CA1206112A - Process for cracking high metals content feedstocks - Google Patents
Process for cracking high metals content feedstocksInfo
- Publication number
- CA1206112A CA1206112A CA000417682A CA417682A CA1206112A CA 1206112 A CA1206112 A CA 1206112A CA 000417682 A CA000417682 A CA 000417682A CA 417682 A CA417682 A CA 417682A CA 1206112 A CA1206112 A CA 1206112A
- Authority
- CA
- Canada
- Prior art keywords
- catalyst
- cracking
- diluent
- cracking catalyst
- alumina
- Prior art date
- Legal status (The legal status is an assumption and is not a legal conclusion. Google has not performed a legal analysis and makes no representation as to the accuracy of the status listed.)
- Expired
Links
Classifications
-
- 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
- C10G11/02—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils characterised by the catalyst used
- C10G11/04—Oxides
-
- 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
- C10G11/02—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils characterised by the catalyst used
- C10G11/04—Oxides
- C10G11/05—Crystalline alumino-silicates, e.g. molecular sieves
Abstract
PROCESS FOR CRACKING
HIGH METALS CONTENT FEEDSTOCKS
Abstract of the Disclosure A process for cracking high metals content feedstocks which comprises contacting said charge stock under catalytic cracking conditions with a novel cata-lyst composition comprising a solid cracking catalyst and a diluent containing a selected alumina or selected alumina in combination with one or more heat-stable metal compounds.
HIGH METALS CONTENT FEEDSTOCKS
Abstract of the Disclosure A process for cracking high metals content feedstocks which comprises contacting said charge stock under catalytic cracking conditions with a novel cata-lyst composition comprising a solid cracking catalyst and a diluent containing a selected alumina or selected alumina in combination with one or more heat-stable metal compounds.
Description
`~` lZO611~
PROCESS FOR CRACKING
HIGH METALS CONTENT FEEDSTOCKS
This invention relates to a process for cracking high metals content feedstocks which comprises contacting said charge stock under catalytic cracking conditions with a novel catalyst composition comprising a solid cracking catalyst and a diluent containing a selected alumina or a selected alumina in combinatio~
with one or more heat-stable metal compounds.
U.S. Pat. No~ 3,944,4a2 to Mitchell et al.
~ discloses a process directed to the catalytic cracking ;of hydrocarbon feeds containing m tals using a~ fluid c~atalyst having improved metals~ tolerant character-istics. Bartholic in U.S. Pat.~No. 4/289,60~5 xelates to a process for the catalytic cracking of`~hydrocarbon feeds containing metals using a catalyst composition containing a solid cracking~ catalyst and calcined microspheres (for example~, calcined;~kaolin clay) havin~
a surface area within the; range~ of ~10 to 15 m2/gram.
We have found that catalytic cracking of high metals content feedstocks 6uch as, for example, those containing iron, vanadium, nic~el and copper, can be S substantially improved by contacting said charge stocks under catalytic cracking conditions wi~h a novel catalyst composition compri6ing a solid cracking catalyst a~d a diluent selected from the group co~sist-ing of alumina (~1203) and alumina in co~bina~ion with one or more heat-stable metal compounds. The improve-ment resides in ~he ability of the cataly~t system to unction well even when the catalyst carries a substan-tially high level of metal on i~s surface, for example, up to 5000 ppm of ni~kel or nickel equivale~ts, or eYen higher, ox up t~ 20,000 ppm o vanadium. By "ppm of nic~el eguivalent" we mean p~m ~ick 1 ~ O.20 ppm vanadium. Thus feeds~ock~ having very high metals content can be sati~f~ctorily u~ed herein.
Thus accordi~g to the pre~ent invention there is provided a process for the catalytic cracki~g o~ a high me~als content charge stock which comprises contacting said charge stock under ca~alytic cr cking conditions with a catalyst ccmposition comprising a cracking catalyst having high activity and a diluent selected f~om ~he group co~sisting of alumina and alumina in combina ion with a heat-stable metal ~ompound, said diluen~ having a surface area of ,about 30 ~o abou~ lOOQ m~/gr~ and a pore volume of about 0.05 to about 2.5 cc/gram.
.
,, '.
-- ~%~
- 2a -The cracking catalyst component of the novel catalys-t composition used in -the novel process herein can be any cracking catalyst of any desired -type having high activity. By ~high activity" we mean catalyst of fresh MAT Activity above about 1.0, preferably up to a~out 4.0, or even higher, where Activity = Wt % Conversion 100 Wt % Converslon The "MAT Activity" was obtained by the use of a microactivi-ty test (MAT) unit similar to the standard Davison MAT (see Ciapetta et al., Oil & Gas Journal, 65, 88 (1967). A11 catalyst samples were tested at 900F.
(482C.~ reaction temperature; 15 weight hourly space velocity; 80 seconds of catalyst contact time; and a catalyst to oil ratio o~ 2.9 with 2O5 grams o~ catalyst.
The charge stock was a Kuwait gas oil havin~ a boiling range of 500F. to 800F. (260C. to 427C.). Inspec-tions o this Kuwait gas oil are shown i~ Table I below.
TABLE I
.
KUW~IT GAS OIL INSPECTIONS
_ Stock MAT
Identification Feedstock Inspections:
Gravity, API, D-287 23.5 Viscosity, SUS D2161, 130F. 94.7 Viscosity, SUS D2161, 150F. 70.5 Viscosi~y, SUS ~2161, 210~. 50.8 Pour Point, D97, F. ~80 Nitrogen, wt % 0.Q74 Sulfur, wt % 2.76 Carbon, Res., D524, wt % 0.23 Bromine No., D1159 5.71 Aniline Point, F. 176.5 Nickel, ppm <0.1 Vanadium, ppm <0.1 Distillation, ~1160 at 760 mm End Point, F. aoo 5 Pct. Cond. 505 Approx. Hydrocarbon Type AnaIysis: Vol. %
Carbon as Aromatics 23.1 Carhon as Naphthenes 10.5 Carbon as Paraffins 66.3 ~ _ Thus, catalytic ~racking catalysts suitable for use herein as host catalyst include amorphous silica~alumina cataly6t6; 6yn~hetic mica-montmorillonite catalysts as defined, for example in U.S. Pat. No.
3~252,889 to Capell et al.; and cross-linked clays (see, for example, Vaughn et al. in U.S~ Pat. Nos.. 4,176,090 and 4,248,739; Vaug~n et al. ~1980~, "Prepar~tion of Molecular Sieves Based on ~illared Interlayered Clays";
Proceedings of the 5th _ternational Conference on 2eolites, Rees, L.V., Editor, Heyden, London, pages 94-101; and Lahav et al., (1978) "Crosslinked Smectites I Synthesis and Properties of ~Iydroxy Aluminum Montmox-illonite'l, ClaY ~ Clay Minerals, 26, p`ages 107-114;
Shabtai, ~. in U.S. PatO No. ~,238,364; and 5habria et al. in U.S. Pat. No. 4,216,188~.
Preferably, ~h~ host catalyst used herein is 15 a catalyst containing a crystalline alwninosilica~e, preferably exchanged wi th rare earth metal cations, sometimes referred to as "raxe earth-exs:hang~d crystal-line. aluminum silicate" or one of the stabilized hydroyen ~eolites. ~os~ preferably, the host ~atalyst 20 i8 a high activit~ cracking cataly~t.
Typical zeolites or molecular siev2~ having cracking activity which can be used herei~ a~ a cata-lytic crac~cing ca~aly~t are well known in ~e art.
Suitable zeoli~es are de~cribed, for exampleO in U. S .
Pat~ No. 3,660,274 ~o Blazek et al., or in U.S. Pat. No.
3, 647, 718 to ~ayden et al .
Syn~hetically prepared zeolite~ are i~itially in the Iorm of alkali metal aIumino6ilicate~. The alkali ~etal ions are exchanged with rare earth metal ions to impart cracking charac~eristics to the zeolites. The zeolites are, of cour~e, cry~talline, ~hree-dim~nsional, stable ~truc~ure~ con~ainin~ a large number of uniorm opening~ or cavitie~ interconnected by smaller, rela-tively uniform holes or channel~. The effective pore . .
~20~ 2 size of synthetic zeolites is suitably between six and 15 A in di~meter. The overall ~ormula for the preferred zeolites can be represented as follows:
M2/n A123 1 5~5-5 Sio2 yH~O
where M is a metal cation and n its ~alence and x varies from 0 to 1 and y is a function of the degree of dehydra-tion and varies from 0 to 9. M is preferably a rare aarth metal cation such as lanthanum, cerium, praseo-dymium, neodymium or mixtures of these.
Zeolites which can be employed herein include both natural and syntheti~ ~eolites. These zeolites include gmelinite, chaba~ite, dachiardite, clinoptilo-lite, faujasite, heulandite, analcite, lev~nite, erio-nite, ~odalite, cancrinite, nepheline, lazurite, scolecite, natrolite, offretite, m~solita, mordenite, brewsterite, ferrierite, a~d the like. The faujasites are preferred. Suitable synthetic zeolites which can be treated in accordance with this invention include zeolites X, Y~ A, h, ZK-4, ~, EF, R, HJ, M, Q, T, W, Z, alpha and beta, Z5M-types and omega. The term "~.eolites" as used her2in contemplates not only alumino~
silicates but substances in whi~h the aluminum is replaced by gallium or boron and substances in which the silicon is replaced ~y germanium.
The preferred zeolites or thi~ invention are the synthetic faujasites of ~he kypes Y ~nd X or mix-tures thereof.
To obtain good cracking activiky the zeolit~s have to be in a proper form. In most cases this involves reducing the alkali metal content of the ~eolite to as low a level as possible. Further, a high alkali m~tal content r~duces the thermal struc~ural stability, and the effective lifetime of the catalyst will be impaired as a consequence thereof. Proc~dures ~%o~ ~
for removing alkali metals and putting the zeolite in the proper form are well known in the art as described in U.S. Pat. No. 3,537,816.
The crystalline aluminosilicate zeolites, such as synthetic faujasite, will under normal conditions crystallize as regularly shaped, discrete par~icles of approximately one to ten microns in size, and, accord-in~ly, this is the size range normally used in commer-cial catalysts. The particle size of the zeolites can be, for example, from about 0.5 to about 10 microns but generally from about 1 to about 2 microns or less.
Crystalline zeolites exhibit both an interior and an exterior surface area, with the largest portion of the total surface area being internal. Blockage of the internal channels by, for example, coke ormation and contamination by metals poisoning will greatly reduce the total surface area.
Especially preferred as the catalytically active component of the catalyst system used he~ein is a crystalline aluminosilicate, such as defined above, dispersed in a refractory metal oxide matrix, for e~ample, as set forth in U.S. Pat~ No. 3,944,482 to Mitchell et al., referred to hereinabove.
The matrix material in the host catalyst an be any well-known heat-stable or refractory metal compounds, for exampl~, metal oxidas, such as silica, magnesia, boron, zirconial or mixture~ of thes materials or suitable large pore clays, cross-linked clays or mixed oxide combinations.
The particular method of fsrming ~he catalyst matrix does not form a part of this inventiQn. Any method which produces the desired cracking activity characteristics can suit~bly be employed. Large pored refractory metal oxide material~ suitable for use as a matrix can be obtained as articles of commerce from catalyst manufacturers or they can be prepared in ways well known in the art ~uch a~ described, for example, in U.S. Pat. No. 2,890,162 The amount of the zeolitic material disper~ed in the matrix can suitably b~ from 3bout 10 to about 60 s weight pexcent, prefer~bly Irom about 10 to about 40 weight percent, but most preferably from about 20 to about 40 weight percent of the final catalyst. The method of forming the final composited catalyst also orms no part of this invention, and any method well known to those skilled in this art is acceptable. For example, inely divided zeolite can be admixed with the finely divided matrix material, and the mixture ~pray dried ~o form the final catalyst. Other suitable methods are described in U~S. Pat. Nos. 3,271,418;
3,717,587; 3,657,lS4; and 3,676,330~ Th~ zeolite can also be grown on ~he matrix material if desired, as deined, for example in U.S. Pat. No. 3,647,718 to Hayden et al., refer~ed to above.
The second component of the catalyst system used in ~he process defined herei~, as a ~eparate and distinct entity, is a diluent selected ro~ ~he group consisting of alumina and alumina irl combinakion wi~h at leas~ one heat-6table metal compound. By "combination"
we mean that the al~nina and the heat-6table ~etal 25 compo~md can be phy~ically discrete co~onents. By "heat-stable metal compoundl' we mean ~o include metal compound~ that: will, Imder the temperature6 exis'cing in a catalytic cracking unit, will no~ decompose, or if they do decompo~e, will decompose ~o compounds that will 30 remain stable in such environment. Exaraples of such heat-stable metal compound~ are the metal oxides of 6ilicon, magne~ m, boron, zircorliurQ, calcium and phosphoru6. The amoun~ o heat-stable c:ompou~d present carl be up ~o about 90 weight percen~ relative to the ,, :
, ~
:
alumina, although, i~ general, the amount will be up to about 50 weight percent or even higher.
The second component must be car~fully selected. In order to obtain the d~sired results herein, it is critical that i~s fresh surface area be in the range of about 30 to about 1000 m2/gram, preferably about 50 to about 600 m2/gram. E~ually cri~ical is the total pore volume, which must be in the range of about O . 05 to about 2 ~ 5 cc/gram, preferably about 0 .1 to about 1.5 cc/gram. It is desirable that the av~rage pore radius be in th~ range of about 10 to about 200 A, pxeferably about 20 to 110 A. The particle size can vary over a wide range, but generally will be in the range of about 20 to about 150 microns, preferably about 20 to about 90 microns.
The weight ratio of the catalytically active component to the diluent (the second component) can be in the range o about 10:90 to about 90:10, preferably in the rang~ o~ about S0:50 to about 70:30.
The catalyst composition defined above possesses a high tolerance to metals and thus is particularly useul in the cracking of high metals content charge stocks. Suitable charge stocks include crude oil, resi~uums or other petroleum fra~tions which are suitable catalytic cracking charge stocks except for the high metals content~. ~ high metals content charge stock for purposes of the process of ~his invention is d~fined as one having a total metals concentration equivalent to or greater than a value of ten as 30 calculated in accordance with the following relationshipo lO[Ni] ~ [V~ -t [F~] 2 10 where [Ni], [V] and [Fe] are the concentrations of nickel, vanadium and iron, respectively, in parts per million by weight. The process is particularly advan-tageou~ when the charge stock metals concentration is S equal ~o or greater than lO0 in the above equation. It is to be understood thexefore that the catalyst composi-tions describe~ above can be used in the catalytic cracking of any hydrocarbon charge stock containing metals, but is particularly useful for the treatment of high metals content charge stocks ~ince th~ useful life of the catalyst is increasedO The charge stocks can also be derived from coal, shale or tar sands. Thus charge stocks which have a metals content value of at least about 10 in accordance with the above equation cannot ~e treated as well as desired economically in commercial processes today due to high catalyst make-up xates, but can now be treated utilizing the catalyst compssitions described and claimed herein. Typical feedstocks are heavy ga~ oils or the heavier fractions of crude oil in which the metal contaminants are concan-trated. Particularly preexred charge stocks for treatmen~ by the process of this invention include deasphalted oils boiling above about 900F. ~482C.~ at atmospheric pressure; heavy gas oils boiling from about 650F. to ~bout 1100F. (343C. to 593C) at atmospheric pressure; atmospheric or vacuum tower bottoms boiling abo~e about 650F.
The preferred method o operating the process of this invention is by fluid catalytic cracking.
Hydrogen is generally not added to ~he reaction.
A sui~able reactor-regenerator for carryi~g ou~ the process claimed herein is shown in the attached Figure I. The cracking occurs in the presence of the fluidized novel catalys~ composition defined herein in an elongated reactor tube 10 which is referred to as a riser. The ri~er has a length to diameter ratio of ~æo~2 above 20 or above 25. The charge stock to be cracked is passed through preheater 2 to heat it to about 600F.
(315.6C.) and is then charged into the bottom of riser 10 to the end of line 14. Steam is introduced into oil inlet line ~4 through line 18. Steam is also intxoduced independently to the bottom of riser 10 -through line 22 to help carxy upwardly into the riser regenerated catalyst which ~lows to the bottom of the riser through transEer line 26.
The oil charge to be cracked in the riser is, for example, a heavy gas oil having a boiling range of about 650F. to abou-t llOO~F. (343 to 593C.). The steam added to the riser can amount to about 10 weight percent based on the oil charge, but the amount of steam can ~ary widPly. The ~atalyst employed is the novel catalyst composition defined above in a fluid form and is added to the bottom of the riser. The riser temperature range is suitably about 900F. to about 1100F. (482C. to 593C~) and is controlled by measuring the temperature of the product from the riser and then adjusting the opening of valve 40 by means of temperature contxoller 42 which regulates the inflow of hot regenera~ed catalyst to the bottom of riser 10. ~he temperature of the regenera~ed cakalys~ i~ above the control temperature in the riser so that the incoming catalyst contributes heat to the cracking reaction. The riser pressure is between ahout 10 and about 35 psig.
Between about 0 and about 5 percent of the oil charge to the risar can be recycled. The residence time of both ~, 30 hydrocarbon and catalyst in the riser i6 very ~mall and ranges from about 0.5 to about 5 seconds. The veloci~y through the riser i~ about 35 to about S5 eet per second and is sufficiently high so ~hat there is little or no slippage between the hydrocarbon and the catalyst flowing through the riser. Therefore no be~ of catalyst is permitted to build up within the riser whereby the density within the rise.r is very low. The density within the riser is a ma~imum of about 4 pounds per cubic foot at the bo~tom of the riser and decreases to a.bout 2 poun~s per cubic ~oot at the top o~ the riser.
Since no dense bed of catalyst is permitted ~o build up within the riser, the space velocity through the riser is unusually high and will have a range between about 100 or ab-)ut 120 and about 600 weight o~ hydrocarbon per hour per instantaneous weight of catalyst in the reacto.r. No signi~icant catalyst bu.ildup within the reactor is permi~ted to occur, and the instantaneous catalyst inventory wi~hln ~he riser is due to a flowing catalyst to oil weight ratio between about 4-1 and about 15:1, the weight ratio corresponding to the feed ratio.
The hydlrocarbon and catalyst exiting from the top o~ each rise;r is passed into a disengaging vessel 44. The top of the riser is capped at 46 so that discha~cge occurs through la~eral slots 50 for proper dispersion. An instantaneous separation between hydroc:arbon and c~at~alyst occurs in the disengaging vessel. The hydrocarbon which separat~s from the cata~yst is primarily gasoline together with some heavier components an~ some lighter gaseous components.
The hydrocarbon ef:fluen~ passes through cyclone system 5~ to separate ca~:alyst fines contained therein and is dis~harged to a ~ractionator through line 56. The catalyst separated from hydrocarbon in disengager 44 immediately drops ~)elow the outlets of the riser so that there is no catal~y~;t level in the disengager but only in a l.ower stripper ~;ection 5~. Steam is in~roduced into catalyst strippe:r section 58 through sparger 60 to remove any entrai~ d hydrocarbon in ~he cataly~t.
Catalys~ leaving stripper 58 passes through transfer line 62 to a regenerator 64. This catalyst contains carbon deposits which tend to lower its cracking ac~ivity ~nd as much car~on as possible must be - ~z~
ol2--burned from the surface of the catalyst. This burning is accomplished by introduction to the regenerator through line 66 of approximately the stoichiometrically re~uired ~mount of air for combuskion of the carbon deposits. The catalyst from the stripper enters the bottom section of the regenerator in a radial and downward direction through transfer line 62. Flue gas lea~ing the dense catalyst bed in regenerator 64 flows through cyclones 72 wherein catalyst fines are separated from flue gas permi~ting the flue gas to leave the regenerator through line 74 and pass through a turbine 76 before leaving for a waste heat boiler wherein any carbon monoxide contained in the flue gas is burned to carbon dioxide to accomplish heat recovery. Turbine 76 compresses atmospheric air in air compressor 78 and this air is charged to the bottom of the regenerator through line 66.
The temperature throughout the dense catalyst bed in the regenerator is about 1250F. (67~.7C.). The temperature of the flue gas leaving the top of the catalyst bed in the regenerator can rise due to afterburning o~ carbon monoxide to carbon dioxide.
Approximately a stoichiometric amount of oxygen is charged to the regenerator, and the r~a~on for this is to minimize afterburning of carbon monoxide to carbon dioxide above the catalyst bed to avoid injury to the e~uipment, since at the temperature of the regenerator flue gas some afterbuxning ~oeR occux. In order to prevent excessively high temperatures in the regenerator flue gas due to afterburni~g, the temperature of the regenerator flue gas is controlled by measuring the temperature of the flue gas entering the cyclones and then venting some of the pres6urized air otherwise destined to be charged to the bottom of the rege~erator 35 through vent line 80 in response to this measurement.
The regenerator reduces the carbon content of the ~Z0~12 catalyst from a~ou~ 0.5 weight percent to about 0.2 weight percent or less. If required, steam is available through line 82 for cooling the regenerator. Makeup catal~st is added to the bottom of the r~generator through line 84. Hopper 86 is disposed at the bottom of the regenerator for receiving regenerated catalyst to be passed to the bottom of the reactor riser through transfex line 26~
While in Figure I it has been shown that the novel catalyst composition herein can be introduced into the system as makeup by way of line 84, it is apparent that the catalyst composition, as makeup, or as fresh catalyst, in whole or in part, can be added to the system at any desirable or suitable point, for example, in line 26 or in line 14. Similarly, the components of the novel catalyst system need not be a~ded together but can be added separately at any of the respective points defined above. The amount added will vary, of course, depending upon the charge stock used, the catalytic cracking conditions in force, the conditions of regenexa-tion, the amount of metals p~esent in the catalyst under e~uilibrium conditionsl etc.
The reaction temperature in accordance with the above described process is at least about 900F.
(482C.)~ The upper limit can be about 1100F.
(593.3C.) or more. The preferred temperature r~nge is about 950F. to about 1050F. (510Co to 565.6C.). The reaction total pressure can vary widely and can be, for example, about 5 ~o about 50 psig (0.34 to 3.4 atmospheres), or preferably, about 20 to about 30 psig (1.36 to 2.04 atmospheres). The maximum residence time is about 5 seconds, and for most charg~ stocks ~he residence tim~ will be about 1.5 to about 2.5 seconds orr l~ss commonly, about 3 to about 4 seconds. For high molecular weight charge skocks, whicA are rich inaromatics, residence times of about 0.5 to about 1.5 seconds are suitable in order ~o crack mono- and di aromatics and naphthenes which ar~ the aromatics which crack most easily and which produce the highest gasoline yield, but to terminate the operation before appreciable cracking of polyaromatics occurs because these ma~erials produce high yields of coke and C~ and lighter gases.
The len~th to diameter ratio of the reactor can vary widely, but the reactor should be elongated to provide a high linear v~locity, such as about 25 to about 75 feet per second; and to this end a length to diameter ratio abov0 about 20 to about 25 is suitable. The reactor can have a uniform diameter or can be provided with a continuous ~aper or a stepwise increase in diameter along the reaction path to maintain a nearly constant velocity along the flow path. The amount of diluent can vary depending upon the xatio of hydrocarbon to diluent desired for control purpo~es. I steam is the diluent employed, a typical amount to be charged can be ~bout 10 percent by volume, which is about 1 percent ~y weight, based on hydrocarbon charge. A suitable but non-limiting proportion of diluent gas, such as steam or nitrogen, to fresh hydrocarbon feed can be about 0.5 to about 10 percent by weight~
The catalyst particle size (of each of the two components, that is, of the ca~alytically-ac~ive co~on-ent and of the diluent) must render it capable of fluidization as a disperse phase in the reactor.
Typical and non-limiting fluid catalyst particle size characteristics are as follows:
Size tMicrons) 0-20 20-45 45-75 > 75 Weight percent 0-5 20 30 35-55 20-40 These particle sizes are usual and are ns~ peculiar to thi~ invention. ~ suitable weight ratio of catalyst to total oil charge is about 4O1 to about 25:1, preferably , %
about 6:1 to about 10:1. The fresh hydrocarbon feed is generally preheated to a temperatuxe of abou~ ~00F. to a~out 700~F. (316C. to 371C.) but is generally not vaporized during preheat and the additional heat reguired to achieve ~he desired reactor temperature is imparted by hot, regenerated catalyst.
Th~ weight ratio of catalyst to hydrocarbon in the feed is varied to affect varia~ions in reactor temperature. Furthermore, ~he higher the tempera~ure o the regenerated catalyst the less catalyst is required to achieve a given rea~tion temperature~ Th~refore, a high regenerated catalyst temperatuxe will permit the very low reactor density level set forth below and thereby help to avoid back mixing in the reactor.
Generally catalyst regeneration can occur at an elevated temperature of ~bouk 1250Fo (676~6C~ ) or more to reduce the level of carbon on the regener~ted catalyst from about 0.6 to about 1.5, generally about 0.05 to 0.3 percent by weight. At usual catalyst to oil ratios in the feed, the quantity of catalyst is moxe than ample to achieve the desired catalytic effeGt and therefore if the temperature of the cat~lyst is high, the ratio can ~e ~afely decreased without i~pairing co~version. Since zeolitic catalysts, for example, are particularly sensitive to the carbon lev~l on the catalyst, regenera~
tion advantageously occurs at elevatad te~peratures in order to lower the carbon level on the ca~alyst to the stated range or loweL. Mareover, since a prime function of ~he ~atalyst is to contribute heat to the reactor, for any given desired reactor temperature ~he higher the temperature of ~he ca~alyst harge, the less catalyst is required. The lower the catalyst charge rate, the lower the densi~y of th~ material in the reactox. A stated, low reactor densities help to avoid backmixing.
Th reactQr linear velocity while not being so high that it induce~ turbulence and excessive hack-mixing, must be sufficiently high that substantially no catalyst accumulation or buildup occurs in the reactor because such accumulation itself leads to backmixing.
(Therefore, the catalys~ to oil weight ratio at any position throughout the reactor is about the same as the catalyst to oil weight ratio in the charge.) Stated another way, catalyst ancl hydrocarbon at any linear position along the reaction path both flow concurrently at about the same linear velocity, thereby avoiding significant slippage of catalyst rPlative to hydro carbon. A buildup of catalyst in the reactor leads to a dense bed and backmixing, which in turn increases the residence time in ~he reactor, for at least a portion of the charge hydrocarbon indu~es aftercrackiny. Avoiding a catalyst buildup in the reactor results in a very low catalyst inventory in the reactor, which in turn results in a high space v~locity. Therefore, a space velocity of over 100 to 120 weight o~ hydrocarbon per hour per weight of catalyst inven~ory is highly desirable. The space valocity should not be below about 35 and can be as high as about 500. ~ue to ~he low catalyst in~entory and low charge ratio of catalyst to hydrocarbon, the density of the material at the inlet of the reactor in the zone where the feed is charged can be only about 1 to less than 5 pounds per cubic foot, although these ranges are non-limiting. An inlet density in the zone where the low molecular weight feed and catalyst is charged below about 4 pounds per cu~ic foot is desirable since this density range is too low ~o encompass dense bed systems which induce backmixing. Although conver-sion ~alls off with a decrease in inlet density to very low levels, it has been found the ex~ent o after-cracking to be a more limiting feature than total conversion of fresh ~eed, even at an inlet density o 3S less than about 4 pounds per cubic oot. At the outlet of the reactor the dens:ity will be about hal o~ the density at the inlet because the cracking operation produces about a four-fold increase in mols of hydro-carbon. The decrease in density througll the reactor can be a measure of conversion.
The above conditions and description of operation are for the preferred fluid bed riser cracking operaticn. For cracking in the older conventional fluid bed operation or in a fixed-bed operatlon, the particu-lar reaction conditions are well known in the art.
A number of catalysts were evaluated for metals tolerance in accordance with the process claimed herein. Each was heat shocked at 1100F. (593C.) for one hour, contaminated with nickel and vanadium by impregnation with nickel and vanadium naphthenates, followed by calcination at 1000F. (538C.) for 10 hours and a steam treatment at 1350F. (732.3C) with about 100 percent steam for 10 hours. The average pore radii were determined after calcination, but before the steam treatment. Each of the catalysts carried on its surface 5000 ppm of nickel e~uivalents (3,800 parts per million of nickel and 6,000 parts per million of vanadium).
The "MAT Activit~:' was obtained by the use of the microactivity test previously described. The gas oil employed was described in Table I.
The catalysts used in the tests included GRZ-l alone and physical mixtures of GRZ-l and one of the following diluents:
Meta kaolin silica ACH-B (cross-linked clay) Aluminalite, ORE Alumina, Alum-Kaolin, AAA, Mg-AAP, and AAP, wherein the weight rativs of GRZ-l to diluent was 60:40.
Each of the above is defined furth~r below:
GRZ~ A commercial cracking catalyst contain-ing a high zeoli~e content composited with a refractory metal oxide matrix.
5 Meta kaolin - A clay predominating in silica and alumina in a 2:1 molar ratio, such as used in in U.S. Pat. No.
4,289,605.
ACH~B - A natural bentonite clay interlayered or ~ cross-linke~ with an aluminum hydroxy oligomer which ~erve as a proppant to expand the clay.
Aluminalite - An alumina hydxogel ormed by : co~gellation of the same polymeric alumina used in ACH~B, ~ogether with a ~uaternary ammonium slilica~e to give a silica-promotad material o~ uniform : ~ porosi~y.
' -lg .
ORE-~lumina - ~ product obtained from Ore Corporation, Cleveland, ~hio, which was washed to remove ~ontaminant chloride. It is believed to be alumina.
~lum~kaolin - The aluminalite defined above deposited on about 50 weight percent kaolin clay.
~AA - A commercial silica-alumina cracking catalyst (non-zeolitic) containing 25 weight percent alumi~a obtained from American Cyanamid, Wayne, New Jexsey.
M~AAP ~ A large-pore magnesium-promoted alumina-aluminum phosphate having the stoichoimetry 4MgO2 13A1~03 10AlP04.
(See U.S. Pat. No~ 4,179,358 to Swift et al.) ~AP - A large-pore, non-promoted alumina~
aluminum phosphate having the stoichio-metry A1203 2AlP04. (See U,S. Pat. No.
4,228,036 to Swift et al.) .
20 Silica - Davison Silica Gel, ~rade 59 obt~ined from Davison Company, Baltimore, Md.
The surface properties of each of the abo~e are set forth below in Table II:
TABLE II
Average Surface Pore Pore Area, Volume, Radius, Catalyst m2/g c~/g A
GRZ-l 222 0.17 16 Meta-Kaolin 10 0.04 80 Silica 249 1.12 90 ACH-B 250 0.16 20 ~luminalite 302 0.43 29 ORE Alumina 101 0.18 37 Alum-Kaolin 180 0.19 21 AAA 525 0.71 27 Mg-AAP .. 183 0.86 94 AAP 1~ 0o84 109 .
Th~ data obtained are tabulated below in Table 20 III:
.
:
:
, $~ ~
.
TABLE III
.... . . . . .
Con~er- C +
sion, 5 Vol. % ~Gasoline) Hydrogen, of Vol. % Carbon, Wt % of Run Fresh of Fresh Wt % of Fresh No. Catalyst Feed Feed Catalyst Feed 1 GRZ-l 60 37.6 5.2 0.58
PROCESS FOR CRACKING
HIGH METALS CONTENT FEEDSTOCKS
This invention relates to a process for cracking high metals content feedstocks which comprises contacting said charge stock under catalytic cracking conditions with a novel catalyst composition comprising a solid cracking catalyst and a diluent containing a selected alumina or a selected alumina in combinatio~
with one or more heat-stable metal compounds.
U.S. Pat. No~ 3,944,4a2 to Mitchell et al.
~ discloses a process directed to the catalytic cracking ;of hydrocarbon feeds containing m tals using a~ fluid c~atalyst having improved metals~ tolerant character-istics. Bartholic in U.S. Pat.~No. 4/289,60~5 xelates to a process for the catalytic cracking of`~hydrocarbon feeds containing metals using a catalyst composition containing a solid cracking~ catalyst and calcined microspheres (for example~, calcined;~kaolin clay) havin~
a surface area within the; range~ of ~10 to 15 m2/gram.
We have found that catalytic cracking of high metals content feedstocks 6uch as, for example, those containing iron, vanadium, nic~el and copper, can be S substantially improved by contacting said charge stocks under catalytic cracking conditions wi~h a novel catalyst composition compri6ing a solid cracking catalyst a~d a diluent selected from the group co~sist-ing of alumina (~1203) and alumina in co~bina~ion with one or more heat-stable metal compounds. The improve-ment resides in ~he ability of the cataly~t system to unction well even when the catalyst carries a substan-tially high level of metal on i~s surface, for example, up to 5000 ppm of ni~kel or nickel equivale~ts, or eYen higher, ox up t~ 20,000 ppm o vanadium. By "ppm of nic~el eguivalent" we mean p~m ~ick 1 ~ O.20 ppm vanadium. Thus feeds~ock~ having very high metals content can be sati~f~ctorily u~ed herein.
Thus accordi~g to the pre~ent invention there is provided a process for the catalytic cracki~g o~ a high me~als content charge stock which comprises contacting said charge stock under ca~alytic cr cking conditions with a catalyst ccmposition comprising a cracking catalyst having high activity and a diluent selected f~om ~he group co~sisting of alumina and alumina in combina ion with a heat-stable metal ~ompound, said diluen~ having a surface area of ,about 30 ~o abou~ lOOQ m~/gr~ and a pore volume of about 0.05 to about 2.5 cc/gram.
.
,, '.
-- ~%~
- 2a -The cracking catalyst component of the novel catalys-t composition used in -the novel process herein can be any cracking catalyst of any desired -type having high activity. By ~high activity" we mean catalyst of fresh MAT Activity above about 1.0, preferably up to a~out 4.0, or even higher, where Activity = Wt % Conversion 100 Wt % Converslon The "MAT Activity" was obtained by the use of a microactivi-ty test (MAT) unit similar to the standard Davison MAT (see Ciapetta et al., Oil & Gas Journal, 65, 88 (1967). A11 catalyst samples were tested at 900F.
(482C.~ reaction temperature; 15 weight hourly space velocity; 80 seconds of catalyst contact time; and a catalyst to oil ratio o~ 2.9 with 2O5 grams o~ catalyst.
The charge stock was a Kuwait gas oil havin~ a boiling range of 500F. to 800F. (260C. to 427C.). Inspec-tions o this Kuwait gas oil are shown i~ Table I below.
TABLE I
.
KUW~IT GAS OIL INSPECTIONS
_ Stock MAT
Identification Feedstock Inspections:
Gravity, API, D-287 23.5 Viscosity, SUS D2161, 130F. 94.7 Viscosity, SUS D2161, 150F. 70.5 Viscosi~y, SUS ~2161, 210~. 50.8 Pour Point, D97, F. ~80 Nitrogen, wt % 0.Q74 Sulfur, wt % 2.76 Carbon, Res., D524, wt % 0.23 Bromine No., D1159 5.71 Aniline Point, F. 176.5 Nickel, ppm <0.1 Vanadium, ppm <0.1 Distillation, ~1160 at 760 mm End Point, F. aoo 5 Pct. Cond. 505 Approx. Hydrocarbon Type AnaIysis: Vol. %
Carbon as Aromatics 23.1 Carhon as Naphthenes 10.5 Carbon as Paraffins 66.3 ~ _ Thus, catalytic ~racking catalysts suitable for use herein as host catalyst include amorphous silica~alumina cataly6t6; 6yn~hetic mica-montmorillonite catalysts as defined, for example in U.S. Pat. No.
3~252,889 to Capell et al.; and cross-linked clays (see, for example, Vaughn et al. in U.S~ Pat. Nos.. 4,176,090 and 4,248,739; Vaug~n et al. ~1980~, "Prepar~tion of Molecular Sieves Based on ~illared Interlayered Clays";
Proceedings of the 5th _ternational Conference on 2eolites, Rees, L.V., Editor, Heyden, London, pages 94-101; and Lahav et al., (1978) "Crosslinked Smectites I Synthesis and Properties of ~Iydroxy Aluminum Montmox-illonite'l, ClaY ~ Clay Minerals, 26, p`ages 107-114;
Shabtai, ~. in U.S. PatO No. ~,238,364; and 5habria et al. in U.S. Pat. No. 4,216,188~.
Preferably, ~h~ host catalyst used herein is 15 a catalyst containing a crystalline alwninosilica~e, preferably exchanged wi th rare earth metal cations, sometimes referred to as "raxe earth-exs:hang~d crystal-line. aluminum silicate" or one of the stabilized hydroyen ~eolites. ~os~ preferably, the host ~atalyst 20 i8 a high activit~ cracking cataly~t.
Typical zeolites or molecular siev2~ having cracking activity which can be used herei~ a~ a cata-lytic crac~cing ca~aly~t are well known in ~e art.
Suitable zeoli~es are de~cribed, for exampleO in U. S .
Pat~ No. 3,660,274 ~o Blazek et al., or in U.S. Pat. No.
3, 647, 718 to ~ayden et al .
Syn~hetically prepared zeolite~ are i~itially in the Iorm of alkali metal aIumino6ilicate~. The alkali ~etal ions are exchanged with rare earth metal ions to impart cracking charac~eristics to the zeolites. The zeolites are, of cour~e, cry~talline, ~hree-dim~nsional, stable ~truc~ure~ con~ainin~ a large number of uniorm opening~ or cavitie~ interconnected by smaller, rela-tively uniform holes or channel~. The effective pore . .
~20~ 2 size of synthetic zeolites is suitably between six and 15 A in di~meter. The overall ~ormula for the preferred zeolites can be represented as follows:
M2/n A123 1 5~5-5 Sio2 yH~O
where M is a metal cation and n its ~alence and x varies from 0 to 1 and y is a function of the degree of dehydra-tion and varies from 0 to 9. M is preferably a rare aarth metal cation such as lanthanum, cerium, praseo-dymium, neodymium or mixtures of these.
Zeolites which can be employed herein include both natural and syntheti~ ~eolites. These zeolites include gmelinite, chaba~ite, dachiardite, clinoptilo-lite, faujasite, heulandite, analcite, lev~nite, erio-nite, ~odalite, cancrinite, nepheline, lazurite, scolecite, natrolite, offretite, m~solita, mordenite, brewsterite, ferrierite, a~d the like. The faujasites are preferred. Suitable synthetic zeolites which can be treated in accordance with this invention include zeolites X, Y~ A, h, ZK-4, ~, EF, R, HJ, M, Q, T, W, Z, alpha and beta, Z5M-types and omega. The term "~.eolites" as used her2in contemplates not only alumino~
silicates but substances in whi~h the aluminum is replaced by gallium or boron and substances in which the silicon is replaced ~y germanium.
The preferred zeolites or thi~ invention are the synthetic faujasites of ~he kypes Y ~nd X or mix-tures thereof.
To obtain good cracking activiky the zeolit~s have to be in a proper form. In most cases this involves reducing the alkali metal content of the ~eolite to as low a level as possible. Further, a high alkali m~tal content r~duces the thermal struc~ural stability, and the effective lifetime of the catalyst will be impaired as a consequence thereof. Proc~dures ~%o~ ~
for removing alkali metals and putting the zeolite in the proper form are well known in the art as described in U.S. Pat. No. 3,537,816.
The crystalline aluminosilicate zeolites, such as synthetic faujasite, will under normal conditions crystallize as regularly shaped, discrete par~icles of approximately one to ten microns in size, and, accord-in~ly, this is the size range normally used in commer-cial catalysts. The particle size of the zeolites can be, for example, from about 0.5 to about 10 microns but generally from about 1 to about 2 microns or less.
Crystalline zeolites exhibit both an interior and an exterior surface area, with the largest portion of the total surface area being internal. Blockage of the internal channels by, for example, coke ormation and contamination by metals poisoning will greatly reduce the total surface area.
Especially preferred as the catalytically active component of the catalyst system used he~ein is a crystalline aluminosilicate, such as defined above, dispersed in a refractory metal oxide matrix, for e~ample, as set forth in U.S. Pat~ No. 3,944,482 to Mitchell et al., referred to hereinabove.
The matrix material in the host catalyst an be any well-known heat-stable or refractory metal compounds, for exampl~, metal oxidas, such as silica, magnesia, boron, zirconial or mixture~ of thes materials or suitable large pore clays, cross-linked clays or mixed oxide combinations.
The particular method of fsrming ~he catalyst matrix does not form a part of this inventiQn. Any method which produces the desired cracking activity characteristics can suit~bly be employed. Large pored refractory metal oxide material~ suitable for use as a matrix can be obtained as articles of commerce from catalyst manufacturers or they can be prepared in ways well known in the art ~uch a~ described, for example, in U.S. Pat. No. 2,890,162 The amount of the zeolitic material disper~ed in the matrix can suitably b~ from 3bout 10 to about 60 s weight pexcent, prefer~bly Irom about 10 to about 40 weight percent, but most preferably from about 20 to about 40 weight percent of the final catalyst. The method of forming the final composited catalyst also orms no part of this invention, and any method well known to those skilled in this art is acceptable. For example, inely divided zeolite can be admixed with the finely divided matrix material, and the mixture ~pray dried ~o form the final catalyst. Other suitable methods are described in U~S. Pat. Nos. 3,271,418;
3,717,587; 3,657,lS4; and 3,676,330~ Th~ zeolite can also be grown on ~he matrix material if desired, as deined, for example in U.S. Pat. No. 3,647,718 to Hayden et al., refer~ed to above.
The second component of the catalyst system used in ~he process defined herei~, as a ~eparate and distinct entity, is a diluent selected ro~ ~he group consisting of alumina and alumina irl combinakion wi~h at leas~ one heat-6table metal compound. By "combination"
we mean that the al~nina and the heat-6table ~etal 25 compo~md can be phy~ically discrete co~onents. By "heat-stable metal compoundl' we mean ~o include metal compound~ that: will, Imder the temperature6 exis'cing in a catalytic cracking unit, will no~ decompose, or if they do decompo~e, will decompose ~o compounds that will 30 remain stable in such environment. Exaraples of such heat-stable metal compound~ are the metal oxides of 6ilicon, magne~ m, boron, zircorliurQ, calcium and phosphoru6. The amoun~ o heat-stable c:ompou~d present carl be up ~o about 90 weight percen~ relative to the ,, :
, ~
:
alumina, although, i~ general, the amount will be up to about 50 weight percent or even higher.
The second component must be car~fully selected. In order to obtain the d~sired results herein, it is critical that i~s fresh surface area be in the range of about 30 to about 1000 m2/gram, preferably about 50 to about 600 m2/gram. E~ually cri~ical is the total pore volume, which must be in the range of about O . 05 to about 2 ~ 5 cc/gram, preferably about 0 .1 to about 1.5 cc/gram. It is desirable that the av~rage pore radius be in th~ range of about 10 to about 200 A, pxeferably about 20 to 110 A. The particle size can vary over a wide range, but generally will be in the range of about 20 to about 150 microns, preferably about 20 to about 90 microns.
The weight ratio of the catalytically active component to the diluent (the second component) can be in the range o about 10:90 to about 90:10, preferably in the rang~ o~ about S0:50 to about 70:30.
The catalyst composition defined above possesses a high tolerance to metals and thus is particularly useul in the cracking of high metals content charge stocks. Suitable charge stocks include crude oil, resi~uums or other petroleum fra~tions which are suitable catalytic cracking charge stocks except for the high metals content~. ~ high metals content charge stock for purposes of the process of ~his invention is d~fined as one having a total metals concentration equivalent to or greater than a value of ten as 30 calculated in accordance with the following relationshipo lO[Ni] ~ [V~ -t [F~] 2 10 where [Ni], [V] and [Fe] are the concentrations of nickel, vanadium and iron, respectively, in parts per million by weight. The process is particularly advan-tageou~ when the charge stock metals concentration is S equal ~o or greater than lO0 in the above equation. It is to be understood thexefore that the catalyst composi-tions describe~ above can be used in the catalytic cracking of any hydrocarbon charge stock containing metals, but is particularly useful for the treatment of high metals content charge stocks ~ince th~ useful life of the catalyst is increasedO The charge stocks can also be derived from coal, shale or tar sands. Thus charge stocks which have a metals content value of at least about 10 in accordance with the above equation cannot ~e treated as well as desired economically in commercial processes today due to high catalyst make-up xates, but can now be treated utilizing the catalyst compssitions described and claimed herein. Typical feedstocks are heavy ga~ oils or the heavier fractions of crude oil in which the metal contaminants are concan-trated. Particularly preexred charge stocks for treatmen~ by the process of this invention include deasphalted oils boiling above about 900F. ~482C.~ at atmospheric pressure; heavy gas oils boiling from about 650F. to ~bout 1100F. (343C. to 593C) at atmospheric pressure; atmospheric or vacuum tower bottoms boiling abo~e about 650F.
The preferred method o operating the process of this invention is by fluid catalytic cracking.
Hydrogen is generally not added to ~he reaction.
A sui~able reactor-regenerator for carryi~g ou~ the process claimed herein is shown in the attached Figure I. The cracking occurs in the presence of the fluidized novel catalys~ composition defined herein in an elongated reactor tube 10 which is referred to as a riser. The ri~er has a length to diameter ratio of ~æo~2 above 20 or above 25. The charge stock to be cracked is passed through preheater 2 to heat it to about 600F.
(315.6C.) and is then charged into the bottom of riser 10 to the end of line 14. Steam is introduced into oil inlet line ~4 through line 18. Steam is also intxoduced independently to the bottom of riser 10 -through line 22 to help carxy upwardly into the riser regenerated catalyst which ~lows to the bottom of the riser through transEer line 26.
The oil charge to be cracked in the riser is, for example, a heavy gas oil having a boiling range of about 650F. to abou-t llOO~F. (343 to 593C.). The steam added to the riser can amount to about 10 weight percent based on the oil charge, but the amount of steam can ~ary widPly. The ~atalyst employed is the novel catalyst composition defined above in a fluid form and is added to the bottom of the riser. The riser temperature range is suitably about 900F. to about 1100F. (482C. to 593C~) and is controlled by measuring the temperature of the product from the riser and then adjusting the opening of valve 40 by means of temperature contxoller 42 which regulates the inflow of hot regenera~ed catalyst to the bottom of riser 10. ~he temperature of the regenera~ed cakalys~ i~ above the control temperature in the riser so that the incoming catalyst contributes heat to the cracking reaction. The riser pressure is between ahout 10 and about 35 psig.
Between about 0 and about 5 percent of the oil charge to the risar can be recycled. The residence time of both ~, 30 hydrocarbon and catalyst in the riser i6 very ~mall and ranges from about 0.5 to about 5 seconds. The veloci~y through the riser i~ about 35 to about S5 eet per second and is sufficiently high so ~hat there is little or no slippage between the hydrocarbon and the catalyst flowing through the riser. Therefore no be~ of catalyst is permitted to build up within the riser whereby the density within the rise.r is very low. The density within the riser is a ma~imum of about 4 pounds per cubic foot at the bo~tom of the riser and decreases to a.bout 2 poun~s per cubic ~oot at the top o~ the riser.
Since no dense bed of catalyst is permitted ~o build up within the riser, the space velocity through the riser is unusually high and will have a range between about 100 or ab-)ut 120 and about 600 weight o~ hydrocarbon per hour per instantaneous weight of catalyst in the reacto.r. No signi~icant catalyst bu.ildup within the reactor is permi~ted to occur, and the instantaneous catalyst inventory wi~hln ~he riser is due to a flowing catalyst to oil weight ratio between about 4-1 and about 15:1, the weight ratio corresponding to the feed ratio.
The hydlrocarbon and catalyst exiting from the top o~ each rise;r is passed into a disengaging vessel 44. The top of the riser is capped at 46 so that discha~cge occurs through la~eral slots 50 for proper dispersion. An instantaneous separation between hydroc:arbon and c~at~alyst occurs in the disengaging vessel. The hydrocarbon which separat~s from the cata~yst is primarily gasoline together with some heavier components an~ some lighter gaseous components.
The hydrocarbon ef:fluen~ passes through cyclone system 5~ to separate ca~:alyst fines contained therein and is dis~harged to a ~ractionator through line 56. The catalyst separated from hydrocarbon in disengager 44 immediately drops ~)elow the outlets of the riser so that there is no catal~y~;t level in the disengager but only in a l.ower stripper ~;ection 5~. Steam is in~roduced into catalyst strippe:r section 58 through sparger 60 to remove any entrai~ d hydrocarbon in ~he cataly~t.
Catalys~ leaving stripper 58 passes through transfer line 62 to a regenerator 64. This catalyst contains carbon deposits which tend to lower its cracking ac~ivity ~nd as much car~on as possible must be - ~z~
ol2--burned from the surface of the catalyst. This burning is accomplished by introduction to the regenerator through line 66 of approximately the stoichiometrically re~uired ~mount of air for combuskion of the carbon deposits. The catalyst from the stripper enters the bottom section of the regenerator in a radial and downward direction through transfer line 62. Flue gas lea~ing the dense catalyst bed in regenerator 64 flows through cyclones 72 wherein catalyst fines are separated from flue gas permi~ting the flue gas to leave the regenerator through line 74 and pass through a turbine 76 before leaving for a waste heat boiler wherein any carbon monoxide contained in the flue gas is burned to carbon dioxide to accomplish heat recovery. Turbine 76 compresses atmospheric air in air compressor 78 and this air is charged to the bottom of the regenerator through line 66.
The temperature throughout the dense catalyst bed in the regenerator is about 1250F. (67~.7C.). The temperature of the flue gas leaving the top of the catalyst bed in the regenerator can rise due to afterburning o~ carbon monoxide to carbon dioxide.
Approximately a stoichiometric amount of oxygen is charged to the regenerator, and the r~a~on for this is to minimize afterburning of carbon monoxide to carbon dioxide above the catalyst bed to avoid injury to the e~uipment, since at the temperature of the regenerator flue gas some afterbuxning ~oeR occux. In order to prevent excessively high temperatures in the regenerator flue gas due to afterburni~g, the temperature of the regenerator flue gas is controlled by measuring the temperature of the flue gas entering the cyclones and then venting some of the pres6urized air otherwise destined to be charged to the bottom of the rege~erator 35 through vent line 80 in response to this measurement.
The regenerator reduces the carbon content of the ~Z0~12 catalyst from a~ou~ 0.5 weight percent to about 0.2 weight percent or less. If required, steam is available through line 82 for cooling the regenerator. Makeup catal~st is added to the bottom of the r~generator through line 84. Hopper 86 is disposed at the bottom of the regenerator for receiving regenerated catalyst to be passed to the bottom of the reactor riser through transfex line 26~
While in Figure I it has been shown that the novel catalyst composition herein can be introduced into the system as makeup by way of line 84, it is apparent that the catalyst composition, as makeup, or as fresh catalyst, in whole or in part, can be added to the system at any desirable or suitable point, for example, in line 26 or in line 14. Similarly, the components of the novel catalyst system need not be a~ded together but can be added separately at any of the respective points defined above. The amount added will vary, of course, depending upon the charge stock used, the catalytic cracking conditions in force, the conditions of regenexa-tion, the amount of metals p~esent in the catalyst under e~uilibrium conditionsl etc.
The reaction temperature in accordance with the above described process is at least about 900F.
(482C.)~ The upper limit can be about 1100F.
(593.3C.) or more. The preferred temperature r~nge is about 950F. to about 1050F. (510Co to 565.6C.). The reaction total pressure can vary widely and can be, for example, about 5 ~o about 50 psig (0.34 to 3.4 atmospheres), or preferably, about 20 to about 30 psig (1.36 to 2.04 atmospheres). The maximum residence time is about 5 seconds, and for most charg~ stocks ~he residence tim~ will be about 1.5 to about 2.5 seconds orr l~ss commonly, about 3 to about 4 seconds. For high molecular weight charge skocks, whicA are rich inaromatics, residence times of about 0.5 to about 1.5 seconds are suitable in order ~o crack mono- and di aromatics and naphthenes which ar~ the aromatics which crack most easily and which produce the highest gasoline yield, but to terminate the operation before appreciable cracking of polyaromatics occurs because these ma~erials produce high yields of coke and C~ and lighter gases.
The len~th to diameter ratio of the reactor can vary widely, but the reactor should be elongated to provide a high linear v~locity, such as about 25 to about 75 feet per second; and to this end a length to diameter ratio abov0 about 20 to about 25 is suitable. The reactor can have a uniform diameter or can be provided with a continuous ~aper or a stepwise increase in diameter along the reaction path to maintain a nearly constant velocity along the flow path. The amount of diluent can vary depending upon the xatio of hydrocarbon to diluent desired for control purpo~es. I steam is the diluent employed, a typical amount to be charged can be ~bout 10 percent by volume, which is about 1 percent ~y weight, based on hydrocarbon charge. A suitable but non-limiting proportion of diluent gas, such as steam or nitrogen, to fresh hydrocarbon feed can be about 0.5 to about 10 percent by weight~
The catalyst particle size (of each of the two components, that is, of the ca~alytically-ac~ive co~on-ent and of the diluent) must render it capable of fluidization as a disperse phase in the reactor.
Typical and non-limiting fluid catalyst particle size characteristics are as follows:
Size tMicrons) 0-20 20-45 45-75 > 75 Weight percent 0-5 20 30 35-55 20-40 These particle sizes are usual and are ns~ peculiar to thi~ invention. ~ suitable weight ratio of catalyst to total oil charge is about 4O1 to about 25:1, preferably , %
about 6:1 to about 10:1. The fresh hydrocarbon feed is generally preheated to a temperatuxe of abou~ ~00F. to a~out 700~F. (316C. to 371C.) but is generally not vaporized during preheat and the additional heat reguired to achieve ~he desired reactor temperature is imparted by hot, regenerated catalyst.
Th~ weight ratio of catalyst to hydrocarbon in the feed is varied to affect varia~ions in reactor temperature. Furthermore, ~he higher the tempera~ure o the regenerated catalyst the less catalyst is required to achieve a given rea~tion temperature~ Th~refore, a high regenerated catalyst temperatuxe will permit the very low reactor density level set forth below and thereby help to avoid back mixing in the reactor.
Generally catalyst regeneration can occur at an elevated temperature of ~bouk 1250Fo (676~6C~ ) or more to reduce the level of carbon on the regener~ted catalyst from about 0.6 to about 1.5, generally about 0.05 to 0.3 percent by weight. At usual catalyst to oil ratios in the feed, the quantity of catalyst is moxe than ample to achieve the desired catalytic effeGt and therefore if the temperature of the cat~lyst is high, the ratio can ~e ~afely decreased without i~pairing co~version. Since zeolitic catalysts, for example, are particularly sensitive to the carbon lev~l on the catalyst, regenera~
tion advantageously occurs at elevatad te~peratures in order to lower the carbon level on the ca~alyst to the stated range or loweL. Mareover, since a prime function of ~he ~atalyst is to contribute heat to the reactor, for any given desired reactor temperature ~he higher the temperature of ~he ca~alyst harge, the less catalyst is required. The lower the catalyst charge rate, the lower the densi~y of th~ material in the reactox. A stated, low reactor densities help to avoid backmixing.
Th reactQr linear velocity while not being so high that it induce~ turbulence and excessive hack-mixing, must be sufficiently high that substantially no catalyst accumulation or buildup occurs in the reactor because such accumulation itself leads to backmixing.
(Therefore, the catalys~ to oil weight ratio at any position throughout the reactor is about the same as the catalyst to oil weight ratio in the charge.) Stated another way, catalyst ancl hydrocarbon at any linear position along the reaction path both flow concurrently at about the same linear velocity, thereby avoiding significant slippage of catalyst rPlative to hydro carbon. A buildup of catalyst in the reactor leads to a dense bed and backmixing, which in turn increases the residence time in ~he reactor, for at least a portion of the charge hydrocarbon indu~es aftercrackiny. Avoiding a catalyst buildup in the reactor results in a very low catalyst inventory in the reactor, which in turn results in a high space v~locity. Therefore, a space velocity of over 100 to 120 weight o~ hydrocarbon per hour per weight of catalyst inven~ory is highly desirable. The space valocity should not be below about 35 and can be as high as about 500. ~ue to ~he low catalyst in~entory and low charge ratio of catalyst to hydrocarbon, the density of the material at the inlet of the reactor in the zone where the feed is charged can be only about 1 to less than 5 pounds per cubic foot, although these ranges are non-limiting. An inlet density in the zone where the low molecular weight feed and catalyst is charged below about 4 pounds per cu~ic foot is desirable since this density range is too low ~o encompass dense bed systems which induce backmixing. Although conver-sion ~alls off with a decrease in inlet density to very low levels, it has been found the ex~ent o after-cracking to be a more limiting feature than total conversion of fresh ~eed, even at an inlet density o 3S less than about 4 pounds per cubic oot. At the outlet of the reactor the dens:ity will be about hal o~ the density at the inlet because the cracking operation produces about a four-fold increase in mols of hydro-carbon. The decrease in density througll the reactor can be a measure of conversion.
The above conditions and description of operation are for the preferred fluid bed riser cracking operaticn. For cracking in the older conventional fluid bed operation or in a fixed-bed operatlon, the particu-lar reaction conditions are well known in the art.
A number of catalysts were evaluated for metals tolerance in accordance with the process claimed herein. Each was heat shocked at 1100F. (593C.) for one hour, contaminated with nickel and vanadium by impregnation with nickel and vanadium naphthenates, followed by calcination at 1000F. (538C.) for 10 hours and a steam treatment at 1350F. (732.3C) with about 100 percent steam for 10 hours. The average pore radii were determined after calcination, but before the steam treatment. Each of the catalysts carried on its surface 5000 ppm of nickel e~uivalents (3,800 parts per million of nickel and 6,000 parts per million of vanadium).
The "MAT Activit~:' was obtained by the use of the microactivity test previously described. The gas oil employed was described in Table I.
The catalysts used in the tests included GRZ-l alone and physical mixtures of GRZ-l and one of the following diluents:
Meta kaolin silica ACH-B (cross-linked clay) Aluminalite, ORE Alumina, Alum-Kaolin, AAA, Mg-AAP, and AAP, wherein the weight rativs of GRZ-l to diluent was 60:40.
Each of the above is defined furth~r below:
GRZ~ A commercial cracking catalyst contain-ing a high zeoli~e content composited with a refractory metal oxide matrix.
5 Meta kaolin - A clay predominating in silica and alumina in a 2:1 molar ratio, such as used in in U.S. Pat. No.
4,289,605.
ACH~B - A natural bentonite clay interlayered or ~ cross-linke~ with an aluminum hydroxy oligomer which ~erve as a proppant to expand the clay.
Aluminalite - An alumina hydxogel ormed by : co~gellation of the same polymeric alumina used in ACH~B, ~ogether with a ~uaternary ammonium slilica~e to give a silica-promotad material o~ uniform : ~ porosi~y.
' -lg .
ORE-~lumina - ~ product obtained from Ore Corporation, Cleveland, ~hio, which was washed to remove ~ontaminant chloride. It is believed to be alumina.
~lum~kaolin - The aluminalite defined above deposited on about 50 weight percent kaolin clay.
~AA - A commercial silica-alumina cracking catalyst (non-zeolitic) containing 25 weight percent alumi~a obtained from American Cyanamid, Wayne, New Jexsey.
M~AAP ~ A large-pore magnesium-promoted alumina-aluminum phosphate having the stoichoimetry 4MgO2 13A1~03 10AlP04.
(See U.S. Pat. No~ 4,179,358 to Swift et al.) ~AP - A large-pore, non-promoted alumina~
aluminum phosphate having the stoichio-metry A1203 2AlP04. (See U,S. Pat. No.
4,228,036 to Swift et al.) .
20 Silica - Davison Silica Gel, ~rade 59 obt~ined from Davison Company, Baltimore, Md.
The surface properties of each of the abo~e are set forth below in Table II:
TABLE II
Average Surface Pore Pore Area, Volume, Radius, Catalyst m2/g c~/g A
GRZ-l 222 0.17 16 Meta-Kaolin 10 0.04 80 Silica 249 1.12 90 ACH-B 250 0.16 20 ~luminalite 302 0.43 29 ORE Alumina 101 0.18 37 Alum-Kaolin 180 0.19 21 AAA 525 0.71 27 Mg-AAP .. 183 0.86 94 AAP 1~ 0o84 109 .
Th~ data obtained are tabulated below in Table 20 III:
.
:
:
, $~ ~
.
TABLE III
.... . . . . .
Con~er- C +
sion, 5 Vol. % ~Gasoline) Hydrogen, of Vol. % Carbon, Wt % of Run Fresh of Fresh Wt % of Fresh No. Catalyst Feed Feed Catalyst Feed 1 GRZ-l 60 37.6 5.2 0.58
2 Meta-Kaolin~ 49.4 33.8 3.1 0.34
3 Silica* 51.0 32.1 4.6 0.27
4 ACH-B* 54.4 36.7 3.9 0.43 Aluminalite* 64.9 37.8 6.8 0.78 6 ORE Alumina* 64.6 43.4 4.3 0.25 7 Alum-Kaolin* 65.6 37.3 8.2 0.77 8 AAA~ . 65.3 38.3 8.1 0.64 9 Mg-AAP* 58.6 35.1 6.6 0.57 AAP* 620 2 37.3 6.~ 0.65 _ _~____ (~ GRZ-l diluted with indicated additive. Resultant catalyst contained GRZ-l and diluent ln a we~ght ratio o~ 60:40.
All catalysts contaminated with 5000 parts per million of nickel equivalents.
The unusual results obtained by operation in accordance with the process defined herein are seen from the data in Table:III. Thus, in Run No. 1, wherein the process wa~ operated with a commercially available high activity catalyst, which has excellent metals tolerant characteris~ics when used in catalyti crack-ing of hydrocarbonaceQus feeds, excellent results were obtained, even with the catalyst carrying 5000 ppm nickel e~uivalents. When in Run No. 2, . the zeolite catalyst of Run No. 1 was diluted with meta-kaolin in a wei~ht ratio of 60:40, following `:
~$.~
the teachings of U.S. Pat. Nc. 4,289,605 of Bartholic, inferior results were obtained compared with those obtained in ~n No. 1, in tha~ conversion was reduced to 49.43 percent, with a drop in gasoline production.
Using silica in place of meta-kaolin in Run No. 3 similarly produced inferior results. However, when the zeolitic catalyst was combined with a h~at-stable refractory aluminum oxide alone, as in Run No. 6 or wi.th aluminum oxide in combination with other heat-stable metal compounds in ~uns Nos. 4, 5 and 7 to 10, con~er-sions and amounts of gasoline were almost as good or even better than the results obtained in Run No. 1.
This is surpri~ing, in that the diluents used in Runs Nos. 4 to 10 do not contain zeolite, and yet when a portion of the ca~alytically active componen~ was replaced with such diluent, excellen~ results were ~till obtained.
An additional series of run~ was carried out similarly to Runs NOB. 1 to 10 above wherein catalyst composition mixtures were e~ployed containing 60 weight p~xcent of GRZ-l and 40 wei~ht percent of ORE-Alumina.
In Run No. 11 the catalyst composition was free of any metal Gontent~ while in Runs Nos~ 1~, 13, 14 and 15 the catalyst compositions used carried 0.5/ 1.0, 1.5 and 2.0 25 weighk percent vanadium, respectively, on th~ir sur-faces. Vanadium was deposited on the catalyst composi~
tion surfaces using vanadi~m naphthe~ate following the procedure of Run~ Nos. 1 to 10. The data o~tained are tabulated below in Table IV.
TABLE IV
Conver- C ~
sion, 5 Hydrogen, Vol. % (Gasol;ne) Carbon, Wt % of Run Vanadium, of Fresh Vol. % of Wt % of Fresh No. Wt %Feed Fresh Feed Catalyst Feed ll 0 72.0 50.0 ~.2 0.03 l2 0.56~.5 48.9 4.0 O.l9 l3 l.0~7.0 44.0 5.3 0.37 l~ l.56~.0 40.0 4.6 0.42 2.054.0 33.0 4.5 0.41 ... .... ~
Still additional run~ were carried out as above but wherein GRZ-l alone was used free of any metal content and with vanadium presen~ in amounts o l.0/ 1.5 and 2.0 weight percents. The data obtained are tabu-lated below in Table V.
TABLE V
_ , .. . . .
Conver- C ~
sion, 5 Hydrogen, Vol. % ~Gasoline) Carbon, Wt % of Run Vanadium, oP Fresh YolO % ofWt % oP Fresh No. Wt % Feed - Fresh Feed Catalyst Feed 16 0 78.0 54.0 5.8 0.02 l7 l.0 57.0 40.0 2.5 0.20 18 1.5 41.0 31.0 1.8 0.13 l9 2.0 2B.O 13.0 1.4 O.ll ;
The advantages o~ operating the defined ~rocess using the novel catalyst herein are further apparent from the data in Tahles IV and V. It can be seen from Table IV tha~ even when the catalyst compo~i
All catalysts contaminated with 5000 parts per million of nickel equivalents.
The unusual results obtained by operation in accordance with the process defined herein are seen from the data in Table:III. Thus, in Run No. 1, wherein the process wa~ operated with a commercially available high activity catalyst, which has excellent metals tolerant characteris~ics when used in catalyti crack-ing of hydrocarbonaceQus feeds, excellent results were obtained, even with the catalyst carrying 5000 ppm nickel e~uivalents. When in Run No. 2, . the zeolite catalyst of Run No. 1 was diluted with meta-kaolin in a wei~ht ratio of 60:40, following `:
~$.~
the teachings of U.S. Pat. Nc. 4,289,605 of Bartholic, inferior results were obtained compared with those obtained in ~n No. 1, in tha~ conversion was reduced to 49.43 percent, with a drop in gasoline production.
Using silica in place of meta-kaolin in Run No. 3 similarly produced inferior results. However, when the zeolitic catalyst was combined with a h~at-stable refractory aluminum oxide alone, as in Run No. 6 or wi.th aluminum oxide in combination with other heat-stable metal compounds in ~uns Nos. 4, 5 and 7 to 10, con~er-sions and amounts of gasoline were almost as good or even better than the results obtained in Run No. 1.
This is surpri~ing, in that the diluents used in Runs Nos. 4 to 10 do not contain zeolite, and yet when a portion of the ca~alytically active componen~ was replaced with such diluent, excellen~ results were ~till obtained.
An additional series of run~ was carried out similarly to Runs NOB. 1 to 10 above wherein catalyst composition mixtures were e~ployed containing 60 weight p~xcent of GRZ-l and 40 wei~ht percent of ORE-Alumina.
In Run No. 11 the catalyst composition was free of any metal Gontent~ while in Runs Nos~ 1~, 13, 14 and 15 the catalyst compositions used carried 0.5/ 1.0, 1.5 and 2.0 25 weighk percent vanadium, respectively, on th~ir sur-faces. Vanadium was deposited on the catalyst composi~
tion surfaces using vanadi~m naphthe~ate following the procedure of Run~ Nos. 1 to 10. The data o~tained are tabulated below in Table IV.
TABLE IV
Conver- C ~
sion, 5 Hydrogen, Vol. % (Gasol;ne) Carbon, Wt % of Run Vanadium, of Fresh Vol. % of Wt % of Fresh No. Wt %Feed Fresh Feed Catalyst Feed ll 0 72.0 50.0 ~.2 0.03 l2 0.56~.5 48.9 4.0 O.l9 l3 l.0~7.0 44.0 5.3 0.37 l~ l.56~.0 40.0 4.6 0.42 2.054.0 33.0 4.5 0.41 ... .... ~
Still additional run~ were carried out as above but wherein GRZ-l alone was used free of any metal content and with vanadium presen~ in amounts o l.0/ 1.5 and 2.0 weight percents. The data obtained are tabu-lated below in Table V.
TABLE V
_ , .. . . .
Conver- C ~
sion, 5 Hydrogen, Vol. % ~Gasoline) Carbon, Wt % of Run Vanadium, oP Fresh YolO % ofWt % oP Fresh No. Wt % Feed - Fresh Feed Catalyst Feed 16 0 78.0 54.0 5.8 0.02 l7 l.0 57.0 40.0 2.5 0.20 18 1.5 41.0 31.0 1.8 0.13 l9 2.0 2B.O 13.0 1.4 O.ll ;
The advantages o~ operating the defined ~rocess using the novel catalyst herein are further apparent from the data in Tahles IV and V. It can be seen from Table IV tha~ even when the catalyst compo~i
5 tion herein carried up to 1 weight percent vanadium (10,000 ppm ~ranadium) in ~UIl N<:. 13, the level of conversion and the amount of gasoline produced was not greatly reduced over the same composition free of metal in Run No. 11. In fact even at a level of 2 . O weight 10 percent o~ vanadium in Run No. 15, con~Tersion and amount of gasoline produced was still appreciable. Comparable runs with ~Z-l alone produced far inferior results.
Obviously many modifications and variations of the inventio~, as herein above set for~h/ can be made without departing from the spirit and scope thereof and, therefore, only such limitations should be imposed as are i~dicated in the appended claims~
Obviously many modifications and variations of the inventio~, as herein above set for~h/ can be made without departing from the spirit and scope thereof and, therefore, only such limitations should be imposed as are i~dicated in the appended claims~
Claims (31)
1. A process for the catalytic cracking of a high metals content charge stock which comprises contacting said charge stock under catalytic cracking conditions with a catalyst composition comprising a cracking catalyst having high activity and a diluent selected from the group consisting of alumina and alumina in combination with a heat-stable metal compound, said diluent having a surface area of about 30 to about 1000 m2/gram and a pore volume of about 0.05 to about 2.5 cc/gram.
2. The process of claim 1 wherein said diluent has a surface area of about 50 to about 600 m2/gram and a pore volume of about 0.1 to about 1.5 cc/gram.
3. The process of claim 1 wherein said diluent has an average pore radius of about 10 to about 200 A.
4. The process of claim 2 wherein said diluent has an average pore radius of about 20 to about 110 A.
5. The process of claim 1 wherein said diluent is alumina.
6. The process of claim 1 wherein said diluent contains alumina and up to about 90 weight percent of heat-stable metal compound.
7. The process of claim 1 wherein said diluent contains alumina and up to about 50 weight percent of a heat-stable metal compound.
8. The process of claim 6 wherein said heat-stable metal compound is at least one metal oxide of silicon, magnesium, calcium, phosphorus, boron or zirconium.
9. The process of claim 7 wherein said heat-stable metal compound is at least one metal oxide of silicon, magnesium, calcium, phosphorus, boron or zirconium.
10. The process of claim 1 wherein the weight ratio of said cracking catalyst to diluent is in the range of about 10:90 to about 90:10.
11. The process of claim 1 wherein the weight ratio of said cracking catalyst to diluent is in the range of about 50:50 to about 70:30.
12. The process of claim 1 wherein said cracking catalyst has a MAT activity above about 1Ø
13. The process of claim 1 wherein said cracking catalyst has a MAT activity of about 1.0 to about 4Ø
14. The process of claim 1 wherein said cracking catalyst is an amorphous silica-alumina catalyst.
15. The process of claim 1 wherein said cracking catalyst is a cross-linked clay.
16. The process of claim 1 wherein said cracking catalyst is a synthetic mica-montmorillonite.
17. The process of claim 1 wherein said cracking catalyst contains a crystalline aluminosilicate.
18. The process of claim 1 wherein said cracking catalyst contains a stabilized hydrogen crystalline aluminum silicate.
19. The process of claim 1 wherein said cracking catalyst contains a rare earth-exchanged crystalline aluminum silicate.
20. The process of claim 1 wherein said cracking catalyst comprises from about ten to about 60 weight percent of a zeolite having cracking characteristics dispersed in a refractory metal oxide matrix,
21. The process of claim 1 wherein said cracking catalyst comprises from about ten to about 40 weight percent of a zeolite having cracking characteristics dispersed in a refractory metal oxide matrix.
22. The process of claim 1 wherein said cracking catalyst comprises from about 20 to about 40 weight percent of a zeolite having cracking characteristics dispersed in a refractory metal oxide matrix.
23. The process of claim 1 wherein the charge stock is a petroleum charge stock boiling above about 343°C. at atmospheric pressure.
24. The process of claim 1 wherein the charge stock is a residual charge stock.
25. The process of claim 20 wherein the zeolite is a synthetic faujasite.
26. The process of claim 20 wherein the zeolite is at least one synthetic faujasite selected from the group consisting of type Y and type X.
27. The process of claim 26 wherein the X and Y
zeolites are rare earth exchanged.
zeolites are rare earth exchanged.
28. the process of claim 20 wherein the matrix is substantially crystalline.
29. The process of claim 20 wherein the matrix is substantially amorphous.
30. The process of claim 1 wherein the charge stock has a total metals concentration as calculated in accordance with the relationship 10[Ni] + [V] + [Fe] ? 10, wherein [Ni], [V] and [Fe] are the concentrations of nickel, vanadium and iron in parts per million by weight.
31. The process of claim 1 wherein the charge stock has a total metals concentration as calculated in accordance with the relationship 10[Ni] + [V] + [Fe] ? 100, wherein [Ni], [V] and [Fe] are the concentrations of nickel, vanadium and iron in parts per million by weight.
Applications Claiming Priority (2)
| Application Number | Priority Date | Filing Date | Title |
|---|---|---|---|
| US35416282A | 1982-03-03 | 1982-03-03 | |
| US354,162 | 1982-03-03 |
Publications (1)
| Publication Number | Publication Date |
|---|---|
| CA1206112A true CA1206112A (en) | 1986-06-17 |
Family
ID=23392117
Family Applications (1)
| Application Number | Title | Priority Date | Filing Date |
|---|---|---|---|
| CA000417682A Expired CA1206112A (en) | 1982-03-03 | 1982-12-14 | Process for cracking high metals content feedstocks |
Country Status (2)
| Country | Link |
|---|---|
| CA (1) | CA1206112A (en) |
| GB (1) | GB2116202A (en) |
Families Citing this family (5)
| Publication number | Priority date | Publication date | Assignee | Title |
|---|---|---|---|---|
| US4859313A (en) * | 1985-02-20 | 1989-08-22 | Uop Inc. | Method for reducing the temperature in a regeneration zone of a fluid catalytic cracking process |
| US4708786A (en) * | 1986-03-26 | 1987-11-24 | Union Oil Company Of California | Process for the catalytic cracking of nitrogen-containing feedstocks |
| US4747935A (en) * | 1986-03-26 | 1988-05-31 | Union Oil Company Of California | Process for the catalytic cracking of feedstocks containing nitrogen |
| US4975180A (en) * | 1989-06-05 | 1990-12-04 | Exxon Research And Engineering Company | Cracking process |
| EP3222345A1 (en) * | 2008-07-30 | 2017-09-27 | King Fahd University Of Petroleum&Minerals | Metallic clay based fcc gasoline sulfur reduction additive compositions |
-
1982
- 1982-11-26 GB GB08233796A patent/GB2116202A/en not_active Withdrawn
- 1982-12-14 CA CA000417682A patent/CA1206112A/en not_active Expired
Also Published As
| Publication number | Publication date |
|---|---|
| GB2116202A (en) | 1983-09-21 |
Similar Documents
| Publication | Publication Date | Title |
|---|---|---|
| US4465779A (en) | Modified cracking catalyst composition | |
| US4465588A (en) | Process for cracking high metals content feedstock | |
| US4466884A (en) | Process for cracking high metals content feedstocks using a cracking catalyst mixture containing antimony and/or tin | |
| US3944482A (en) | Process for the cracking of high metals content feedstocks | |
| US4311580A (en) | Selective vaporization process and dynamic control thereof | |
| US5002653A (en) | Catalytic cracking process with vanadium passivation and improved | |
| CA1206458A (en) | High performance catalysts for carbometallic oil conversion and their manufacturing and use | |
| US4263128A (en) | Upgrading petroleum and residual fractions thereof | |
| CA2384030C (en) | Method for reducing gasoline sulfur in fluid catalytic cracking | |
| US4988654A (en) | Dual component cracking catalyst with vanadium passivation and improved sulfur tolerance | |
| CA1217753A (en) | Dual function cracking catalyst (dfcc) composition | |
| US4414098A (en) | Upgrading carbo-metallic oils with used catalyst | |
| CA1189845A (en) | Process and catalyst for the conversion of carbo- metallic containing oils | |
| US4650564A (en) | Process for cracking high metals content feedstocks | |
| US4944865A (en) | Process for cracking high metals content feedstocks | |
| CA1088912A (en) | Hydrocarbon catalytic cracking process | |
| CA1206112A (en) | Process for cracking high metals content feedstocks | |
| AU643736B2 (en) | Dual component cracking catalyst and process for vanadium passivation and improved sulfur tolerance | |
| EP0914401A1 (en) | Catalysts for optimum bottoms cracking of heavy feeds | |
| AU2002258676B2 (en) | Zeolite based catalyst of ultra-high kinetic conversion activity | |
| GB2116062A (en) | Catalytic cracking composition | |
| Nishimura | Development of catalytic cracking process and catalysts | |
| Occelli | Recent trends in fluid catalytic cracking technology | |
| US4859313A (en) | Method for reducing the temperature in a regeneration zone of a fluid catalytic cracking process | |
| CA1041031A (en) | Process for the cracking of high metals content feed stocks |
Legal Events
| Date | Code | Title | Description |
|---|---|---|---|
| MKEX | Expiry |