CA1089390A - Process for cracking high metals containing hydrocarbons - Google Patents

Abstract

PROCESS FOR CRACKING HIGH METALS
CONTAINING HYDROCARBONS

ABSTRACT OF THE DISCLOSURE
The cracking of high metals-containing hydrocarbons to obtain high yields of gasoline and reduced yields of carbon and hydrogen is accomplished while employing a catalyst containing a high concentration of metal contaminants.

Description

BAC~GROUND OF THE INVENTION
Conventionally feedstocks to catalytic cracking processes operated to obtain a high yield of gasoline and other low boiling fractions must contain very low concentrations of metals, normally less than 2 parts per million (ppm) and prefer-ably no greater than 1 ppm. The metals in the feed to the process are accumulated on the catalyst as hereafter described, substantially reducing the activity of the catalyst which results in low conversion of the feed to the lower boiling range products.
The metals present in the petroleum charge stocks to the catalytic cracking processes are generally in an organo-metallo form, such as in a porphyrin ring or as a naphthenate.
These metals tend to deposit in a relatively non-volatile form onto the catalyst during the cracking process, and regeneration of the catalyst to remove coke does not remove these contaminant metals. Metals found to be present in hydrocarbon feeds to catalytic processes include nickel, vanadium, copper, chromium and iron.
A number of methods have been proposed to reduce the co~centration of metals in feedstocks to catalytic cracking processes which typically employ zeolitic cracking catalysts.
It has been suggested that the contaminated feed be pretreated to reduce the concentration of metals to below about 1 ppm or ,, - , .
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to exclude by fractionation the heavier gas oil and residual fractions where the major concentration of metal contaminants occur. However, as the necessity increases for converting heavier feedstocks to lower boiling product fractions in order to satisfy the increasing demands of the market place for gasoline products, it is evident that an improved catalytic process that will permit the charging of feedstocks which contain relatively high concentrations of metals, such as residual containing hydrocarbons, is needed.
By the invention a process for the catalytic cracking of feedstocks containing at least 2 ppm nickel equivalent metal contaminants is provided with the process operated continuously until the concentration of contaminant metals in the catalyst exceeds 2,000 ppm to obtain high yields of gasoline products while producing relatively low yields of hydrogen and coke.
Thus according to the present invention there is provided a continous process which comprises contacting a hydrocarbon feed containing at least 2 ppm nickel equivalents with a catalyst under catalytic cracking conditions until said catalyst contains greater than 2000 ppm nickel equivalents as metal contaminants, and continuously recovering therefrom a gasoline boiling range product fraction, said catalyst consisting essentially of greater than 50 weight percent of a refractory oxide and greater than 30 weight percent of a zeolitic component, said catalyst further having a surface area of at least 150 square meters per gram and an average pore diameter of at least 20 A and a nitrogen pore volume of at least 0.12 cc per gram when the maximum catalyst temperature employed in preparing the catalyst does not exceed 1025F.

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10l~9390 The catalysts employed in the process of this invention 21re zeolitic-containing catalysts wherein the concentration of the zeolite component is greater than 30 percent by weight. Suitable c!atalysts comprise a crystalline aluminosilicate dispersed in a refractory metal oxide matrix such as disclosed in U. S. Letters Patent 3,140,249 and 3,140,253 to C. J. Plank and E. J. Rosinski.
Suitable matrix materials comprise inorganic oxides such as amorphous and semi-crystalline silica-aluminas, silica-magnesias, silica-alumina-magnesia, alumina, titania, zirconia, and mixtures 10 thereof.
Zeolites or molecular sieves having cracking activity and suitable in the preparation of the catalysts of this invention are crystalline, three-dimensional, stable structures containing a - 2a -large number of uniform openings or cavities interconnected by smaller, relatively uniform holes or channels. The formula for the zeolites can be represented as follows.
XM2/no:Al2o3:l~5-6 5 SiO2 YH2O

where M is a metal cation and n its valence; x varies from 0 to 1;
and y is a function of the degree of dehydration and varies from 0 to 9. M is preferably a rare earth metal cation such as lanthanum, cerium, praseodymium, neodymium or mixtures thereof.
Zeolites which can be employed in the practice of this invention include both natural and synthetic zeolites. These natural occurring zeolites include gmelinite, chabazite, dachiardite, clinoptilolite, faujasite, heulandite, analcite, levynite, erionite, sodalite, cancrinite, nepheline, lazurite, scolecite, natrolite, offretite, mesolite, mordenite, brewsterite, ferrierite, and the like. Suitable synthetic zeolites which can be employed in the inventive process include zeolites X, Y, A, L, ZK-4, B, E, F, H, J, M, Q, T, W, Z, alpha and beta, ZSM-types and omega. The effective pore size of synthetic zeolites are suitable between 6 and 15 A in diameter. The term "zeolites" as used herein contemplates not only aluminosilicates but substances in which the aluminum are replaced by gallium and substances in which the silicon is replaced by germanium. The preferred zeolites are the synthetic faujasites of the types Y and X or mixtures thereof.
It is also well known in the art that to obtain good cracking activity the zeolites must be in good cracking form. In most cases this involves reducing the alkali metal content of the zeolite to as low a level as possible as a high alkali metal content reduces the thermal structural stability, and the effective lifetime of the catalyst is impaired. Procedures for removing 10~3390 alkali metals and putting the zeolite in the proper form are well known in the art and are as described in U. S. ~etters Patent 3,534,816.
The amount of the zeolitic material to be dispersed in the matrix is greater than 30 percent by weight. Conventional methods can be employed to form the final catalyst composite.
For example, finely divided zeolite can be admixed with the finely divided matrix material, and the mixture spray dried to form the catalyst composite. Other suitable methods of dispersing 0 - the zeolite materials in the matrix materials are described in U. S. Patents 3,271,418; 3,717,587; 3,657,154; and 3,676,330, In addition to having a concentration of zeolitic material greater than 30 percent by weight, the catalyst composite employed in the process of this invention as manufactured should have a surface area of at least 150 square meters per gram and an average pore diameter of at least 20 A units, and a pore volume of at least 0.12 cc per gram as determined by the nitroqen adsorption test method described by E. V. Ballou, o. K. Dollen, in Analytical Chemistry, Volume 32, page 532, 1960 when the maximum catalyst temperature employed in preparing the catalyst composite does not exceed 1025F. (552~C.) The above-identified catalysts are employed in the cracking of high metals content feed stocks. Such feedstocks include heavy gas oils, residuum or other petroleum fractions which are suitable catalytic cracking feed or charge stocks except for the high metals concentrations. The charge stocks to the inventive process can also be derived from coal, shale, or tar sands. The high metals content charge stocks employed in this inventive process are those having a total metals .
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concentration of at least 2.0 as calculated in accordance with the following relationship:

[Ni] + 0.2[V] ~ 2.0 whexe [Ni] and [V] are the concentration of nickel and vanadium, respectively, in parts per million by weight. Charge tocks having a concentration of metals of at least 2.0 ppm cannot be treated economically in existing commercial processes due to the high catalyst make-up rates required to maintain adequate product yields. By the invention, these charge stocks can be catalytically cracked economically employing the heretofore described catalyst composite and the hereafter described process conditions.
A preferred method of operating the process of this invention i8 by fluid catalytic cracking using riser outlet temperatures between about 900 and 1100F. (482 to 593C.) Under such conditions, the cracking occurs in the presence of a fluidized composited catalyst in an elongated reactor tube which i9 commonly referred to as a riser. Generally, the riser has a length to diameter ratio of above 20. The charge stock is passed through a preheater to heat the feed to a temperature of about 600F. (316C.) and then charged into the bottom of the riser.
In operation, a residence time of up to five seconds and catalyst to oil weight ratios of about 4:1 to about 12:1 to 15:1 are employed. Steam can be introduced into the oil inlet line to the riser and/or introduced independently to the bottom of the riser so as to assist in carrying regenerated catalyst upwardly through the riser. Regenerated catalyst at temperatures generally between 1100 and 1350F. (593 to 7320C.)is introduced into the bottom of the riser.
The riser system at a pressure in the range of about 5 to about 50 psig (0.35 to 3.5 Kg/cm2) is normally operated 10~9390 with catalyst and hydrocarbon feed flowing concurrently into and upwardly into the riser at about the same velocity, thereby avoiding any significant slippage of catalyst relative to hydro-carbon in the riser and avoiding formation of a catalyst bed in the reaction flow stream. In this manner the catalyst to oil ratio thus increases significantly from the riser inlet along the reaction flow stream.
The riser temperature drops along the riser length due to heating and vaporization cf the feed and by the slightly endothermic nature of the cracking reaction and heat loss to the atmosphere. As nearly all the cracking in the system occurs within one or two seconds, it is necessary that feed vaporization occurs nearly instantaneously upon contact of feed and regenerated catalyst at the bottom of the riser. Therefore, at the riser inlet, the hot, regenerated catalyst and preheated feed, generally together with a mixing agent such as steam, nitrogen, methane, ethane or other light gas, are intimately admixed to achieve an equilibrium temperature nearly instantaneously.
In those instances where residual components are included in the feed, the inlet temperature should be relatively high to vaporize the major portion of the feed as residual feed components that do not vaporize remain on the hot catalyst and tend to be converted to coke, thereby resulting in a loss of useful product and the lowering of catalyst activity. This deleterious effect occurs in addition to the deposition upon the catalyst of the metals content of the residual oil, especially nickel and vanadium, further tending to reduce catalyst activity and selectivityO
Catalytic processes employing conventional catalysts can operate at contaminated metals higher than a 1000 ppm nickel equivalents but product yields are adversely affected. It is also known that some conventional commercial cracking catalysts have been employed under cracking conditions such that the cataly~t may accumulate about 1500 ppm nickel equivalents, Bly employing the defined catalyst in the manner of this invention, the contaminant metals level on the catalyst can exceed 3000 ppm nickel equivalents with le~s than 4 volume % reduction in the production of gasoline and lighter productsbased upon the feed. Yields of gasoline and carbon are unaffected significantly up to contaminant levels of 5000 ppm nickel equivalents. Although -hydrogen yields increase with increasing metals contamination above 3000 ppm, the rate of increase is substantially less than that normally obtained in conventional hydrocarbon cracking processes. Thus, the process of this invention enables the cracking process to be operated efficiently with the metal contaminant concentration on the catalyst up to at least 5000 ppm nickel equivalents.
The process of this invention has a number of advantages over conventional catalytic cracking processes by providing an economically attractive method to include higher metals-containing gas oils as a feed to the catalytic cracking process. As previously indicated, because of the loss of selectivity to high value products (loss of conversion and yield of gasoline, and gain in coke and light gases) with the increased metals contamination on conventional zeolite catalysts, most refiners attempt to maintain a low metals level on cracking catalysts. A method of controlling the metals concentration on the catalyst is to omit feed components containing high concentration of metals. A second costly method is to increase the catalyst makeup rate higher than that required to maintain activity or to satisfy unit losses.

The following examples are presented to illustrate objects and advantages of the invention. However, it is not intended that the invention should be limited to the specific em~bodiments presented therein.

In this example, a fluid catalytic cracking process, illustrative of the invention, was employed to process a hydro-carbon cracking feed characterized as follows:

Gravity: API 25.0 Sulfur: wt% 0.31 Nitrogen: wt% 0.12 Carbon Residue, Rams, ASTM D525:
wt% 0 77 Aniline Point, ASTM D611: F 199 (92.8C) Viscosity, ASTM D2161, 210F (98.9C): SUS 49.8 Pour Point, ASTM D97: F + 90 (+32.2C) Nickel: ppm 1.2 Vanadium: ppm 0.4 -Vacuum Distillation, ASTM D1160:
F
10% at 760 mm Hg 622 (327.8C) 30~ 716 (380C) 50% 797 (425C) 70% 885 (473.9C) 90~ 1055 (568.3C) Calc. Carbon Type compositions, vol. Fract.
Aromatics (CA) 0.15 Naphthenes (CN) 0.26 Parafins (Cp) 0-59 ' ~

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10~390 ~
The catalyst employed in this example contained greater than 30.0 weight percent zeolite, greater than 50 weight percent clay, and was further characterized as follows:
Alumina: wt~ 29.3 Surface Area: m2/g 253 Nitrogen Pore Volume: cc/g 0.15 Nitrogen Avg. Pore Diameter: A 24 Apparent Bulk Density: g/cc 0.785 Compacted Bulk Density: kg/dm 0.903 Particle Size Distribution 0-20 Microns 8.0 20-40 Microns 20.5 40-80 Microns 33.7 80 Microns 37.8 ~80/ c40~ 1.23 :
The above-described catalyst after being subjected to a heat shock treatment for 1.5 hours at 649C. and a steam purge for eight (8) hours at 718C. and containing 912 ppm nickel and 187 ppm vanadium was charged to a riser reactor with the above-described hydrocarbon feed. The riser reactor was operated at the following conditions:
Hydrocarbon Feed Preheat Temperature: F 496 (258C) ; Catalyst Preheat Temperature:
F 1202 (650C) Catalyst To Oil Ratio: wt cat/wt Fresh Feed (FF) 8.7 Reaction Zone Average Temperature:
F 994 t534C) `30 Riser Outlet Temperature: F 982 (528C) Contact Time, Based on Feed:
sec 3.18 Carbon On Catalyst: wt%
Spent 0.84 Regenerated 0.30 Riser Pressure: psig 26.0 (1.82 kg/cm ) The products obtained during this run were as follows:

g _ r33~0 Product Yields: Vol ~ Of FF
Slurry Oil ~650+F (343+C) TBP] 12. 0 Furnace Oil [650F (343C) TBP] 10.2 Debut. Gaso. [430F (221C) TBP EP] 56.2 Depent. ~aso. [430F (221C) TBP EP] 43.3 Heavy Gaso. [430F (221C) TBP EP] 20.9 Depentanized Light Gasoline 22.4 Light Hydrocarbons:
Total Pentanes-Pentenes 12.9 l-Pentane 7.9 N-Pentane 1.0 Pentenes 4.1 Total Butanes-Butenes 18.9 l-Butane 8.4 N-Butane 2.2 Butenes 8.3 Total Propane-Propylene 11.7 Propane 3.2 Propylene 8.5 Total C3+ Liquid Yield: Vol % FF 108.9 Conv. To 430F (221C) EP Gaso. And Lighter WT % Of FF 76.1 VOL % Of FF 77.8 Product Yields: wt ~ Of FF
C2 And Lighter 2.9 Total Ethane-Ethylene 1.4 Ethane 0.6 Ethylene 0.8 Methane 1.1 Hydrogen 0.18 Hydrogen Sulfide 0.1 Coke By Flue Gas An~lysis9.7 .

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1~89390 In a second run the above-described catalyst now containing 2051 ppm nickel and 412 ppm vanadium and the above-described hydrocarbon feed, now containing 44.5 ppm nickel and 8.8 ppm vanadium, was charged to a riser reactor operated at the following conditions:
Hydrocarbon Feed Preheat Temperature: F 530 (277C) Catalyst Preheat Temperature:
F 1194 (646C) Catalyst To Oil Ratio: wt cat/wt Fresh Feed (FF) 8.3 Reaction Zone Average Temperature:
F 993 (534C) Riser Outlet Temperature: F 985 (530C) Contact Time, Based on Feed:
sec 3.24 Carbon On Catalyst: wt%
Spent 0.84 Regenerated 0.3 Riser Pressure: psig26.2 (1.82 kg/cm ) The products obtained during this run were as follows:
Product Yields: Vol % Of FF
Slurry Oil [650+F (343+C) TBP] 11.8 Furnace Oil [650F (343C) TBP] 10.3 Debut. Gaso. [430F (221C) TBP EP] 53.1 Depent. Gaso. [430F (221C) TBP EP] 40.5 Heavy Gaso [430F (221C) TBP EP] 20.6 Depentanized Light Gasoline 20.0 ~ight Hydrocarbons:
Total Pentanes-Pentenes 12.5 l-Pentane 8.6 N-Pentane 0.8 Pentenes 3.2 Total Butanes-Butenes19.1 l-Butane 8.4 N-Butane 2.2 Butenes 8.4 Total Propane-Propylene11.7 Propane 3.1 Propylene 8.7 Total C3+ Liquid Yield: Vol % FF 106.0 Conv. To 430F (221C) EP Gaso.
And Lighter WT % Of FF 76.2 VOL % Of FF 77.9 --~
Product Yields: wt % Of FF
C2 And Lighter 3.0 Total Ethane-Ethylene1.4 Ethane 0.6 Ethylene 0.8 Methane 1.2 Hydrogen 0.31 -Hydrogen Sulfide 0.1 Coke By Flue ~as Analysis 11.0 A comparison of the results obtained in the two runs of this example, with the riser reactor operated at essentially the same conditions, demonstrates that by the invention catalyst containing above 2000 ppm nickel equivalents can be employed in the cracking of feedstocks containing 45 ppm nickel equivalents with no reduction in conversion and only a minimal reduction in the yield of debutanized gasoline (56.2 to 53.1 volume percent of fresh feed). Coke production when operating under the process conditions of this invention increased only 13.4% and the pro-duction of light gases (C2 and lighter) was not increased.

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In this example the catalytic cracking catalyst com-pGSition of Example 1, now containing 3614 ppm nickel and 681 ppm vanadium, was employed in a fluid catalytic cracking process with the hydrocarbon feed of the second run of Example 1 (containing 44.5 ppm nickel and 8.8 ppm vanadium). Regenerated catalyst and the hydrocarbon feed was charged to a riser reactor operated at the following conditions:

Hydrocarbon Feed Preheat Temperature: F 516 (269C) Catalyst Preheat Temperature:F 1186 (641C) Catalyst To Oil Ratio: wt cat/wt Fresh Feed (FF) 8.3 Reaction Zone Average Temperature F 981 (527C) Riser Outlet Temperature: F 969 (521C) Contact Time, Based on Feed: sec 3.33 Carbon On Cata~yst: wt%
Spent 0.84 Regenerated 0.30 Riser Pressure: psig 26.3 (1.82 kg/cm ) The products obtained during this run were as follows:
Product Yields: Vol % Of FF

Slurry Oil [650+F (343+C) TBP 14.9 Furnace Oil [650F (343C) TBP 11.9 Debut. Gaso. [430F (221C) TBP EP] 52.3 Depent. Gaso. [430F (221C) TBP EP] 40.9 Heavy Gaso. [43OF (221C) TBP EP] 22.4 Depentanized Light Gasoline 18.4 Light Hydrocarbons:

Total Pentanes-Pentenes 11.5 l-Pentane 7.1 N-Pentane 0.7 Pentenes 3.7 Total Butanes-Butenes 17.9 l-Butane 7.3 N-Butane 1.7 Butenes 8.9 Total Propane-Propylene10.5 Propane 2.6 Propylene 7.9 Total C3~ Liquid Yield: Vol % FF 107.6 Conv. To 430 F (221C) EP Gaso.
And Lighter WT % Of FF 71.4 ; VOL % Of FF 73.2 Product Yields: wt % Of FF
C2 And Lighter 2.7 Total Ethane-Ethylene1.2 Ethane 0.5 Ethylene 0.7 Methane l.1 Hydrogen 0.41 Hydrogen Sulfide 0.0 Coke By Flue Gas Analysis 8.3 A comparison of the results obtained in this run with the first run of Example 1 (illustrative of a conventional process) demonstrates that by the invention catalyst containing above 3500 ppm nickel equivalents can be employed in the cracking of feedstocks containing 45 ppm nickel equivalents with only 4.6 volume percent reduction in conversion and only 3.9 volume percent reduction in the yield of debutanized gasoline. Coke production when operating under the run of this example was less than that produced under the first run of Example l and the production of light gases (C2 and lighter) was also less than that produced in the first run of Example l.

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In this example and in subsequent Example 4, a fluid catalytic cracking process illustrative of the invention, was employed to process a hydrocarbon cracking feed characterized as follows:
Gravity: API 24.7 Sulfur: wt% 0.17 Nitrogen: wt% 0.13 Carbon Residue, Rams, ASTM D525:
wt% 0 4 Aniline Point, ASTM D611: F 185 (85C) Viscosity, ASTM D2161, 210F (98.9C): SUS 45.7 Pour Point, ASTM D97: F + 100 (+38.8C) Nickel: ppm 1.0 Vanadium: ppm 0.2 Vacuum Distillation, ASTM D1160:
F
10% at 760 mm Hg 598 (314.4C) 30% 707 (375C) 50% 786 (418.9C) 70% 860 (460C) 90% 995 (535C) Calc. Carbon Type compositions, vol. Fract.
Aromatics (CA) 0.17 Naphthenes (CN) 0.27 Paraffins (Cp) 0.56 lU~3~3~
The catalyst emplo~ed in this example contained greater than 30.0 weight percent zeolite, greater than 50 weight percent clay, and was further characterized as follows:
Alumina: wt% 55.0 Surface Area: m /g 326 Nitrogen Pore Volume: cc/g 0.36 Nitrogen Avg. Pore Diameter: A 44 Apparent Bulk Density: g/cc 0.734 Compacted Bulk Density: kg/cm3 0.808 Particle Size Distribution 0-20 Microns 1.0 20-40 Microns 8.2 40-80 Microns 30.6 ~80 Microns 60.2 > 80/ ~40~ 6.54 The above-described catalyst after being subjected to a heat shock treatment for 1.5 hours at 649C. and a steam purge for eight (8) hours at 718C. and containing 9 ppm nickel and 2 ppm vanadium was charged to a riser reactor with the above-described hydrocarbon feed. The riser reactor was operated at the following conditions:
Hydrocarbon Feed Preheat Temperature F 536 (280C) Catalyst Preheat Temperature: F 1088 (586C) Catalyst To Oil Ratio: wt cat/wt Fresh Feed (FF) 8.5 Reaction Zone Average Temperature: F 981 (527C) Riser Outlet Temperature: F 975 (524C) Contact Time, Based on Feed:
sec 8.98 Carbon On Catalyst: wt%
Spent 0.84 Regenerated 0.30 Riser Pressure: psig 26.1 (1.82 kg/cm ) The products obtained during this run were as follows:

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Product Yields: Vol ~ Of FF
Slurry Oil [650+F (343+C) TBP] 2.9 Furnace Oil [650F (343C) TBP] 9.4 Debut. Gaso. [43OF (221C) TBP
EP] 60.8 Depent. Gaso. ~430F (221C) TBP EP] 45-3 Heavy Gaso. [430F (221C) TBP EP] 23.1 Depentanized Light Gasoline 22.6 Light Hydrocarbons:
Total Pentanes-Pentenes15.0 l-Pentane 10.9 N-Pentane 1.3 Pentenes 2.8 Total Butanes-Butenes 23.0 l-Butane 11.4 N-Butane 3.5 Butenes 8.1 Total Pxopane-Propylene13.5 Propane 4.6 Propylene 8.9 Total C3+ Liquid ~ield: Vol % FF 109.6 Conv. To 430F (221C) EP Gaso. And Lighter WT % Of FF 86.2 VOL % Of FF 87.7 Product Yields: wt % Of FF
C2 And Lighter 3.7 Total Ethane-Ethylene 2.1 Ethane 0.9 Ethylene 1.2 Methane 1.4 ; Hydrogen 0.07 Hydrogen Sulfide 0.1 Coke By Flue Gas Analysis 10.8 :

101~390 In a second run the above-described catalyst of this example now containing 2523 ppm nickel and 449 ppm vanadium and the above-described hydrocarbon feed, now containing 44.0 ppm nickel and 7.5 ppm vanadium, was charged to a riser reactor operated at the following conditions:
E~ydrocarbon Feed Preheat Temperature: F 522 (272C) -Catalyst Preheat Temperature: F 1128 (609C) Catalyst To Oil Ratio: wt cat/wt Fresh Feed tFF) 8.3 Reaction Zone Average Temperature:
F 985 (529C) Riser Outlet Temperature: F 981 (527C) Contact Time, Based on Feed:
sec 6.72 Carbon On Catalyst: wt%
Spent 0.84 Regenerated 0.3 Riser Pressure: psig25.9 (1.82 kg/cm ) The products obtained during this run were as follows:
Product Yields: Vol ~ Of FF
Slurry Oil ~650+F (343+C) TBP] 3.5 Furnace Oil [650F (343C) TBP] 8.2 Debut. Gaso. [430F (221C) TBP EP] 61.3 Depent. Gaso. [430F (221C) TBP EP] 48.2 Heavy Gaso. [430F (221C) TBP EP] 25.5 Depentanized Light Gasoline 22.7 Light Hydrocarbons:
Total Pentanes-Pentenes 13.1 l-Pentane 8.6 N-Pentane 1.0 Pentenes 3.5 10~9390 Total Butanes-Butenes 21.8 l-Butane 9.5 N-Butane 2.5 .` Butenes 9.8 Total Propane-Propylene12.5 Propane 3.0 Propylene 9.5 Total C3+ Liquid Yield: Vol % FF 107.2 Conv. To 430F (221C) EP Gaso.
And Lighter WT & Of FF 86.9 VOL % Of FF 88.4 Product Yields: wt % Of FF
C2 And Lighter 3.0 Total Ethane-Ethylene1.4 Ethane 0.6 Ethylene 0.8 : Methane 1.1 Hydrogen 0.31 Hydrogen Sulfide 0.1 Coke By Flue Gas Analysis 12.4 1(~l3~390 A comparison of the two runs of this example demon- :
; strates that when operated in accordance to the invention at metals contaminants in excess of 2500 ppm nickel equivalents that there is no reduction in conversion and debutanized gasoline. The production of light gases employing the catalyst composition containing in excess of 2500 ppm nickel equivalents is not increased when compared with the first run of this example and the production of coke is increased by only 14.7 weight percent.

In this example the catalytic cracking catalyst com-position of Example 2, now containing 4918 ppm nickel and 834 ppm vanadium, was employed in the fluid catalytic cracking pro- ~ -cess with the hydrocarbon feed of the second run of Example 3 tcontaining 44.0 ppm nickel and 7.5 ppm vanadium). Regenerated catalyst in the hydrocarbon feed was charged to a riser reactor operated at the following conditions:

Hydrocarbon Feed Preheat Temperature:F 516 (269C) Catalyst Preheat Temperature: F 1137 (614C) Catalyst To Oil Ratio: wt cat/wt Fresh Feed ~FF) 8.5 Reaction Zone Average Temperature:
F 987 (531C) Riser Outlet Temperature: F 981 (527C) Contact Time, Based on Feed:
: sec ~.96 Carbon On Catalyst: wt%
Spent 0.84 Regenerated 0.30 Riser Pressure: psig 24.9 (1.75 kg/cm ) The products obtained during the run of this example were as follows:

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Product Yields: Vol % Of FF
Slurry Oil [650+F (343+C) TBP] 4.1 Furnace Oil [65OF (343C) TBP] 13.0 Debut. Gaso. [430F (221C) TBP
EP] 59.2 Depent. Gaso. [430F (22lC) TBP EP] 47.0 Heavy Gaso. [43OF (221C) TBP EP] 24.7 Depentanized Light Gasoline 22.3 Light Hydrocarbons:
Total Pentanes-Pentenes12.2 -l-Pentane 6.9 N-Pentane 0.8 Pentenes 4.5 Total Butanes-Butenes19.5 l-Butane 7.3 N-Butane 1.8 Butenes 10.3 Total Propane-Propylene11.1 Propane 2.4 Propylene 8.7 Total C3+ Liquid Yield: Vol % FF 106.9 Conv. To 430F (221C) EP Gaso. And Lighter WT % Of FF 81.3 VOL % Of FF 82.9 Product Yields: wt % Of FF
C2 And Lighter 2.9 Total Ethane-Ethylene 1.4 Ethane 0.6 Ethylene 0.8 Methane 1.0 Hydrogen 0.44 Hydrogen Sulfide 0.1 Coke By Flue Gas Analysis 10.9 .

A comparison of the results obtained in the run of this example with the first run of Example 3 demonstrates that t:he process of this invention when employing a catalyst compo-~;ition containing in excess of 5000 ppm nickel equivalents as metal contaminants only reduces the conversion from 87.7 to 82.9 volume percent based upon the feed. The reduction in debutanized gasoline produced by the run of this example when compared with the first run of Example 3 was only 1.6 volume percent based upon the feed to the catalytic cracking zone. Less light gases were produced in the run of this example and there was no significant increase in coke production when comparing the run of this example with the first run of Example 3.
Although the invention has been described with refer-ence to specific embodiments, references, and details, various modifications and changes will be apparent to one skilled in the art and are contemplated to be embraced in this invention.

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Claims (7)

THE EMBODIMENTS OF THE INVENTION IN WHICH AN EXCLUSIVE
PROPERTY OR PRIVILEGE IS CLAIMED ARE DEFINED AS FOLLOWS:

1. A continuous process which comprises contacting a hydrocarbon feed containing at least 2 ppm nickel equivalents with a catalyst under catalytic cracking conditions until said catalyst contains greater than 2000 ppm nickel equivalents as metal contaminants, and continuously recovering therefrom a gasoline boiling range product fraction, said catalyst consisting essentially of greater than 50 weight percent of a refractory oxide and greater than 30 weight percent of a zeolitic component, said catalyst further having a surface area of at least 150 square meters per gram and an average pore diameter of at least 20 .ANG. and a nitrogen pore volume of at least 0.12 cc per gram when the maximum catalyst temperature employed in preparing the catalyst does not exceed 1025°F.

2. The process of claim 1 wherein said process is con-ducted until said catalyst contains at least 3000 ppm nickel equivalents as metal contaminants.

3. The process of claim 1 wherein said process is conducted until said catalyst contains at least 5000 ppm nickel equivalents as metal contaminants.

4. The process of claim 1 wherein said gasoline boiling range product fraction is a debutanized gasoline and wherein catalytic cracking conditions are maintained so that the change in the volume percent of debutanized gasoline obtained when conducting the process and employing a catalyst having at least 3000 ppm nickel equivalents metal contaminants is not more than 4 volume percent less, based on the feed to the to the process, than the volume percent of the debutanized gasoline obtained from said process with the catalyst contain-ing no more than 1000 ppm nickel equivalents as metal contaminants.

5. The process of claim 1 wherein said catalytic cracking hydrocarbon feed is selected from the group consisting of petroleum gas oils and residuums.

6. The process of claim 1 wherein said catalytic cracking conditions comprise a fluid catalytic cracking riser having an outlet temperature between about 900 and 1100°F., a riser residence time of up to 5 seconds, a catalyst to oil weight ratio of between about 4:1 and about 12:1, and a pressure in the range from about 5 to about 50 psig.

7. The process of claim 6 wherein said catalytic cracking hydrocarbon feed is selected from the group consisting of petroleum gas oils and petroleum residuums.