Note : Les descriptions sont présentées dans la langue officielle dans laquelle elles ont été soumises.
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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1089390
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 -
390
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
-- 4
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lVI~9;~90 :
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
1089390
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.
390
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.
EXAMPLE 1
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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-- 10 --
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
390
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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390
EXAMPLE 2
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
- 13 -
1089390
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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lV89390 ~:
EXAMPLE 3
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
- 15 -
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:
- 16 -
. ' ' ~ ~ .
10~3390 :~
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
- 17 -
:
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
- 18 -
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
-- 19 - -
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.
EXAMPLE 4
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:
.
~ - 20 -
1~8~3~(~
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
- 21 -
.
1089390
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.
.,
. :
- 22
. .
