feed such as a heavy hydrocarbon residuum, vacuum or atmospheric crude bottom, pitch, asphalt or mixtures thereof wherein hot coke particles are initially dispersed in fluidizing gases such as steam, light hydrocarbons, an inert gas or mixtures thereof.
  • Bowles 3,406,112 discloses recovering a hydrocarbon product stream boiling below about C6 or C5 hydrocarbons which are employed in an amount sufficient to form a suspension with zeolite catalyst particles in a lower portion of a riser reaction zone before charging the oil feed in contact therewith.
  • Owen 3,849,291 discloses the use of dry gas or wet gas products of cracking as diluent material in the riser cracking operation disclosed.
  • Owen 3,894,932 discloses the use of a C3-C4 hydrocarbon gas mixture in the bottom portion of a riser to form an upflowing suspension of zeolite catalyst particles prior to contact with an oil feed.
  • Myers et al 4,431,515 (Ashland Oil, Inc.) is directed to a carbometallic oil conversion process using hydrogen in a riser reactor and comprising a high metals containing catalyst.
  • This patent discloses the addition of hydrogen to the riser reaction zone to reduce the formation of conjugated diolefins. Thus, it is postulated that the concentration of diolefins is reduced and coke production with the metals containing catalyst is also reduced.
  • the patent discusses the use of hydrogen gas admixed with the hydrocarbon feed provided by gas streams comprising 60 and 80% or more hydrogen. The patent is silent with respect to the essence of invention described in the present application.
  • gasoline and other liquid hydrocarbon fuels boil in the range of about 38°C (100°F) to about 343°C (about 650°F), however, the crude oil from which these fuels are made is a diverse mixture of hydrocarbons and other compounds which vary widely in molecular weight and therefore boil over a wider range.
  • crude oils are known in which 30% to 60% or more of the total volume is composed of compounds boiling at temperatures above 343°C (650°F).
  • 650°F Among these crudes are crudes in which about 10% to about 30% or more of the total volume consists of compounds which are so heavy in molecular weight that they boil above 552°C (1025°F), or at least will not boil below 552°C (1025°F) at atmospheric pressure.
  • the Fluid Catalytic Cracking (FCC) process was developed for cracking or breaking the molecules of high molecular weight, high boiling compounds into smaller molecules which boil over an appropriate boiling range.
  • FCC Fluid Catalytic Cracking
  • the FCC process has reached a highly advanced state, and many modified forms and variations have been developed, their unifying factor is that a vaporized hydrocarbon feedstock which contains high molecular weight, high boiling components is caused to crack at an elevated temperature in contact with a cracking catalyst that is suspended in the feedstock vapors.
  • the catalyst is separated from the desired products.
  • the present invention is concerned with using hydrocarbon feedstocks which have Ramsbottom carbon values which exhibit a substantially greater potential for coke formation than does the usual FCC feedstock.
  • Ramsbottom carbon values on the order of about 0.1 to about 1.0 are regarded as indicative of acceptable feed.
  • Conventional FCC practice has employed as feedstock that fraction of crude oil which boil at about 343°C (650°F) to about 538°C (1000°F), and is relatively free of coke precursors and heavy metal contaminants.
  • VGO vacuum gas oil
  • the various heavy metals in carbometallic oil are not of equal catalyst poisoning activity, it is convenient to express the poisoning activity of an oil containing a given poisoning metal or metals in terms of the amount of a single metal which is estimated to have equivalent poisoning activity.
  • the heavy metals content of an oil can be expressed by the following formula (patterned after that of W. L, Nelson in Oil and Gas Journal, page 143, Oct. 23, 1961) in which the content of each metal present is expressed in parts per million of such metal, -as metal, on a weight basis, based on the weight of feed.
  • Nickel Equivalents Ni+(V/4.8)+(Fe/7.1)+(Cu/1.23)
  • the above formula can also be employed as a measure of the accumulation of heavy metals on the cracking catalyst, except that the quantity of metal employed in the formula is based on the weight of catalyst (moisture free basis) instead of the weight of feed.
  • the present invention is concerned with the processing of feedstocks containing heavy metals substantially in excess of that in conventional FCC processing, and which therefore have potential for accumulating on and poisoning the catalyst.
  • the present invention is notable in providing a simple, relatively straightforward and highly productive approach to the conversion of oil feeds to various lighter products, such as gasoline.
  • the feedstock is comprised of oil which boils above about 343°C (650°F).
  • oil, or at least the 343°C+ (650°F) portion thereof is characterized by a heavy metals content of at least about 4, preferably more than about 5, and most preferably at least about 5.5 ppm Nickel Equivalents of heavy metals by weight and by a carbon residue on pyrolysis of at least about 1% and more preferably at least about 2% by weight or more.
  • a catalyst which contains from about 1000 to about 70,000 ppm of metals such as nickel, incremented iron, copper and/or vanadium or its oxides is referred to herein as a contaminated catalyst because such catalyst tends to encourage unless passivated the formation of coke during the cracking process. Thus, in most instances, the catalyst is continuously replaced to maintain these metals in low concentrations on the catalyst up to 20,000 ppm Ni + V.
  • a catalyst which contains these metals especially in the above concentration range, because they are considered to have the ability to activate hydrogen, when hydrogen is introduced in the cracking system.
  • a second reaction also based on LeChatalier's principle, is also inhibited by the addition of hydrogen to the catalytic cracking system.
  • Many of the metals referred to above return to the reactor as oxides and are undoubtedly reduced quickly to metals or lower valent oxides in the reactor by scavenging of the hydrogen produced in the above reaction.
  • the driving force for such reaction therefore, is also to the right in the above equation and this driving force may also encourage the formation of conjugated diolefins.
  • the catalyst is projected in a direction established by the elongated riser reaction zone or an extension, thereof, whereby vaporous products, having lesser momentum, are caused to make a change of direction, resulting in ballistic separation of products from catalyst, thus avoiding secondary cracking of products.
  • the separated catalyst is stripped to remove high boiling components and other entrained or absorbed hydrocarbons, and then regenerated by burning the coke in at least one regeneration zone with an oxygen-containing combustion-supporting gas under conditions of time, temperature and atmosphere sufficient to reduce the carbon on the regenerated catalyst to about 0.05% or less by weight.
  • the coke is burned in a zone wherein the molar ratio of CO:C0 2 is maintained at a level of at least about 0.20. to about 0.25, more preferably at least about 0.3 and still more preferably at least about 0.5.
  • the regenerated catalyst is stripped to remove entrained air, and thereafter recycled to the reactor for contact with fresh feed.
  • the present invention is directed to maintaining a special relationship in operating parameters in the hydrocarbon feed catalytic cracking step which thereby affects the severity of the catalyst regeneration required and employed.
  • a substantial improvement in hydrocarbon product selectivity is achieved at reduced coke make by the catalyst activity charged to the oil feed cracking operation.
  • the method of operation of this invention provides a catalyst which is of a higher order of activity than previously achieved.
  • the present invention is concerned with effecting more selective crystalline zeolite catalytic cracking of residual oil fractions of crude oils particularly comprising nickel and vanadium metal contaminants in an amount within the range of 1000 to 20,000 ppm in combination with high molecular weight polycyclic hydrocarbon materials contributing a Ramsbottom carbon value up to about 8 during catalytic cracking thereof.
  • the present invention is concerned with using a dry gas stream comprising hydrogen with some limited e 3 -plus carbon producing components therein under catalyst suspension forming conditions.
  • a combination of one or more regenerated catalyst cooling fluids such as steam, water and combinations thereof is used with the dry gas to form a rising suspension with hot regenerated catalyst particles adjusted to a temperature particularly suitable for effecting catalytic cracking of a residual oil feed comprising gas oils with or without higher boiling vacuum bottoms generally boiling above about 552°C (1025°F).
  • a gas stream suitable for the purpose and essence of this invention is a commercially available refinery product dry gas stream comprising less than 10 vol. % of C 3 plus hydrocarbons and preferably comprises hydrogen in an amount of at least 10 vol. %. Such a hydrogen containing gas is economically recoverable from one or more refinery gas streams.
  • the catalyst employed in the catalytic cracking-hydrocarbon conversion operation of the present invention may be substantially any fluid .
  • crystalline zeolite cracking catalyst of the prior art comprising rare earth and/or hydrogen ions in the crystal structure of the zeolite.
  • the zeolite is dispersed in a siliceous-clay matrix material which may or may not provide some cracking activity. That is, the matrix may be selected from silica-alumina, silica-zirconium or silica-chromium mixture which is promoted with one or more metal additives which are effective in passivating accumulated metal contaminants.
  • Some additive material which may be used include rare earth metals providing excess lanthanum, and compounds of antimony and titanium.
  • the cracking catalyst employed in the method of this invention may comprise the active crystalline zeolite component in an amount less than about 40 wt. % and more usually in an amount within the range of 5 to 20 wt. % as equilibrium catalyst.
  • the catalyst employed may be selected from one described in U.S. 4,440,868 or U.S. 4,435,515.
  • a preferred catalyst may be one selected from application USSN 483,061 filed April 7, 1983 (docket 6193AUS), each of which is incorporated herein by reference thereto.
  • a particularly preferred class of catalysts includes those that are capable of activating hydrogen and that have pore structures into which molecules of feed 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.
  • the zeolite-containing catalysts used in the present invention may include any zeolite, whether natural, semi-synthetic or synthetic, alone or in admixture with other materials which do not significantly impair the suitability of the catalyst, provided the resultant catalyst has the activity and pore structure referred to below.
  • the 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 about 1% to about 60%, more preferably about 1 to about 40% and most typically about 5 to about 40% by weight of the zeolite dispersed in the carriers, based on the total weight of catalyst (water free basis) of the porous refractory inorganic oxide alone or in combination with any of the known adjuvants for promoting or suppressing various desired and undesired reactions, some of which are discussed below.
  • catalysts having an overall particle size in the range of about 5 x 10 -6 meters to about 160 x 10 -6 meters (about 5 to about 160 microns), more preferably about 40 x 10 -6 meters to about 120 x 10 -6 meters (about 40 to about 120 microns), and containing a proportionately major amount in the 40 x 10 -6 meters to about 80 x 10 -6 meters (40 to about 80 microns) range.
  • a catalyst initially having a relatively high level of cracking activity and selectivity, and providing high levels of conversion and productivity at low residence times may be expressed in terms of the conversion produced during actual operation or by standard catalyst activity test. (See the classical Shankland and Schmitkons "Determination of Activity and Selectivity of Cracking Catalyst", Proc. API 2.7 (III), 1947, pp. 57-77).
  • catalysts which, in the course of extended operation in the process, are sufficiently active for sustaining a level of conversion of at least about 50% or more preferably at least about 60%. In this connection, conversion is expressed in liquid volume percent, based on fresh feed.
  • the preferred catalyst may also be defined as one which, in its virgin or equilibrium state, exhibits a specified activity expressed as a volume percentage derived by the MAT (micro-activity test).
  • MAT micro-activity test
  • the preferred catalysts When characterized on the basis of MAT activity, the preferred catalysts may be described on the basis of their MAT activity "as introduced” into the process of the present invention, or on the basis of their “as withdrawn” or equilibrium MAT activity, or on both of these bases.
  • a preferred MAT activity for virgin and non-virgin catalyst "as introduced" in the process of the present invention is at least about 60%, but it will be appreciated that, particularly in the case of non-virgin catalysts supplied at high addition rates, lower MAT activity levels may be acceptable.
  • An acceptable equilibrium MAT activity level of catalyst which has been used in the process of the present invention is above 20%, preferably at least about 40% or more preferably about 60% or more are preferred values.
  • the weight ratio of catalyst to fresh feed (feed which has not previously been exposed to cracking catalyst under cracking conditions) used in the present invention is in the range of about 3 to 18.
  • Preferred ratios may be about 4 to 12, depending on the coke forming tendencies of the feed.
  • controlling the catalyst to oil ratio at relatively low levels within the aforesaid ranges tends to reduce the coke yield of the oil, based on fresh feed.
  • Catalyst may be added continuously or periodically, such as, for example, to make up for normal losses of catalyst from the system. Moreover, catalyst addition may be conducted in conjunction with withdrawal of catalyst, such as, for example, to maintain or increase the average activity level of the catalyst in the unit or to maintain a constant amount of metal on catalyst.
  • the rate at which virgin catalyst is added to the unit may be in the range of about .285 kilograms per m 3 of feed (0.1 to about 3 lb/bbl) to about 8.55 kilgrams per cubic meter of feed or more (about 0.03 to 1 weight % of the feedstock) or more, depending on metal content in-the feed, and the level of metal allowed to reside on the equilibrium catalyst.
  • equilbrium catalyst is employed, a replacement rate as high as about 14.25 kilograms per cubic meter of feed (about 5 pounds per barrel) or more can be practiced. Where circumstances are such that the conditions in the unit tend to promote more rapid deactivation, one may employ rates of addition greater than those stated above; but in the opposite circumstances, lower rates of addition may be employed.
  • the invention may be practiced with catalyst bearing accumulations of heavy metals which heretofore would have been considered quite intolerable in conventional fluid catalytic cracking (FCC), vacuum gas oil (VGO) operations.
  • catalyst bearing heavy metals accumulations in the range of about 1000 to about 20,000 ppm Ni+V on the average, is contemplated.
  • the accumulation may also be in the range of about 4000 to 50,000 ppm and more likely in the range of 5000 to about 30,000 ppm.
  • the higher foregoing ranges are based on parts per million of heavy metal, including nickel, vanadium, incremental iron (that additional iron accumulated while being used) and copper, in which the metals are expressed as metal, by weight, measured on and abased on regenerated equilibrium catalyst, i.e.
  • the catalyst composition may also include one or more combustion promoters which are useful in the subsequent step of regenerating the catalyst.
  • combustion promoters which are useful in the subsequent step of regenerating the catalyst.
  • coke is burned off in a regeneration step, in which coke is converted to combustion gases including carbon monoxide and/or carbon dioxide.
  • Various substances e.g. Pt, Pd, rare earths, are known which, when incorporated into a cracking catalyst in small quantities (or added with the feed stock), tend to promote conversion of coke to carbon monoxide and/or carbon dioxide. Promoters of combustion to carbon monoxide tend to lower the temperature at which a given degree of coke removal can be attained, thus diminishing the potential for thermal deactivation of the catalyst.
  • Such promoters normally used in effective amounts ranging from a trace up to about 10% to 20% by weight of catalyst, may, for example, be of any type which generally promotes combustion of carbon under regenerating conditions.
  • the amount of additional materials which may be present in the feed may be varied as desired; but said amount will preferably be sufficient to substantially heat balance the- process.
  • These materials may for example be introduced into the reaction zone in a weight ratio relative to feed of up to about 0.4, preferably in the range of about 0.02 to about 0.4, more preferably about 0.03 to about 0.3 and most preferably about 0.05 to about 0.25.
  • a preferred embodiment is to have hydrogen sulfide dissolved therein within the above ranges, based on the total amount of feed. Alternately, about 500 ppm to about 5000 ppm of hydrogen sulfide should be dissolved in the recycled liquid water. Hydrogen sulfide gas, in the above weight ratio ranges, may also be added as the additional material instead of hydrogen sulfide dissolved in recycled liquid water.
  • the process of the present invention employees ballistic separation of catalyst and vapors at the downstream end of a progressive flow type riser, such as is taught in U.S. 4,066,533 and 4,070,159 to Myers et al, the disclosures of which are hereby incorporated by reference thereto.
  • the catalyst riser residence time may or may not be the same as that of the vapor.
  • the ratio of average catalyst reactor residence time versus vapor reactor residence time i.e. slippage, may be in the range of about 1 to about 5, more preferably about 1 to about 4, and most preferably about 1.1 to about 3, with about 1,2 to about 2 being the preferred range.
  • vapor riser residence time and vapor-catalyst contact time in the riser are substantially the same for at least about 80% of the riser length.
  • Regeneration of catalyst may be performed at a temperature in the range of about 593°C (1100°F) to about 871 0 C (1600°F), measured at the catalyst regenerator outlet.
  • This temperature may be in the range of about 649°C (1200°F) to about 816°C (1500°F), more preferably in the range of about 677°C (1250°F) to about 774°C (1425°F) and optimally about 704°C (1300°F) to about 746°C (1375°F) or about 760°C (1400°F).
  • a stripper which are sufficient to reduce potentially volatile hydrocarbon material borne by the stripped catalyst to about 10% or less by weight carried to the regenerator.
  • 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 FCC/VGO operations, such as for instance flue gas from the regenerator, as well as other refinery stream gases such as hydrotreater off-gas (H 2 S containing), hydrogen and others.
  • the stripper may be operated at a temperature above about 482°C (900°F). Stripping operations in which the temperature of the spent catalyst is raised to higher temperatures is also within the scope of the present invention.
  • coke should be understood to include any residual unvaporized feed or hydrocarbonaceous material present on the catalyst after stripping thereof.
  • the substantial levels of conversion accomplished by the process of the present invention result in relatively large yeilds of coke, such as for example about 4% to about 17% by weight based on fresh feed, more commonly about 6% to about 14% and most frequently about 6% to about 12%.
  • the resultant coke laydown may be in excess of about 0.3%, more commonly in excess of about 0.5% and very frequently in excess of about 1% of coke by weight, based on the weight of moisture free virgin or regenerated catalyst.
  • Such coke laydown may range as high as about 2%, or about 3%, or even higher, although coke in the range of 0.5 to about 1.5% is more commonly experienced.
  • the sub-process of regeneration may be carried out to the above-mentioned low levels of coke on regenerated catalyst with oxygen supplied to the one or more stages of regeneration in the stoichiometric amount required to burn all hydrogen in the coke to H 2 0 and to burn all carbon in the coke to CO and/or C0 2 and to burn all sulfur in the coke to S0 2 .
  • the coke includes other combustibles, the aforementioned stoichiometric amount can be adjusted to include the amount of oxygen required to burn them.
  • Multi-stage regeneration offers the technique of combining oxygen deficient regeneration with control of the CO:CO 2 molar ratio and still provide means by which coke on catalyst is reduced preferably to 0.05% or lower. Thus, about 65% to about 80% by weight of the coke on the catalyst is removed in a first stage of regeneration in which the molar ratio of CO:C0 2 is controlled.
  • the last weight percent of the coke originally present up to the entire amount of coke remaining after the preceding stage can be removed in a subsequent stage of regeneration in which more oxygen is present.
  • a particularly preferred embodiment of the present invention is two-stage catalyst regeneration at a maximum temperature of about 816°C (1500°F) but preferably not above 760°C (1400°F).
  • the second stage temperature is the same or lower than the first stage, with reduction of carbon on catalyst to about 0.05% or less or even about 0.025% or less by weight in the second zone.
  • catalyst can readily be regenerated to carbon levels as low as 0.01% by this technique, even though the carbon on catalyst prior to regeneration is as much as about 1% or greater.
  • 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 additional material, discussed herein.
  • the catalyst discharged from the regenerator is stripped with appropriate stripping gases to remove oxygen containing gases.
  • stripping may for instance be conducted at relatively high temperatures, using steam nitrogen or inert gas(es) as the stripping gas.
  • nitrogen or other inert gases is beneficial from the standpoint of avoiding a tendency toward hydrothermal catalyst deactivation which may result from the use of steam.
  • the present invention is applicable to the catalytic conversion of light gas oil feeds or heavy residual oil feeds comprising vacuum bottoms and portions thereof which have been subjected to a previous partial hydrogenation operation to remove sulfur and nitrogen compounds, therefrom, and/or which has been partially decarbonized and demetallized by contact with a sorbent material under thermal visbreaking conditions in the presence of a diluent with or without the presence of hydrogen.
  • the sorbent material employed in the visbreaking operation may be relatively inert or of such low catalytic activity that it is no longer suitable for use in a catalytic cracking operation.
  • the essence of this invention is useful in the disclosed combination operation of U.S. 4,434,044 Busch et al, the subject matter of which is incorporated, herein, by reference thereto.
  • the process conditions employed in the catalytic cracking operation of this invention will vary depending upon the composition and boiling range of the oil feed charged. Generally, the regenerated catalyst charged to the riser cracking operation will be at a temperature in the range of 649°C (1200°F) to 816°C (1500°F) and more usually about 704°C (1300°F) to 760 0 C (1400°F).
  • the catalyst to oil ratio and hydrocarbon feed partial pressure will vary with the feed boiling range and volume of gaseous diluent used so that vaporous hydrocarbon conversion products comprising suspended cracking catalyst, lift gas and feed atomizing diluent material will be discharged from the riser reactor cracking zone at a temperature within the range of 482°C (900°F) to 598°C (l100°F) and more usually within the range of about 510°C (950°F) to about 566°C (1050°F).
  • the present processing concept of invention to reduce coke make and improve product selectivity is applicable to the processing disclosure of U.S. 4,434,044 as above identified with modification thereto as required in a metals removing oil feed decarbonizing visbreaking operation with solid sorbent fluid particles followed by catalytic upgrading of the partially demetallized and decarbonized heavy oil feed by the technique of this invention.
  • An important aspect of the combination operation of this referenced patent is related to the light gaseous product recovery steps of figure I, wherein, a fuel gas is recovered from a C 3 -C 4 fraction. This fuel gas comprising hydrogen is particularly suitable when separated from the C 3 -C 4 hydrocarbons for use in accordance with the processing concepts of the present invention.
  • 4,434,044 is particularly incorporated, herein, by reference thereto.
  • the processing concepts of the present invention modify the cracking concepts of the referenced patent to the extent that dry gas without naphtha but comprising steam is used to form an upflowing suspension of the hot regenerated catalyst particle in a lower portion of the riser and, thereafter, contacted with oil feed to be upgraded by crystalline zeolite catalytic cracking as herein provided.
  • the appartus arrangement of Figure V of U.S. Patent 4,434,044 comprising a riser catalytic cracking zone adjacent to a sequence of two stage catalyst regeneration providing for cooling of catalyst passed from said first stage to said second stage of catalyst regeneration is preferably'modified to incorporate a riser reactor of larger diameter in an upper portion than in a lower portion thereof with the oil feed to be cracked being charged to a downstream section of the riser comprising the larger diameter portion, thereof.
  • This riser design is a part of U.S. 4,435,279 Busch et al (docket 6168AUS) and is thus incorporated herein by reference thereto.
  • Table I identifies an analysis of a commercially obtained wet gas employed in the cracking process of the invention. It is to be noted that the
  • C 3 -plus material up to and including C 5 hydrocarbon materials amounts to about 45%.
  • Table II identifies an analysis of a commercially obtained dry gas employed in the cracking process of the invention. It is to be noted that the C 3 -plus material up to and including N-Pentane is in an amount of about 8.68%. Thus, it will be recognized that the commercial processing of refinery gaseous products to obtain an inexpensive dry gas product comprising less than 8 or 10% of e 3 -plus material is economically difficult to achieve and, thus, of little or no interest to a petroleum refiner. It is desirable and essential on the other hand, to pass a suspension upwardly through a riser fluid catalyst cracking zone for contact with charged heavy oil feed. A light gaseous product of cracking is recovered in one aspect of this invention for recycle to the cracking operation and use, therein, as herein discussed.
  • carbometallic containing residual oil feeds comprising Ramsbottoms carbon, sulfur, nickel and vanadium were brought in contact with a typical fluid cracking catalyst comprising a rare earth exchanged crystalline aluminosilicate (faujasite) containing cracking catalyst following regeneration treatment thereof at an elevated temperature, herein defined with a dry gas hydrogen rich product comprising less than about 10% of C 3 -plus materials and a wet gas hydrogen rich product of catalytic cracking herein defined comprising substantial C 4 and C 5 hydrocarbons.
  • a typical fluid cracking catalyst comprising a rare earth exchanged crystalline aluminosilicate (faujasite) containing cracking catalyst following regeneration treatment thereof at an elevated temperature, herein defined with a dry gas hydrogen rich product comprising less than about 10% of C 3 -plus materials and a wet gas hydrogen rich product of catalytic cracking herein defined comprising substantial C 4 and C 5 hydrocarbons.
  • FIGURE II COKE PRODUCTION
  • Figure II is a further plot of the experimental data obtained showing the coke production obtained when converting a residual oil to 221°C (430°F) minus product in the presence of catalyst initially contacted with dry recycle gas or wet recycle gas employed as lift gas in, for example, a riser cracking zone. It will be observed from the plot of Figure II that the use of a hydrogen rich wet recycle gas comprising C 4 and C 5 hydrocarbons in substantial amounts produced considerable more coke in the catalyst than was obtained when using a hydrogen rich dry recycle gas. The high coke deposition contributes to obtaining high catalyst regeneration temperatures exceeding 760°C (1400°F).
  • Figure III identifies the C 5 to 221°C (430°F) gasoline yield provided by the experimental data when using a dry gas or a wet gas as a catalyst lift gas before contacting hot regenerated catalyst with the residual oil feed charged. It is significant to note from this figure that the use of dry gas as a lift gas provided higher yield of gasoline product than was obtained when using the wet gas as a lift gas. Thus, the gasoline product selectivity is considerably improved.
  • FIGURE IV GASOLINE SELECTIVITY
  • Figure IV identifies from a plot of the experimental data, an improved gasoline selectivity curve obtained when using a dry gas to initially contact the hot regenerated catalyst over that obtained when using a hydrogen rich wet recycle gas, herein identified prior to converting the charged residual oil feed.
  • Figure V identifies the improved results obtained in the 332°C (630 * F) plus slurry oil product obtained when employing dry and wet recycle gas product as herein defined.
  • the figure shows that the yield of slurry oil is much less when using a dry recycle gas stream herein defined to initially contact the hot regenerated catalyst.
  • FIGURE VI INCREASED LIGHT CYCLE OIL YIELD
  • Figure VI shows another unpredicted aspect of the invention with respect to the increased light cycle oil yield (LCO) which was obtained when plotting the experimental data obtained.
  • LCO light cycle oil yield
  • This figure shows the compound LCO yield obtained when using a dry recycle gas in lieu of a wet recycle gas stream.
  • This plot of data clearly shows that as the composition of the lift gas changes in its coke make tendency, there is a directional reduction in the deposition of total coke make irrespective of Ramsbottom coke level.
  • Figure VII provides a plot of data obtained for different Ramsbottom carbon oil feeds conversion which identifies the total coke make obtained for different hydrogen containing gas feeds used with the conversion catalyst.
  • the lift gas comprises a significant quantity of C 3 plus material comprising C 5 hydrocarbons- which are cracked to deposit coke on the hot freshly regenerated catalyst prior to contact with the residual oil feed, thereby reducing the catalyst cracking activity and selectivity as shown by the above discussed figures, this contributes to a resultant loss in C 5 plus gasoline product material evaluated to amount to at least 3 to 5 vol.% of desired gasoline forming product material.
  • the resultant precoked catalyst is found to provide increased slurry oil and coke make as shown by the graphs presented as self explanatory.
  • the expanded or larger diameter portion of the riser 2 is provided with a plurality of feed inlet nozzles means 6 adjacent the upper edge of the transition section which are used in a preferred embodiment to charge the oil feed.
  • the vertically spaced apart feed inlet means 5, 7 and 9 provides the operator considerably more latitude in feed contact time with the dry gas-catalyst suspension within the riser reactor before separation of a resultant formed suspension of hydrocarbon product vapors, catalyst and lift gas available as herein discussed.
  • the riser 1-2 configuration of Figure VIII permits achieving relatively high temperature zeolite catalytic upgrading of an oil feed charged to a bottom, intermediate or upper portion of the riser conversion zone but downstream of the formed dry gas-regenerated catalyst suspension to restrict the oil feed contact time with catalyst within the range of a fraction of a second up to 1, 2 or even 3 seconds contact time.
  • the hot regenerated catalyst at a temperature within the range of 649°C (1200°F) to 816°C (1500°F) is initially mixed with a dry lift gas or fluidizing gas as herein provided with the addition of steam and/or water as heat sink material to form an upflowing suspension in the restricted diameter portion thereof at a temperature suitable for effecting catalytic cracking of a downstream charged residual oil feed as by 7 or 9.
  • feed inlet means a 5, 7 and 9 with diluent inlets 6, 8 and 10 permit a substantial variation in feed atomization and partial pressure and contact time as above identified between oil feed and the dry gas-steam suspended catalyst particles.
  • a bottom portion of the riser reactor permits adjustment of the regenerated catalyst temperature by the addition of steam and/or water as a heat sink along with the dry lift gas of a composition particularly identified herein.
  • the contact time between a residual oil feed and catalyst in the riser depending on feed composition and source will be restricted to within the range of 0.5 to about 2 or 3 seconds when contacting an oil feed with catalyst at a temperature in the range of 704°C (1300 0 F) to 760°C (1400°F) to provide a riser outlet temperature within the range of 510°C (950°F) to 593°C (1100°F) and more usually not above 566°C (1050°F).
  • the riser reactor may be substantially any desired vertical length which will be compatable with the adjacent catalyst regeneration appartus whether of single or multiple stages of regeneration as shown, catalyst stripping and catalyst transfer conduit means essential to the combination.
  • riser 2 passes upwardly through a stripping zone 6 to form an annular stripping zone therewith into an upper portion of a larger diameter catalyst disengaging zone in open communication with the annular stripping zone 16.
  • Stripping gas such as steam or other suitable gas is charged to a bottom portion of the stripping zone by conduit 17 for flow upwardly therethrough and counter-current to downflowing catalyst particles.
  • the stripped catalyst is then passed by conduit 19 to catalyst regeneration shown as a sequence of catalyst beds 20 and 36 being regenerated in separate zones to remove carbonaceous deposits of conversion by combustion without exceeding an elevated temperature below about 816°C (1500°F) and preferably restricted to within the range of about 649°C (1200°F) to 816°C (1500°F) and more usually within the range of 704°C (1300 0 F) to 760°C (1400°F).
  • An important aspect of the riser system of this invention is particularly concerned with the method and means for separting the upwardly flowing suspension at the riser upper open end. That is, the suspension of hydrocarbon vapors, catalyst, lift gas and steam is discharged from the upper open end of the riser at a velocity which will impart a greater momentum to the particles of catalyst than to that imparted to the vaporous constituents whereby an upwardly flowing trajectory is established which separates catalyst particles from vaporous material.
  • the vaporous material mixture often referred to as gasiform material in the prior art, passes into an annular cup 11 withdrawal passageway open in the top thereof and thence through radiating conduit means in open communication with cyclone separation means 12 on the outer end of each of said radiating conduits.
  • Vapors separated from entrained catalyst fines in cyclones 12 are recovered by conduits communicating with plenum chamber 13 and product withdrawal conduit 14 for passage to product fractionation and separation in means not shown. Catalyst fines separated in cyclones 12 are removed by diplegs for passage to catalyst stripping and regeneration discussed below.
  • the hydrocarbon conversion operation contemplated to be accomplished in the riser zone herein discussed relies upon the use of fluidizable particles of catalyst of a particle size in excess of 10 x 10 -6 meters (10 microns) and usually providing an average particle size within the range of 60 to 100 x 10 meters (60 to 100 microns) and more usually below about 85 x 10 meters (85 microns).
  • the catalyst is preferably one comprising a crystalline alumisilicate or crystalline zeolite which has been rare earth and/or ammonia exchanged to provide a catalytically active material which is dispersed in a matrix material which may or may not have catalytic activity.
  • a catalyst particularly suitable for use in the process of this invention is a rare earth exchanged faujasite crystalling zeolite comprising a catalyst pore volume and matrix pore size openings which will collect and/or accumulate substantial quantities of metal contaminants and yet retain substantial catalyst cracking activity and selectivity as herein provided.
  • the oil feed such as a residual portion of crude oil charged by feed inlet 5 or 7 may be mixed with steam and/or water such as product sour water charged by conduits 6 or 8.
  • the steam-water mixture may be added by conduit 10.
  • the bottom portion of riser 2 is provided with dry lift gas inlet conduit 4 for charging the lift gas to form a upflowing suspension with hot regenerated catalyst particles charged to a bottom portion of the riser by conduit 3.
  • the dry lift gas may be charged to the riser alone or in combihation with steam and/or water introduced by conduit 43.
  • the lower portion of the riser of restricted diameter may be used to serve seftveral different functions beyond the formation of an upflowing suspension of a desired catalyst particle concentration within the range of 16 to 44 kilograms per cubic meter. That is, the use of a hydrogen containing dry gas herein identified as lift gas may be used as a contaminant metals passivation material to which a passivating metal compound is added to passivate Ni and V. Antimony may be added to passivate accumulated nickel deposits. Vanadium oxide may be passivated by the combination of hydrogen reduction to a lower oxide state providing a high melting point oxide thereof alone or in conjunction with. the addition of titanium, alumina and rare earth metals rich in lanthanum.
  • lift gas may be used as a contaminant metals passivation material to which a passivating metal compound is added to passivate Ni and V.
  • Antimony may be added to passivate accumulated nickel deposits.
  • Vanadium oxide may be passivated by the combination of hydrogen reduction to a lower oxide
  • a hydrogen containing product recycle dry gas be of a composition which severely limits the C 3 plus components of the dry gas to a level inhibiting any significant coking of the catalyst therewith and prior to contact with the heavy oil feed to be cracked.
  • restricting the hydrogen, containing dry gas to a C 3 plus content less than 10%, more preferably less than 8% and most preferably less than 6% improves the gasoline yield, reduces the yield of hydrogen, increases the yield of light cycle oil and reduces the yield of slurry oil and coke.
  • a dry gas product of the cracking operation comprising at least 15 vol. percent hydrogen, less than 10 vol.% of C 3 plus hydrocarbons in admixture with water in an amount sufficient to partially cool the regenerated catalyst to a desired low oil feed conversion level before contact with atomized preheated residual oil charged to the rising dry gas-steam-catalyst suspension.
  • the fluid catalytic cracking of the charged hydrocarbons is effected at a riser pressure above atmospheric pressure and the riser cracking operation of this invention may be effected at a pressure of about 172 x 10 3 to 1,137 x 10 3 Pascals (about 10 to 150 psig) pressure.
  • the atomized oil feed hydrocarbon partial pressure will be substantially reduced by the lift gas-steam mixture and the oil feed atomizing diluent material.
  • the oil feed partial pressure may be in the range of 27.6 to 172 x 10 3 Pascals and the catalyst to oil ratio may be within the range of about 5 to 15, more preferably 6 to 12, and providing for intimate contact between catalyst particles and the atomized oil feed.
  • the combustion apparatus of Figure VIII provides a unique catalyst particle regeneration arrangement permitting close temperature control to minimize particularly hydrothermal deactivation of catalyst particles during the removal of coke deposits by combustion and contributed particularly by gas oil catalytic conversion and/or higher boiling components of residual oil including vacuum resid.
  • the upper chamber portion thereof is of a larger diameter than a bottom chamber portion and separated from one another by a regeneration gas distributor chamber 24 centrally located and supported by an annular baffle member 40 provided with gas flow through passage ways 41.
  • a plurality of radiating arm means 25 from chamber 24 are provided for introducing regeneration gas to a lower bottom portion of catalyst bed 20 being regenerated.
  • Regeneration combustion supporting gas such as air or an oxygen modified gas in conduit 22 admixed with steam in conduit 23 provides a desired concentration of oxygen and amounting to more or less than that requried to achieve a partial removal of carbonaceous deposits from the charged catalyst particles whereby combustion temperatures encountered can be restricted to within a desired range are charged by plenum 24 and radiating arms 25.
  • the regeneration temperature is preferably kept to a low value in the range of 593°C (1100°F) to 871°C (1600°F), preferably 649°C (1200°F) to 815°C (1500°F) and more usually in the range of about 690°C (1275°F) to 760°C (1400°F).
  • a partial removal of carbonaceous material is removed in catalyst bed 20 under conditions producing CO rich containing product flue gases and comprising carbon dioxide, sulfur, nitrogen and water vapor.
  • The, thus, generated flue gases pass through one or more combination of cyclones represented by cyclones 26 to remove entrained catalyst fines recovered by diplegs provided.
  • the flue gases then pass from cyclones 26 to a plenum chamber 27 for recovery therefrom by conduit 28.
  • Such CO rich containing flue gases are normally passed to a CO boiler not shown to generate process steam.
  • the partially regenerated catalyst comprising bed 20 is removed from a bottom portion thereof for downflow through an external catalyst cooling zone 29 in indirect heat exchange with bayonnet type heat exchange tubes 30 provided and substantially vertically extending therein.
  • High pressure steam of the order of about 3.1 x 10 6 Pascals (450 pounds) steam is generated and recovered as by conduit 34 when charging boiler feed water by conduit 31 to a distributor chamber in the bottom of cooler 29 communicating with said heat exchange tubes 30.
  • the catalyst partially cooled in chamber 29 by an amount in the range of 28°C to 111°C (50°F to 200°F) and more usually in the range of 55 to 83°C (100 to 150°F) is withdrawn and passed by conduit 35 to a bed of catalyst 36 retained in the second stage of catalyst regeneration in chamber 37.
  • a stand pipe 42 communicating between bed 20 and 36 is provided for direct passage of catalyst without cooling from the upper bed to the lower bed when required.
  • the main or primary flow of catalyst between beds is through cooler 29 to maintain desired catalyst temperature restraints in the sequential regeneration system.
  • a temperature restraint in the second stage comprising bed 36 is restricted within the range of 649°C (1200°F) to 816 0 C (1500 0 F) and more usually within the range of 704°C (1300 0 F) to 760°C (1400°F).
  • the temperature of the regenerated catalyst in dense fluid bed 36 may be equal to, above or below the temperature maintained in dense fluid catalyst bed 20 in the first stage of catalyst regeneration.
  • the amount of air or oxygen modified gas charged to catalyst bed 36 by conduit 38 and passing through grid 39 may be equal to or more than that required to complete combustion of residue carbon on the partially regenerated catalyst and provide a CO 2 rich flue gas product which may or may not comprise some unconsumed oxygen. It is preferred that the flue gas passed from the upper dense phase of catalyst bed 36 be free of combustion supporting amounts of CO to prevent after burning from occurring therein.
  • the C0 2 rich flue gas product of the second stage of catalyst regeneration at an elevated temperature passes through openings 41 in baffle 40 into a bottom portion of bed 20 for admixture with the regeneration gas charged by distributor arms 25 thereby contributing heat to the first stage of catalyst regeneration.
  • a velocity of about 24.5 meters/sec (80 ft/sec) the suspension traverses the riser in about 2 seconds.
  • the dry gas-steam-catalyst suspension initially formed consumes a residence time of a fraction of a second up to 0.5 second before contact with the atomized oil feed and providing a hydrocarbon residence contact time with catalyst particles up to about 1 or 1.5 seconds.
  • the short residence times identified are not detrimental to the process and may be used with considerable advantage to maintain desired product selectivity by reducing any tendency of over-cracking to occur.

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EP84112717A 1984-06-13 1984-12-17 Procédé de craquage d'huile résiduelle en utilisant du gaz sec tel que le gaz d'entraînement dans un réacteur à colonne montante Expired EP0171460B1 (fr)

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US4752375A (en) * 1986-09-03 1988-06-21 Mobil Oil Corporation Single riser fluidized catalytic cracking process utilizing a C3-4 paraffin-rich co-feed and mixed catalyst system
EP0280724A1 (fr) * 1986-09-03 1988-09-07 Mobil Oil Corporation Traitement d'hydrocarbures lourds actives d'alimentation
US4802971A (en) * 1986-09-03 1989-02-07 Mobil Oil Corporation Single riser fluidized catalytic cracking process utilizing hydrogen and carbon-hydrogen contributing fragments
US4826586A (en) * 1986-09-03 1989-05-02 Mobil Coil Corporation Single riser fluidized catalytic cracking process utilizing a C3-4 paraffin-rich co-feed and mixed catalyst system
EP0323297A1 (fr) * 1987-12-30 1989-07-05 Société Anonyme dite: COMPAGNIE DE RAFFINAGE ET DE DISTRIBUTION TOTAL FRANCE Procédé de conversion d'hydrocarbures en lit fluidisé
US4853105A (en) * 1986-09-03 1989-08-01 Mobil Oil Corporation Multiple riser fluidized catalytic cracking process utilizing hydrogen and carbon-hydrogen contributing fragments
US4863585A (en) * 1986-09-03 1989-09-05 Mobil Oil Corporation Fluidized catalytic cracking process utilizing a C3-C4 paraffin-rich Co-feed and mixed catalyst system with selective reactivation of the medium pore silicate zeolite component thereofo
FR2628436A1 (fr) * 1988-03-10 1989-09-15 Total France Procede et dispositif de vapocraquage d'hydrocarbures en phase fluidisee de particules caloporteuses
US4957617A (en) * 1986-09-03 1990-09-18 Mobil Oil Corporation Fluid catalytic cracking
US5506365A (en) * 1987-12-30 1996-04-09 Compagnie De Raffinage Et De Distribution Total France Process and apparatus for fluidized-bed hydrocarbon conversion
EP1046696A2 (fr) * 1999-04-23 2000-10-25 China Petrochemical Corporation Procédé de conversion catalytique pour la production d'essence enrichie en isobutane et en isoparaffines
WO2007149921A1 (fr) * 2006-06-22 2007-12-27 Shell Oil Company Procédés de production d'un produit brut à partir d'une charge d'alimentation sélectionnée
US7678342B1 (en) * 1999-04-23 2010-03-16 China Petrochemical Corporation Riser reactor for fluidized catalytic conversion
US8263008B2 (en) 2008-12-18 2012-09-11 Uop Llc Apparatus for improving flow properties of crude petroleum
US8608944B2 (en) 2005-12-23 2013-12-17 Research Institute Of Petroleum Processing Sinopec Catalytic conversion method of increasing the yield of lower olefin
US8951406B2 (en) 2011-07-29 2015-02-10 Saudi Arabian Oil Company Hydrogen-enriched feedstock for fluidized catalytic cracking process
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WO2024059600A1 (fr) * 2022-09-14 2024-03-21 Dow Global Technologies Llc Procédés de déshydrogénation d'hydrocarbures utilisant des régénérateurs

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FR2624762B1 (fr) * 1987-12-21 1990-06-08 Total France Procede et dispositif de regeneration de catalyseur en lit fluidise
JP4859358B2 (ja) * 2004-09-22 2012-01-25 日揮触媒化成株式会社 接触分解ガソリンの脱硫触媒およびそれを用いた接触分解ガソリンの脱硫方法
DE112008002718T5 (de) * 2007-10-10 2010-09-09 Shell Internationale Research Maatschappij B.V. Systeme und Verfahren zum Herstellen eines Mitteldestillatproduktes und niedere Olefine aus einem Kohlenwasserstoffeinsatzgut
RU2474605C2 (ru) * 2007-11-29 2013-02-10 Шелл Интернэшнл Рисерч Маатсхаппий Б.В. Установки и способы для получения среднедистиллятного продукта и низших олефинов из углеводородного исходного сырья
FR2953851B1 (fr) * 2009-12-14 2012-12-21 Total Raffinage Marketing Procede de craquage catalytique avec maximisation des bases gazoles
FR2966160B1 (fr) * 2010-10-14 2013-11-15 IFP Energies Nouvelles Procede de craquage catalytique adapte au traitement de charges a faible carbon conradson comportant le recycle d'une coupe cokante selon une technologie nouvelle
CN109297742B (zh) * 2018-11-21 2024-03-12 上海市计量测试技术研究院 餐饮油烟发生器

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EP0280724A4 (fr) * 1986-09-03 1989-10-25 Mobil Oil Corp Traitement d'hydrocarbures lourds actives d'alimentation.
US4752375A (en) * 1986-09-03 1988-06-21 Mobil Oil Corporation Single riser fluidized catalytic cracking process utilizing a C3-4 paraffin-rich co-feed and mixed catalyst system
EP0280724A1 (fr) * 1986-09-03 1988-09-07 Mobil Oil Corporation Traitement d'hydrocarbures lourds actives d'alimentation
US4802971A (en) * 1986-09-03 1989-02-07 Mobil Oil Corporation Single riser fluidized catalytic cracking process utilizing hydrogen and carbon-hydrogen contributing fragments
US4826586A (en) * 1986-09-03 1989-05-02 Mobil Coil Corporation Single riser fluidized catalytic cracking process utilizing a C3-4 paraffin-rich co-feed and mixed catalyst system
US4966681A (en) * 1986-09-03 1990-10-30 Mobil Oil Corporation Multiple riser fluidized catalytic cracking process utilizing a C3 -C4 paraffin-rich co-feed and mixed catalyst system
EP0259156A1 (fr) * 1986-09-03 1988-03-09 Mobil Oil Corporation Procédé de craquage catalytique fluidisé avec des fragments réactifs
US4853105A (en) * 1986-09-03 1989-08-01 Mobil Oil Corporation Multiple riser fluidized catalytic cracking process utilizing hydrogen and carbon-hydrogen contributing fragments
US4863585A (en) * 1986-09-03 1989-09-05 Mobil Oil Corporation Fluidized catalytic cracking process utilizing a C3-C4 paraffin-rich Co-feed and mixed catalyst system with selective reactivation of the medium pore silicate zeolite component thereofo
US4957617A (en) * 1986-09-03 1990-09-18 Mobil Oil Corporation Fluid catalytic cracking
FR2625509A1 (fr) * 1987-12-30 1989-07-07 Total France Procede et dispositif de conversion d'hydrocarbures en lit fluidise
EP0323297A1 (fr) * 1987-12-30 1989-07-05 Société Anonyme dite: COMPAGNIE DE RAFFINAGE ET DE DISTRIBUTION TOTAL FRANCE Procédé de conversion d'hydrocarbures en lit fluidisé
US5506365A (en) * 1987-12-30 1996-04-09 Compagnie De Raffinage Et De Distribution Total France Process and apparatus for fluidized-bed hydrocarbon conversion
FR2628436A1 (fr) * 1988-03-10 1989-09-15 Total France Procede et dispositif de vapocraquage d'hydrocarbures en phase fluidisee de particules caloporteuses
US7678342B1 (en) * 1999-04-23 2010-03-16 China Petrochemical Corporation Riser reactor for fluidized catalytic conversion
EP1046696A3 (fr) * 1999-04-23 2001-01-03 China Petrochemical Corporation Procédé de conversion catalytique pour la production d'essence enrichie en isobutane et en isoparaffines
EP1046696A2 (fr) * 1999-04-23 2000-10-25 China Petrochemical Corporation Procédé de conversion catalytique pour la production d'essence enrichie en isobutane et en isoparaffines
US8608944B2 (en) 2005-12-23 2013-12-17 Research Institute Of Petroleum Processing Sinopec Catalytic conversion method of increasing the yield of lower olefin
WO2007149921A1 (fr) * 2006-06-22 2007-12-27 Shell Oil Company Procédés de production d'un produit brut à partir d'une charge d'alimentation sélectionnée
US8263008B2 (en) 2008-12-18 2012-09-11 Uop Llc Apparatus for improving flow properties of crude petroleum
US9157037B2 (en) 2008-12-18 2015-10-13 Uop Llc Process for improving flow properties of crude petroleum
US8951406B2 (en) 2011-07-29 2015-02-10 Saudi Arabian Oil Company Hydrogen-enriched feedstock for fluidized catalytic cracking process
WO2021262639A1 (fr) * 2020-06-25 2021-12-30 Saudi Arabian Oil Company Procédé de valorisation d'huile lourde à l'aide d'hydrogène et d'eau
US11286429B2 (en) 2020-06-25 2022-03-29 Saudi Arabian Oil Company Process for heavy oil upgrading utilizing hydrogen and water
WO2024059600A1 (fr) * 2022-09-14 2024-03-21 Dow Global Technologies Llc Procédés de déshydrogénation d'hydrocarbures utilisant des régénérateurs

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JPH0226663B2 (fr) 1990-06-12
CA1265464A (fr) 1990-02-06
AU588528B2 (en) 1989-09-21
MX165471B (es) 1992-11-12
ATE36554T1 (de) 1988-09-15
AU4311185A (en) 1985-12-19
RU2091433C1 (ru) 1997-09-27
JPS614785A (ja) 1986-01-10
DE3473473D1 (en) 1988-09-22
EP0171460B1 (fr) 1988-08-17
IN162877B (fr) 1988-07-16

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