Note : Les descriptions sont présentées dans la langue officielle dans laquelle elles ont été soumises.
7~i~
001 -1 -
002 TWO-BED CATALYTIC HYDROPROCESSING FOR
003 HEAVY ~IYDROCARBON FEEDSTOCKS
Q05 BACKG~OUND_OF _THE INVENTION
Q06 This invention relates to catalytic hydroprocessing
C07 of heavy hydrocarbon feedstocks such as crude oil, topped
008 crude, reduced crude, a~mosphexic residual oil, vacuum residual
009 oil, deasphalted atmospheric or vacuum residua, vacuum gas oil,
010 coal liquefaction product fractions such as solvent refined
011 coal (SRC) and liquid solvent refined coal (SRC II), shale oil,
012 oil from tar sands, and other heavy hydrocarbonaceous
013 materials. Heavy hydrocarbon feedstocks suitable for pro-
014 cessing according to this invention include those feedstocks
015 containing significant quantities, e.g. at least about 95~ by
016 weight materials boiling above 200C and particularly those
017feedstocks containing at least 25%, 50%, or 75~ material
015boiling above 300C or above 450C. The hydroprocessing pro-
019 cess of the present invention can perform hydrodemetalation
020 hydrodesulfurization, hydrodenitrlfication, hydrocracking,
021 and/or hydrogenation of olefinic and aromatic hydrocarbons.
022 The process is particularly useful for the hydrocracking of
023 heavy feedstocks containing nitrogen and sulfur prior to fluid
024 catalytic cracking.
025DESCRIPTION C)~ THE PRIt:)R ART
026A number of workers have described hydroprocessing of
027 heavy hydrocarbons by sequential catalytic steps. Some cata-
028 lyst arrangements provide high metals capacity in a first
02~ catalyst bed to reduce fouling of subse~uent catalysts. Other
030 systems employ successive catalytic zones individually
031 optimized for demetalation, desulfurization and denitrlflca-
032tion. U.S. Patents 4,019,976 and 3,159,568 describe the use of
033 two catalyst beds wherein the second bed contains a rnore active
034catalyst than the first. U.S. Patent 3,437,588 describes the
035 use of a mixture of hydrogenation catalysts on supports having
03620-lC0 A pores~ U.S. Patents 3,977,961 and 3,977,962 describe
037 two-stage catalyst systems containing 100-275 ~ pores in the
001 -2-
002 irst stage and 100-200 A pores in the second stage. ~.S.
003 Patent 3,696,027 describes hydrodesulfurization using a cata-
004 lyst having graded macroporosity (pores greater than 500A).
005 The graded catalyst is packed in a downflow reactor with pore
006 volume varying from greater than 30~ macropores in the upper
007 section to less than 5~ macropores in the lower sections. U.S.
008 Patents 3,254,017 and 3,535,225 describe t~o-stage hydro-
009 cracking using a first stage large pore catalyst and second
010 stage zeolites. U.S. Patent 3,385,731 suygests two-stage hydro-
011 cracking using a large pore zeolite having 10 to 13 A pores in
012 the first stage and a small pore zeolite having 4 to 6 A pores
013 in the second stage. Two-stage catalyst beds wherein the pores
014 of the second stage catalyst are larger than those of the first
015 stage catalyst are depicted in U.S. Patents 3,730,879,
016 3,766,058, and 4,048,0600
Q17 SUMMARY OF THE INVENTION
018 It is an object of this invention to provide a two-
019 stage catalyst system capable of effectively hydrocracking,
020 hydrodenitrifying and hydrodesulfurizing heavy hydrocarbon feed--
021 stocks. It is a further object to provide such a catalyst
022 system useful in a single reactor under hydroprocesslng condi-
023 tions without the need for separation of reaction products
024 between catalyst stages. It is a further object to prov1de
025 such a catalyst system having enhanced hydrocracking, hydrodeni-
0~6 trification, and~'or hydrodesulfuri3ation activity. Another
027 object is to provide a two-stage catalyst system having
028 enhanced fouling resistance for hydrocracking, hydrodenitrifica-
029 tion and/or hydrodesulfurization relative to elther of the
030 catalyst stages alone. A further object is to provide such a
031 catalyst configuration which ls capable of providing feedstock
032 of reduced metals, sulfur, and nitrogen content for subsequent
033 fluid catalytic cracking.
034 These and other objects are provided according to the
035 present invention in a process for hydroprocessing a heavy
036 hydrocarbon feedstock comprising the steps of:
037 (a) contacting said hydrocarbon feedstock with hydrogen
--3--
under hydroprocessing conditions in the presence of a Eirs-t
hydroprocessing catalyst comprising refractory suppor-t material
and at least one metal, metal oxide or metal sulfide of groups
VIb and VIII elements, said first hydroprocessing catalyst
having an average pore diameter in -the range of 60-lS0 ~;
(b) contacting at least a por-tion of -the hydrocarbon pro-
duct from said step (a) under hydroprocessing conditions with a
second hydroprocessing catalyst comprising refractory support
material and at least one metal, metal oxide or metal sulfide
of groups VIb and VIII elements, said second hydroprocessing
catalyst having an average pore diameter of 30-70 ~ and smaller
than the average pore diameter of said first hydroprocessing
catalyst, said first and second hydroprocessing catalysts con-
stituting a synergistic hydroprocessing combination wherein said
hydroprocessing conditions include a temperature at which at
least one of the hydrocracking activity, the hydrodesulfuriza-
tion activity, and the hydrodenitrification activity of said
synergistic hydroprocessing combination exceeds the correspond-
ing activity of said first hydroprocessing catalyst and said
second hydroprocessing catalyst alone. Preferably the first
stage catalyst has at least 40% and more preferably 50% pore
volume in pores >80~. The second stage catalyst preferably has
at least 50% and more preferably at least 90% pore volume present
as pores having diameters smaller than 80A. The feed to the
second contacting step can be the entire liquid hydrocarbon
product of the first contacting step, or the entire product of
the first contacting step, including reaction products of hydro-
desulfurization (e.g. H2S) or hydrodenitrification (e.g. NH3).
The process of this invention is particularly applicable to
-3a-
hydroprocessing heavy hydrocarbon feedstocks containing above
about 1 weight percent sulfur and abou-t 0.1 weight percent
nitrogen. The first catalyst refractory suppor-t material
preferably consists essentially of alumina, the second catalyst
refractory support material preferably consists essen-tially of
alumina and 10 to 70 weight % silica. It is preferred tha-t the
hydrocarbonaceous feed contain less than about 5 weight %
asphaltenes such as a deasphalted atmospheric residuum, deasphal-
ted vacuum residuum, vacuum gas oil, or mixtures thereof.
FIG. 1 is a graphical representation of the hydro-
cracking activity of the catalyst combination of the invention
compared to single catalysts.
7~
001 -4~
002 FIG. 2 is a graphical representation of the hydrode-
003 sulfurization activity of the catalyst combination of this
004 invention compared to single catalysts.
005 FIG. 3 is a graphical representation of the hydrode-
006 nitrification activities of the catalyst combination of this
007 invention compared to single catalysts.
008 DETAILED DESCRIPTION OF THE INVENTION
009 According to this invention, the heavy hydrocarbon
010 feedstock is contacted under hydrprocessing conditions with at
011 least two catalyst beds to cause hydrodesulfurization, hydro-
012 denitrification and/or hydrocracking of the feedstock. The
013 precise hydroprocessing conditions depend primarily upon the
014 extent of reaction needed. Hydroprocessing conditions include
015 temperatures in the range of 250 to 600C, preferably 350 to
016 500~C and most preferably 400 to 450C; total pressures in the
017 range of 30 to Z00 atmospheres, preferably 100 to 170 atmo-
018 spheres and more preferably 120 to 150 atmosph`eres; hydrogen
019 partial pressures in the range of 25 to 190 atmospheres, prefer-
020 ably 90 to 160 atmospheres and most preferably 110 to 140 atmo-
021 spheres, and space velocities (LHSV) of 0.1 to 10, preEerably
022 0.3 to 5 and most preferably 0.5 to 3 hours 1, The catalyst of
023 each stage is comprised of a refractory ceramic support
024 material such as alumina, silica, magnesia, zirconia, or
025 mixtures thereof. The catalysts contain as a hydrogenation com-
02.6 ponent one or more metals, metal oxides or metal sulfides,
027 selected from elements of group VIb and group VIII of the
028 Periodic Table of the Elements as set forth in Handbook of
029 Chemistry and Physics, 45th Ed~ Chesnical Rubber Company,
030 Cleveland, Ohio, 1964. It is preferred that each catalyst
031 contain at least one metal, metal oxide, or metal sulfide from
03~ group VIb and one metal, metal oxide or metal sulfide from
033 group VIII, for example, Co/~o, Ni/~o, Ni/~, etc. The metals
034 should typically be present in quantities of 5 to 25 weight %
035 of group VIb and 1 to 20 weight ~ of group VIII, as metals,
036 based upon total weight of catalyst, as is typical for hydropro-
037 cessing catalysts. Promoters, such as phosphorus or titanium
001 -5-
002 as metals, oxides or sulfides, can be added, if desired. The
003 hydrogenation and promoter metals or metal compounds can be
004 included in the catalyst by any of the well known methods of im~
005 pregnation of a refractory shaped support, coprecipitating
006 comulling, cogelling, etc.
007 The catalyst supports of the first and second stage
008 differ in the pore size distribution. Catalyst for the fi.rst
009 bed, i.e., the bed which first encounters the heavy hydrocarbon
010 feedstock, has an average pore diameter within the range of
011 60-150, preferably 80-120A. The pore volume distribution is
012 such that at least 40%, preferably at least ~5%, and most
013 preferably at least 50~ of its pore volume is present in pores
014 having diameters larger than 80A. The catalyst for the second
015 bed is characterized by an average pore diameter within the
01~ range of 30-70, and preferably 40-60A. The pore volume distri~
017 bution is such that at least 50%, preferably at least 75~ and
018 most preferably at least 90% of its pore volume is present in
019 pores smaller than 80A in diameter. It has been found that a
020 combination of hydroprocessing catalysts having such pore
021 structures illustrates excellent hydrocracking, hydrodenitri-
022 fication and hydrodesulfuri~ation activities as well as
023 enhanced fouling resistance relative to either catalyst alone.
024 Consequently, the two-bed catalyst configuration of this inven-
025 tion can provide synergistic hydrodenitrification, hydrodesul-
026 furization, and/or hydrocracking activities.
027 When heavy feeds are hydroprocessed in order to pro-
028 duce lighter components, i.e~ hydrocracking, the second stage
029 catàlyst should have higher acidity, hence increased hydro-
030 cracking activity, than the ~irst stage catalyst. The first
031 stage catalyst reduces the nitrogen content of the feedstock
03~ before it contacts the second stage catalyst thereby preserving
033 the acidity of the second stage. For example, the first stage
034 catalyst support material can be alumina and the second stage
035 catalyst support material can be alumina containing 10 to 70
036 weight % silica, more preferably ~0-60 weight ~ silica.
037 Activated forms of alumina such as beta, gamma, etc. can be
001 -6-
002 used in either stage if desired. Either catalyst can contain
003 zeolitic components; however, little improvement in hydro-
004 cracking is obtained unless temperatures above about 430C are
OOS used. Consequently, the process oE this invention operates
006 satisfactorily when one or both the first and second stage
007 catalysts are free of zeolitic components.
008 The hydroprocessing conditions of the first and
009 second catalyst beds can be the same or different. For
010 particularly heavy feedstocks, hydrogenation conditions should
011 be more severe in the first catalyst bed. It is believed that
012 the first catalyst bed results in some hydrocracking of heavy
013 materials to molecules more able to diffuse into the second
014 stage catalyst pores.
015 The first and second catalyst stages can be operated
016 as fluidized beds, moving beds, or fixed beds. When both beds
017 are operated as fixed beds, they can be disposed in fluid
01~ communication in a single reactor or reaction zone. No other
019 group VIb or VIII metal-containing catalytic material need be
020 present between the two catalyst stages, e.g. the stages can be
021 unseparated or separated only b~ porous support material or
022 reactor internals. It may be desirable, however, to include
023 inexpensive support catalysts between the beds, such as alumina
024 impregnated with less than 10 weight percent total metals, as
025 metals.
0~6 The catalysts in the fixed beds can be irregular
027 particulates or any of the other conven~ional catalyst shapes
028 and sizes. The catalysts are preferably in the form of
029 extrudate, spheres, pellets, trilobes, etc. having diameters of
030 1/8 inch or less.
031 In order to preserve the catalytic activity of the
032 catalyst beds, the feedstock entering the first catalyst bed
033 should contain no more than about 25 ppmw, total V, Ni, and Fe
034 as metals. Ine~pensive guard catalysts such as red mud, etc.
035 can be employed to demetalize the feed to the required metals`
036 level. Due to the small pore size of the first stage catalyst,
037 the feedstock should be substantially free of large asphaltene
molecules. The feed preferably contains less than 5 weight % asphaltenes,
more preferably less than 0.02 weight % asphaltenes. Asphaltenes are defined
as hydrocarbonaceous materials which are soluble in benzene but not in n-}leptane.
~xamples of such low asphaltene feedstocks are solvent (e.g. liquefied propane)
deasphalted atmospheric or vacuum residual oils, and vacuum gas oils, etc. from
the fractionation of crude oil, shale oil, oil from tar sand, dissolved coal
or other coal liquefaction products. The relative amounts of first and second
stage catalyst in the process can range from 1:10 to 10:1, depellding upon the
feedstock. Feedstocks higher in metals and heavy components generally require
a greater proportion of first stage catalyst.
Catalysts useful in the first and second stages according to this
invention can be obtained commercially, either as supports which can be
impregnated or as catalysts containing the desired level of metals, metal
oxides, or metal sulfides. Catalysts suitable for use in the first stage and
second stage reactors can be prepared in the following manner. All percentages
are by weight.
First Stage Catalyst Preparation
Hydrated Kaiser alumina is peptized with concentrated IINO3. The
resulting solution is back-titrated with concentrated Nl~ OH to a pll of 5. The
mixture is extruded at a consistency of about 50% volatiles. The extrudate is
surface-dried at 120C for 2 llours and at 200C for 2 hours. The dried
extrudates are calcined at 750C for one hour in dry air. An impregnant is
prepared by mixing 50 ml of crude phosphomolybdic acid (a mixture of 20 parts
MoO3, 2 parts H3PO4 and 48 parts water and containing 2.5 weight % P, and 20.1
weight % Mo) with 0.8 ml of 85% 113PO4, followed by heating to 45C. 7 ml of
an aqueous NiCO3, containing 62.8% NiO is added. After the solution becomes
clear, it is cooled to 25C and diluted wilh water to a volume of 55cc per 100
grams of catalyst to be produced. The resulting impregnant solution is sprayed
onto the extrudate under vacuum. The sprayed extrudate is allowed
7~
to stand at room temperature for one hour and is then surface-dried at 120C
for one hour. The dried catalyst is then calcined in 570 liters/hr of dry
air for 6 hours at 95C, 4 hours at 230C, ~ hours at 400C and 4 hours at
510 C and allowed to coo]. The pore diameter can be varied, if desired, by
conventional techniques, such as varying the back titration of the extrusion
mix as described in United States 4,082,697.
Second Sta~e Catalyst Preparation (100 gram basis)
A first solution is prepared by combining 407cc of H20, 66 grams
of a 21.5% aqueous AlC13 solution, 57 grams of 30.7% aqueous NiC12 solution
and 369 grams of an aqueous solution containing 5.15% TiC14, 81.3% AlC13,
and 13.5% acetic acid. The heat of solvation produces a final temperature
of about 40C for the first solution. A second solution is prepared by
mixing 94 grams of a 28.7% aqueous SiO2 with 283cc H20. The second solution
is added slowly to the first solution, with mixing. The pH is increased to
4.5 with aqueous N~140H (8.0%) whereupon a gel is formed. After a pH of 4.5
is obtained, 281 grams aqueous ammonium paratungstate [(N~l4)6W7024-6H20]
solution, containing 6~97%W, is added and the pH is raised to 7.5 by adding
8.0% NH40H, whereupon gellation is completed and coprecipitation occurs.
IE an aluminosilicate zeolite (e.g. ultrastable Y-type) component is
desired, the necessary amount of a 25% aqueous zeolite slurry is added.
The catalyst gel is aged at 75 C Eor one hour and filtered. The filter
cake is surface-dried and twice extruded. The extrudates are washed and
calcined in air for 4 hours at 200C and 5 hours at 510C. Those familiar
with the art of preparing cogelled catalyst supports can vary the pore
size distribution, if desired, using conventional techniques such as adding
a detergent as described in IJnited States Patent 3,657,151.
EXPERI~ENTAL
Several catalyst configurations were tested in a fixed bed pilot
plant reactor. The feed was a mixture of Alaskan North Slope solvent
deasphalted vacuum residuum and
27~
001 -9-
002 vacuum gas oil (3:2 vol. ratio). The properties of the feed
003 are shown in Table I.
004 TABLE I
006 Specific Gravity 0.95
007 S, wt. % 1.6
008 N, ~t. ~ 0.4
009 Ni/V/Fe, ppmw 8/7/4
010 Ramsbottom carbon, wt. ~ 3.0
011 Distillation, ASTM D1160
012 Start 258C
013 10 401
014 30 471
015 50 521
016 70 585
017 End Point ~90
018 % Recovery 73
019 The feed was passed through the pilot plant reactor
020 containing fixed beds of catalystsO The liquid hourly space
021 velocity was 1.0, total pressure was 135 atmospheres and
022 hydrogen pressure was 110 atmospheres. The temperature was
023 varied in order to measure the reaction constant. The hydro~
024 processing reactor atmosphere was provided by recycle gas from
02S the reactor at 890 cubic meters per cubic meter of feedstock.
026 The properties of catalysts tested are shown in Table
027 II. Catalyst C contained an ultrastable Y zeolite component.
' 01 ~o
--I ~ 't' ~r ~1
--~ ~ U~ 9
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O
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3 ~ X ::~
u~ ~
O O O O O O O O
O O O O O C:) O O O O O O C~ O O
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0 0 1
002 Table 3 sets forth the pore size distribution of the
003 catalysts. The pore distribution was obtained by nitrogen
004 adsorption technique, using a Digisorb 2500, Micrometrics
005 Instrument Corporation. The average pore diameter, as used
006 herein, is obtained by dividing the measured pore volume, in
007 cc/gm by the measured surface area, in m2/gm and multiplying
008 the result by 40,000.
009 TABLE 3
011 Catalyst A B C D E
012
013
014 Pore Volume, cc/gm .40 .38 .40 .49 .40
015 Avg. pore O
016 diameter, A 10745 48 86 60
017 ~ volume present
018 as pores having
019 diameters, A
020 <30 ~
021 30-60 10appr.10098 15 85
022 60-80 40 - appr.2 50 15
023 80-90 25 - - 20
024 90-100 15 - - appr.12
025 100-150 10 - - apprO 3
~26
027
028 The catalysts were tested individually and in beds con-
029 taining various combinations of two catalysts in equal volumes,
030 with a layer of the first catalyst directly over a layer of the
031 second catalyst. Reaction constants KHCR (hydrocracklng), KHDS
032 (hydrodesulfurization) and KHDN (hydrodenitrification) were
033 computed for the plant runs. These reaction constants are
034 plotted as a func~ion of catalyst temperature for single cata-
035 lyst and 50/50 volume catalyst mixtures in Figures 1, 2 and 3.
036 ~HCR is equal to LHSV ln l-Xo
037 l-x
OOl ~12
00~ where
003 x = liquid volume percent oE product boiling below 343C
004 and XO = liquid volume percent of feed material boiling
005 below 343C.
006 KHDS = (LEISV) 7 ln S~ where Sf = % sul~ur in the feeds~)c~
007 Sp
008 and Sp = ~ sulfur in the product
009 KHDN = LHSV ln Nf where Nf = ~ nitrogen in the feedstock
010 Np
011 and Np = % nitrogen in the product.
012 Because the ordinate has a logarithmic scale, the slope of the
013 straight lines is equal to dlog K/dT, which is a measure o~ the
014 activation energy or the reaction. FIGS. 1, 2 and 3 demon-
015 strate that the combinations of catalysts A and B and A and C
OlG are synergistic in that the combinations have greater activa-
017 tion energies (greater slopes) than either of the catalysts
018 alone. Because the reaction constants for the combined cata-
019 lysts increase more rapidly with temperature than do the
020 -- reaction constants of either catalyst alone, there will
021 necessarily be a temperature above which the combination cata-
022 lyst is more active for the particular reaction than either
023 catalyst alone. This corresponds to the intersection of the
024 appropriate lines connecting the calculated K values. Conse-
025 quently, catalyst combinations for which the slope of ln K vsO
026 T, for at least one of hydrocracklng, hydrodesulfurization, and
027 hydrodenitrification, is greater than for either catalyst
028 component alone are defined as synergistic hydroprocessing
029 combinations. FIG. 1 demonstrates that at temperatures above
030 about 410C, the combination catalysts A/B and A/C have greater
031 hydrocracking activity than catalysts A or B alone. Catalyst
032 D/C would demonstrate greater activity than catalyst D above
033 about 427C. FIG. 2 shows that above about 415C and 421C
034 catalysts A/B and A/C, respectively, have greater desulfuriza-
035 tion activity than either catalyst A or B. Catalyst D/C, on
036 the other hand, will not surpass catalyst D in
037 hydrodesulfurization temperature until much higher temperatures
038 are employed.
001 -13-
002 FIG. 3 shows that above about 410C and 416C, catalysts A/3
003 and ~/C have higher hydrodenitrification activitles than either
004 catalysts A or B alone. Again, nnuch higher temperatures will
005 be required before catalyst D/C becomes more active than
006 catalyst D. In order to best achieve the objects o~ this
007 invention, the catalyst components of the two-stage catalyst
008 should be selected so that enhanced hydrocracking, hydro-
009 desulfurization, and/or hydrodenitrification activities are
010 achieved at the desired hydroprocessing temperature, e.g. most011 preferably in the 350 to 500C range. Table 4 sets forth the
012 product distribution and the precise reaction conditions of the
013 various runs. As seen, the combination of catalysts A and B
014 and C and D produce significantly higher naphtha (C5-205C) and
015 diesel (205 to 343C) fractions than do the individual catalyst
016 components.
017 Additional pilot plant runs were made to determine the
018 fouling rate of the combination bed of this invention relatlve
019 to larger pored catalyst, which would ordinarily be expected to
020 be the more resistant to fouling. The feedstock was an Alaskan
021 North Slope deasphalted oil having the composition set ~orth in
022 Table 5. In each case the feed was passed downwardly through a
023 bed having an impregnated A1203 guard catalyst compeising 40%
024 of the bed volume and situated above the catalyst sample
025 tested. The guard catalyst was a commercially low density
02~ catalyst containing about 2~ Co and 4~ ~lo present as oxides,
027 and having a pore volume of 0.67 cc/gm and an average pore
028 diameter of 80-lOOA.
7~ `
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er O ~ J O O O O O ~ ~ t`l
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coO ~ ~ ~ ~ JJ oo o o o
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e ~ E ~ c ~ (~
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O O O O O O ~ O .~ ~ .~ ~1 ~ ~ .~ ~ ~ ~ ~ ~ ~ ~ ~
o o o o o o o o o o o o o ~ o o o o o o o o o
001 -15-
002 TABLE 5
003
004 Specific Gravity .955
005 S, wt. ~ 1.4
006 N, wt. % 0.5
007 Ni/V/Fe, ppmw 10/5/3
008 Ramsbottom carbon, wt. ~ 3.8
010 Distillation, ASTM D1160
011 Start 385C
012 5 442
013 10 491
014 30 547
015 50 581
016 End Point 591
017 % Recovery 55
018
019 Table 6 sets forth the results of these fouling
020 tests. The hydrocracking fouling of the combination catalyst
021 charge proceeded at about the same rate as the single charge,
022 while the hydrodesulurization and hydrodenitrification fouling
023 rates of the combination catalyst A/B were approximately half
024 the ouling rates of catalyst A alone.
7~
001 -16-
002 TABLE 6
004 FO~LING TEST
006 427C
007 1.0 LHSV
008 105 atmospheres H2
011 Catalyst Charge, Vol. % 40~ Guard A12O3 40% Guard A12O3
012 60% A 30~ A
013 , 30~ B
015 ~HCR (Hr
016 SORl0.25, 22 LV% <343C 0.25, 22 LV% <343C
017 1600 Hr0.19, 17 LV% <343C 0.20, 18 LV~ <343C
018 HCR Fouling Rate 1C/Hr) 0.003 0.003
020 RHDS ~Hr
021 SORl 4.5, 150 ppm S 4.5, 150 ppm S
022 1600 Hr 3.2, 560 ppm S 3.7, 340 ppm S
024 HDS Fouling Rate (C/Hr) 0.0077 0.0044
.
026 KHDN (Hr
027 SORl 1.9, 700 ppm N 1.8, 780 ppm N
028 1600 Hr 1.0, 1730 ppm N 1.3, 1280 ppm N
030 HDN Fouling Rate (F/Hr) 0.023 0.011
oo323 1 Start of Run
~8~7~i~
001 -17-
002 DESCRIPTION OF T~E PREFERRED CONFIG[~RATION
003 A solvent-deasphalted vacuum gas oil having character-
004 istics as shown in Table 5 is introduced with a hydrogen-con-
005 taining gas into the upper portion of a downflow, fixed bed
006 catalytic reactor having at least 3 layers of catalyst
007 material The first, or upper, layer is a bed of guard cata-
008 lyst such as alumina particles about 5 mm. in diameter having a
009 pore volume of about 0.7 cc/gram and an average pore diameter
010 of about 80 to 100 A. The second catalyst layer is comprised
011 of 2.5 mm~ particles of alumina impregnated with nickel,
012 molybdenum, and phosphorous compounds and calcined to provide a
013 catalyst containing about 2-5 wt. % Ni as NiO, 8-15 wt. ~ Mo as
014 MoO3 and 1-4 wt. ~ P as P2O5. The catalyst of the second layer
015 has a pore volume of about 0.4 cc/gram, an average pore
016 diameter of 80-120 A, and at least 50% pore volume in pores of
017 80 to 150 A diameter. The third catalyst layer is comprised of
018 1/10 inch particles of cogelled SiO2/A12O3 particles having a
019 SiO2/A12O3 ratio of about one-to-one and containing about O-9
020 wt. % Ni as NiO~ 4-25 wt.~ W--as WO3 and about 4 wt. ~ Ti as
021 TiO2. The third stage catalyst has a pore volume of about 0.4
022 cc/g and at least 90% pore volume in pores of 30 to 80 A. The
Q23 first catalyst occupies about 40~ of the volume of the beds in
024 the reactor. The second and third beds are of equal volumeO
025 Additional catalys'cs such as alumina containing no more than
026 about 5-10% group VIb or VIII metals as metals can be used as
027 support catalysts between the beds or elsewhere in the reactor.
028 The reactor is operated at a liquid hourly space velocity,
029 based on the second and third bed volumes of 1.7. The total
030 pressure is 140 atmospheres with a 100-atmosphere E12 pressure.
031 The H2 flow rate is 140,000 liters/min. and the reactor
032 temperature is 425C. The product leaves the reactor below the
033 third catalyst layer and is fractionated at atmospheric
034 pressure. H2 is recovered from the vapor fracticn and recycled
035 to the reactor. Intermediate cuts of naphtha (C5-200C) and `
036 diesel (200-350C), are recovered. The 350C+ bottom fraction,
037 having significantly reduced nitrogen and sulfur contents, is
038 passed as feed to a conventional fluld catalytic cracking unit.
7~
0 0 1
002 The examples and embodiments herein are provided to
003 illustrate the invention and are not intended to b~ exhaustive
004 or limiting, the scope of the invention being limited only by
005 the claimsO
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