US3594309A - Conversion and desulfurization of hydrocarbonaceous black oils - Google Patents

Conversion and desulfurization of hydrocarbonaceous black oils Download PDF

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US3594309A
US3594309A US771248A US3594309DA US3594309A US 3594309 A US3594309 A US 3594309A US 771248 A US771248 A US 771248A US 3594309D A US3594309D A US 3594309DA US 3594309 A US3594309 A US 3594309A
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boiling
temperature
reaction zone
desulfurization
catalytic
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Frank Stolfa
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Universal Oil Products Co
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Universal Oil Products Co
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    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G69/00—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one other conversion process
    • C10G69/02—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one other conversion process plural serial stages only
    • C10G69/06—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one other conversion process plural serial stages only including at least one step of thermal cracking in the absence of hydrogen
    • 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
    • C10G49/00—Treatment of hydrocarbon oils, in the presence of hydrogen or hydrogen-generating compounds, not provided for in a single one of groups C10G45/02, C10G45/32, C10G45/44, C10G45/58 or C10G47/00
    • C10G49/007—Treatment of hydrocarbon oils, in the presence of hydrogen or hydrogen-generating compounds, not provided for in a single one of groups C10G45/02, C10G45/32, C10G45/44, C10G45/58 or C10G47/00 in the presence of hydrogen from a special source or of a special composition or having been purified by a special treatment

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  • ABSTRACT OF THE DISCLOSURE A process for converting sulfurous, hydrocarbonaceous black oils into lower-boiling, normally liquid hydrocarbon products of reduced sulfur content.
  • the process involves the integration of hydrogenative cracking and fixedbed catalytic desulfurization, and is especially applicable to those hydrocarbon charge stocks containing less than 150 ppm. of metallic contaminants.
  • the charge stock is initially subjected to fixed-bed catalytic hydrogenation and desulfurization.
  • a high-boiling concentrate is thermally-cracked in the presence of dissolved hydrogen.
  • the process described herein is adaptable to the desulfurization of petroleum crude oil residuals having relatively low metals content-ie. containing less than about 150 p.p.m. of total metals. More specifically, the present invention is directed toward a combination process for converting and reducing the sulfur concentration of hydrocarbonaceous charge stocks commonly referred to in the art as black oils.
  • the various black oil charge stocks can be classified as (l) high metals residuals or (2) low metals residuals.
  • the present invention is primarily directed to the processing of those hydrocarbonaceous black oils having a low metals content of less than about 150 p.p.m., computed as if existing in the elemental state.
  • a Iblack oil is generally characterized as a heavy lcarbonaceous material of which more than about 10.0% by volume has a normal boiling point above a temperature of 1050 F. (referred to as nondistillables). Such material generally has a gravity less than about 20.0 API and sulfur concentrations greater than about 2.0% by weight. Conradson carbon residue factors exceed 1.0% by Weight, and a great proportion of black oils indicate a Conradson carbon residue factor above 10.0.
  • hydrocarbonaceous black oils to the conversion and desulfurization of which the present invention is directed, include a crude tower bottoms product having a gravity of about 14.3 API, and contaminated by the presence of about 3.0% by weight of sulfur, 3830 p pm. of total nitrogen, S5 ppm. of total metals and about 11.0% by weight of asphaltenes.
  • Another typical charge stock is a vacuum column bottoms product derived from a Middle-East crude oil. This vacuum bottoms product has a gravity of ⁇ 60 API, an average molecular weight of about 620, an ASTM 20.0% volumetric distillation temperature of about 1035 F., and contains about 4000 p.p.m. of nitrogen, 5.5% by ice weight of sulfur, p.p.m.
  • the present invention affords the conversion of such material into lower-boiling, normally liquid hydrocarbon products, and further converts a considerable quantity of nondistillables. Additionally, the normally liquid product of the process has been substantially reduced in sulfur content-ie. less than about 1.0% by weight of sulfur.
  • the present invention is founded on recognition of the fact that an acceptable degree of desulfurization of low metals-containing black oils is possible at relatively mild operating severities which favor extended catalyst life without effecting a significant degree of asphaltene polymerization.
  • an essential feature of my invention resides in the subsequent processing of the liquid product efiiuent from the fixed-bed catalytic reaction zone. Therefore, as hereinafter set forth in greater detail, the reaction zone effluent is separated at a temperature of from about 100 F.
  • a principal object of my invention is to provide an economical process for effecting the desulfurization and hydrogenative conversion of low metal-containing black oils.
  • a corollary objective is to extend the period of acceptable, economical catalyst life while desulfurizing and converting hydrocarbonaceous black oils containing less than about p.p.m. of total metals.
  • Another object is to convert heavy hydrocarbon charge stocks, a significant amount of which exhibits a boiling range above a temperature of 1050 F., into lower-boiling distillable hydrocarbons having a sulfur concentration less than about 1.0% by weight.
  • my invention relates to a process for the con-version of a sulfurous, hydrocarbonaceous charge stock, of which at least about 10.0% boils above a temperature of about 1050 EF., into lower-boiling hydrocarbon products, which process comprises the steps of: (a) heating said charge stock to a temperature of from 500 F.
  • the total charge to the iirst iixed-bed catalytic reaction zone consisting primarily of fresh charge stock, a recycled portion of the rst liquid phase, a recycled hydrogen-rich vapor phase and make-up hydrogen, required to supplant that which is consumed within the overall process and to maintain pressure, is heated to a temperature within the preferred range of from about 650 F. to about 750 F
  • the precise temperature is controlled within the aforesaid range by monitoring the temperature of the reaction zone product effluent.
  • the iirst reaction zone effluent being introduced into the irst separation zone is at a temeprature of from about 700 F. to about 800 F. in order that the portion of the first liquid phase being subjected to the subsequent thermal cracking reaction zone, or coil, contains from about 10.0 mol. percent to about 40.0% of dissolved hydrogen.
  • the principal function of the present invention resides in the production of maximum quantities of distillable hydrocarbons which have been reduced in sulfur concentration. Through the utilization of the present process, this is accomplished in a highly economical fashion while avoiding the difficulties and pitfalls of currently-practiced processing techniques. Paramount is the extension of the period of time during which the fixed-bed catalytic composite functions in an acceptable manner. With respect to the processing of high metals black oils, being those containing more than about '150 p.p.m. of total metals, it has been found that a successful operation involves initially hydrovisbreaking the fresh hydrocarbon charge stock in the presence of limited quantities of hydrogen.
  • the residual charge stock is catalytically desulfurized, and at least partially converted, at relatively mild hydrogenation severities which favor extended catalyst life.
  • the catalytically converted product effluent is separated into a principally vaporous phase and a principally liquid phase, at least a portion of the latter being utilized as the charge to a non-catalytic thermal cracking reaction zone.
  • the total charge to the xed-bed catalytic reaction zone includes the fresh hydrocarbon charge stock, a recycled hydrogen-rich gaseous phase, make-up hydrogen and a recycled diluent, the source of the latter being hereinafter set forth.
  • This mixture is raised to a temperature of from about 500 F. to about 750 F., and preferably from about 650 F. to about 750 F. as measured at the inlet to the catalyst bed.
  • the inlet temperature is controlled at a level such that the temperature of the reaction product effluent, or the maximum catalyst bed temperature does not exceed about 800 F.
  • the reaction zone will be maintained under an imposed pressure of from about 1000 to about 4000 p.s.i.g., and the hydrocarbon charge stock will contact the catalytic composite at a liquid hourly space velocity of from about 0.5 to about 10.0, based upon the fresh hydrocarbon charge stock, exclusive of any recycled diluent.
  • the hydrogen concentration will be in the range of from about 5000 to about 50,000 standard cubic feet per barrel, while the combined feed ratio, defined as total volumes of liquid charge per volume of fresh hydrocarbon charge, is in the range of from about 1.1:1 to about 3.5:1.
  • the catalytic composite disposed within the fixed-bed catalytic reaction, or conversion zone can be characterized as containing a metallic component having hydrogenation activity, which component is combined with a suitable refractory inorganic oxide carrier material of either synthetic, or natural origin.
  • a suitable refractory inorganic oxide carrier material of either synthetic, or natural origin.
  • the precise composition and method of manufacturing the carrier material is not considered essential to the present invention, although a siliceous carrier, such as 88.0% by weight of alumina and 12.0% by weight of silica, or 63.0% by weight of alumina and 37.0% by weight of silica, or 68.0% by weight of alumina, 10.0% by weight of silica and 22.0% by weight of boron phosphate are generally preferred.
  • Suitable metallic components having hydrogenation activity are those selected from the ⁇ group consisting of the metals of Group VI-B and VIII of the Periodic Table, as set forth in the Periodic Table of the Elements, E. H. Sargent & Company, 1964.
  • the catalytic composites may comprise one or more metallic components from the group of molybdenum, tungsten, chromium, iron, cobalt, nickel, platinum, iridium, osmium, rhodium, ruthenium, and mixtures thereof.
  • concentration of the catalytically active metallic component, or components is primarily dependent upon a particular metal as well as the physical and/ or chemical characteristics of the charge stock.
  • the metallic components of Group VI-B are generally present in an amount within the range of from about 1.0% to about 20.0% by weight, the iron-group metals in an amount within the range of about 0.2% to about 10.0% by weight, whereas the noble metals of Group VIII are preferably present in an amount within the range of from about 0.1% to about 5.0% by weight, all of which are calculated as if these components existed within the catalytic composite in the elemental state.
  • the catalytic first reaction zone is maintained at a pressure of about 2900 p.s.i.g., the yfirst separation zone, or hot separator will function at about 2780 p..i.g.
  • the distillable portion of the charge stock comprises those hydrocarbons which can be distilled at a temperature lbelow about 1050 F.
  • the charge stock contains asphaltics, this temperature becomes limiting in order to prevent cracking.
  • analyses have indicated end boiling points as high at 1100 F. to 1150 F.
  • the total product effluent from the first catalytic reaction zone is passed into a first separation zone hereinafter referred to as the hot separator.
  • the principal function served by the hot separator is to separate the mixed-phase product effluent into a principally Vapor phase rich in hydrogen and a principally liquid phase containing from about 10.0 mol percent to about 40.0 mol percent of dissolved hydrogen.
  • the total reaction product effluent is utilized as a heat-exchange medium in order to lower the temperature thereof to a level in the range of from about 700 F. to about 800 F., and preferably below a level of 750 F.
  • the principally vaporous phase from the hot separator is introduced into a second separation zone hereinafter referred to as the cold separator.
  • the cold separator operating at substantially the same pressure as the hot separator, but at a significantly lower temperature in the range of about 60 F. to about 140 F., serves to concentrate the hydrogen in a second principally vaporous phase.
  • This hydrogenrich vaporous phase comprising about 80.0 mol percent hydrogen, and only about 2.2 mol percent propane and heavier hydrocarbons, is made available for use as a recycle stream to be combined with the fresh black oil charge stock. Butanes and heavier hydrocarbons are condensed in the cold separator, and removed therefrom as a second principally liquid phase.
  • the liquid phase from the hot separator is in part recycled to combine with the fresh hydrocarbon charge stock to serve as a diluent for the heavier constituents thereof.
  • the quantity of the liquid phase diverted in this manner is such that the combined feed ratio to the catalytic reaction Zone, being defined as total Volumes of liquid charge per volume of fresh liquid charge, is within the range of from about 1.1:1 to about 3.5: 1.
  • the remaining portion of the principally liquid phase from the hot separator is introduced into a thermal cracking reaction zone, or coil, at a reduced pressure in the range of from about 200 p.s.i.g. to about 500 p.s.i.g. and at a temperature of from about 700 to about 750 F.
  • the thermally-cracked product effluent is introduced into a fractionation zone, or distillation column maintained at conditions of temperature and pressure such that light, normally gaseous hydrocarbons and gasoline boiling range hydrocarbons, having a nominal end boiling point of about 400 F., are removed as an overhead fraction, middle-distillate hydrocarbons boiling from about 400 F. to about 650 F. are removed as a side-cut fraction, and material boiling above about 650 F. are removed as a bottoms fraction.
  • the second principally liquid phase from the cold separator is also introduced into the distillation column, and at a locus intermediate that from which the middle-distillates are withdrawn, and that at which the thermally-cracked product effluent is introduced.
  • the 650 F. plus hydrocarbonaceous material, withdrawn from the bottom portion of the distillation column, is introduced into a vacuum flash column maintained at about 20 to about 60 mm. Hg absolute.
  • the vacuum flash zone serves as the third separation zone, the principal function of which is the concentration and recovery of an asphaltic residuum substantially free from distillable hydrocarbons.
  • vacuum gas oil streams are recovered from the vacuum flash column as a separate light vacuum gas oil (LVGO) and a heavy vacuum -gas oil (HVGO), although in some instances, a medium vacuum gas oil is also recovered.
  • LVGO light vacuum gas oil
  • HVGO heavy vacuum -gas oil
  • a slop-wax stream boiling within a temperature range of about 980 F. to about 1150 F., is withdrawn from the vacuum flash column, and at least a part recycled to combine with the fresh hydrocarbon charge stock in an amount of from 1.0% to 20.0%.
  • the amount of the slop-wax is such that the combined feed ratio to the catalytic reaction zone, including the recycled portion of the first liquid phase from the hot separator, continues to be maintained in the range of from about 1.1:1 to about 3.5 :1, as hereinafter set forth.
  • a portion of the slop-wax and/or heavy vacuum gas oil may be recycled to combine with the first liquid phase being introduced to the thermal cracking reaction zone. The amount so recycled is such that the combined feed ratio to the thermal cracking reaction coil is above about 1.2:l, and generally not higher than about 4.011.
  • the principal advantage, or benefit, attendant the use of my invention resides in an extension of the period of acceptable catalyst life with respect to a fixed-bed catalytic reaction zone. This stems primarily from the fact that desulfurization, to a level less than about 1.0% by weight, is effected at a relatively low severity of operation with the result that the atmosphere within the reaction zone is not conducive to the formation of polymer products, containing sulfur, Otherwise resulting from the presence of hydrocarbon-insoluble asphaltenes.
  • Another advantage resulting from the low severity of operation within the fixed-bed catalytic reaction zone involves the dissolution of hydrogen in the normally liquid heavier portion of the product effluent, at least a portion of which is utilized as the charge to the thermal cracking reaction zone.
  • the vacuum flash column is significantly reduced in size as a result of the fractionation zone being utilized to separate the thermal reaction zone product effluent to concentrate the material boiling above a temperature of about 650 F.
  • This affords an added advantage with respect to the overall economics of the process.
  • the introduction of the second principally liquid phase from the cold separator, into the distillation column at a point intermediate the withdrawal of the middledistillates and the introduction of the thermally-cracked product effluent, further insures that the bottoms stream contains a minimum quantity of material boiling below 650 F. Since the hydrocarbonaceous material boiling below about 650 F. is separately recovered in the distillation zone, there is effected a beneficial reduction in the size of the vacuum flash column.
  • the drawing will be described in connection with the conversion and desulfurization of a vacuum column bottoms product derived from a full boiling range crude oil.
  • the vacuum bottoms product has a gravity of 11.7 API, an average molecular weight of about 600, an ASTM initial boiling point of about 860 F. with about 44.0% by volume being distillable at a temperature of 1050 F., contains 4600 p.p.m. of nitrogen, 1.92% by weight of sulfur, 105 p.p.m. of vanadium and nickel, has a Conradson carbon residue factor of 12.8% by weight and contains about 3.45% by weight of heptaneinsoluble asphaltenes.
  • This vacuum column bottoms is intended for conver- -sion to maximum distillable hydrocarbons recoverable by ordinary distillation and commonly utilized fractionation facilities.
  • the charge stock is processed in a fixed-bed catalytic conversion zone in admixture with about 10,000 standard cubic feet per barrel of hydrogen, based upon fresh feed, at an inlet catalyst bed temperature of about 685 F. and an inlet pressure of about 2650 p.s.i.g.
  • the liquid hourly space velocity, based upon fresh feed only, is 0.5 and the combined feed ratio, with respect to total liquid feed, is about 2.0: 1.
  • the product streams are intended to be a 400 F. end point gasoline fraction, a 400 F. to 500 F. kerosene cut, a middle-distillate fraction boiling from 500 F. to 650 F., and a gas oil stream having an intial boiling point of about 650 F. and containing all the remaining distillable hydrocarbons in the product eflluent.
  • the vacuum column bottoms in an amount of about 3,600 bbl/day, is introduced into the process by way of line 1, is admixed with about 2.0% by volume of a slop-Wax recycle from line 2, 3,600 bbl./day of a hot separator liquid stream in line 3 and a recycled hydrogen-rich gaseous stream in line 4 (about 10,000 s.c.f./bbl. of hydrogen), the mixture continuing through line 1 into heater 6.
  • heater 6 Not illustrated in the drawing is the technique whereby the total charge to heater 6 is first heated by way of conventional heatexchange with various hot effluent streams, 'and to a temperature of about 635 F. Heater 6 raises the temperature of the feed mixture to 685 F., the heated mixture passing through line 7 into fixed-bed catalytic reactor 8.
  • a principally vaporous phase is withdrawn from hot separator 10 via line 11 and, after cooling by conventional means, is introduced into cold separator 12 at about 100 F. and a pressure of 2550 p.s.i.g.
  • a principally liquid phase is withdrawn from separator 10, by way of line 15, and a portion thereof (3,600 bbL/day) is diverted through line 3 to combine with the fresh charge in line 1 and the slop-wax recycle (about 72 bbl./day) from line 2.
  • the remainder continues via line into thermal reaction coil 16 at a pressure of about 250 p.s.i.g. and a temperature of about 750 F.
  • a hydrogen-rich gaseous phase is removed from cold separator 12 by way of line 4 through the use of compressive means not illustrated in the drawing. After make-up hydrogen is introduced via line 5, the gaseous mixture continues through line 4 to be combined with the liquid feed mixture in line 1.
  • a liquid phase from cold separator 12 is introduced into fractionator, or distillation column 14 by way of line 13.
  • Tables I and -II present the component analyses of the streams resulting from the eparations effected in hot separator 10 and cold separator 12.
  • the indicated analyses in rnol percent, do not account for various recycle streams and/or quench streams.
  • the analyses for the principally vaporous phase in line 11 and the liquid phase in line 15 are given in the following Table I.
  • fractionator 14 The principally liquid phase from cold separator 12 is introduced into fractionator 14 by way of line 13, at a locus above that at which the thermally-cracked product euent is introduced thereto via line 17.
  • the thermallycracked effluent is at a pressure of about p.s.i.g., and a temperature of 930 F. These will be changed, as will the temperature and pressure of the cold separator liquid in line 13 to the extent necessary to conform to the conditions imposed upon the distillation column in order to provide the desired product cuts.
  • fractionator 14 will function to provide an overhead fraction consisting of a minor amount of normally gaseous components and those normally liquid hydrocarbons in the gasoline boiling range-ie.
  • the heavier material at a temperature of about 800 F. is introduced into vacuum flash zone 22 functioning at about 30 mm. of Hg, absolute.
  • the vacuum flash column serves to concentrate a residuum fraction, line 25, which may be blended with thc heavy gas oil (850 F. to about 980 F.), line 24, and at least a portion of the slop-wax (980 F.1150 F.) not being recycled through line 2.
  • This mixture of slop-wax, residuurn and heavy vacuum gas oil is well-suited as a fuel oil since its sulfur content is well within the common specification of 1.0% by weight.
  • a light vacuum gas oil is removed from vacuum flash zone 22 by way of line 23.
  • the relatively minor quantity of hydrocarbonaceous material boiling below 650 F., is removed from the vacuum flash column by conventional jets which are not illustrated in the drawing.
  • the C7-400 F. fraction has an API gravity of 46.7 and a sulfur content less than 0.1% by weight; that of the 400 F.-500 F. fraction is about 0.210% by weight with a gravity of 34.5 API; the 500 F. lto 650 F. fraction has a gravity of about 29.6 API, and contains about 0.2% sulfur; and, the ⁇ 650" F.-plus portion has a concentration of 0.25% byweight of sulfur and a gravity of 19.0 API.
  • a process for the conversion of a sulfurous hydrocarbonaceous charge stock, containing less than about 150 p.p.m. metallic contaminants, and of which at least about 10.0% boils above a temperature of about 1050 F., into lower-boiling hydrocarbon products which process comprises the steps of:

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US771248A 1968-10-28 1968-10-28 Conversion and desulfurization of hydrocarbonaceous black oils Expired - Lifetime US3594309A (en)

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JP (1) JPS4830441B1 (pt)
AT (1) AT303941B (pt)
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CS (1) CS166719B2 (pt)
DK (1) DK133474B (pt)
ES (1) ES372921A1 (pt)
FR (1) FR2021705A1 (pt)
GB (1) GB1276864A (pt)
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Cited By (8)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US4446004A (en) * 1982-12-23 1984-05-01 Mobil Oil Corporation Process for upgrading vacuum resids to premium liquid products
US4604185A (en) * 1985-07-02 1986-08-05 Conoco Inc. Co-processing of straight run vacuum resid and cracked residua
US4606812A (en) * 1980-04-15 1986-08-19 Chemroll Enterprises, Inc. Hydrotreating of carbonaceous materials
US4661238A (en) * 1985-09-05 1987-04-28 Uop Inc. Combination process for the conversion of a distillate hydrocarbon to maximize middle distillate production
US4715947A (en) * 1986-11-24 1987-12-29 Uop Inc. Combination process for the conversion of a residual asphaltene-containing hydrocarbonaceous stream to maximize middle distillate production
US4721557A (en) * 1986-10-08 1988-01-26 Uop Inc. Combination process for the conversion of a residual asphaltene-containing hydrocarbonaceous stream to maximize middle distillate production
US4792390A (en) * 1987-09-21 1988-12-20 Uop Inc. Combination process for the conversion of a distillate hydrocarbon to produce middle distillate product
US4961839A (en) * 1988-05-23 1990-10-09 Uop High conversion hydrocracking process

Families Citing this family (1)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US7938953B2 (en) * 2008-05-20 2011-05-10 Institute Francais Du Petrole Selective heavy gas oil recycle for optimal integration of heavy oil conversion and vacuum gas oil treating

Family Cites Families (6)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
FR659583A (fr) * 1927-09-01 1929-07-01 Ig Farbenindustrie Ag Procédé pour la production d'hydrocarbures de grande valeur
FR1313068A (fr) * 1959-06-04 1962-12-28 Inst Francais Du Petrole Procédé continu d'hydrotraitement des produits pétroliers
FR1386576A (fr) * 1963-07-02 1965-01-22 Gulf Research Development Co Procédé de traitement d'un mélange d'hydrocarbures
US3324028A (en) * 1964-04-22 1967-06-06 Gulf Research Development Co Preparation of low sulfur content heavy fuel oils
US3371029A (en) * 1966-11-30 1968-02-27 Universal Oil Prod Co Mixed-phase conversion product separation process
US3371030A (en) * 1966-12-30 1968-02-27 Universal Oil Prod Co Black oil conversion product separation process

Cited By (9)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US4606812A (en) * 1980-04-15 1986-08-19 Chemroll Enterprises, Inc. Hydrotreating of carbonaceous materials
US4446004A (en) * 1982-12-23 1984-05-01 Mobil Oil Corporation Process for upgrading vacuum resids to premium liquid products
US4604185A (en) * 1985-07-02 1986-08-05 Conoco Inc. Co-processing of straight run vacuum resid and cracked residua
US4661238A (en) * 1985-09-05 1987-04-28 Uop Inc. Combination process for the conversion of a distillate hydrocarbon to maximize middle distillate production
EP0288619A1 (en) * 1985-09-05 1988-11-02 Uop Process for maximum middle distillate production with minimum hydrogen consumption
US4721557A (en) * 1986-10-08 1988-01-26 Uop Inc. Combination process for the conversion of a residual asphaltene-containing hydrocarbonaceous stream to maximize middle distillate production
US4715947A (en) * 1986-11-24 1987-12-29 Uop Inc. Combination process for the conversion of a residual asphaltene-containing hydrocarbonaceous stream to maximize middle distillate production
US4792390A (en) * 1987-09-21 1988-12-20 Uop Inc. Combination process for the conversion of a distillate hydrocarbon to produce middle distillate product
US4961839A (en) * 1988-05-23 1990-10-09 Uop High conversion hydrocracking process

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JPS4830441B1 (pt) 1973-09-20
GB1276864A (en) 1972-06-07
AT303941B (de) 1972-12-11
FR2021705A1 (pt) 1970-07-24
DK133474C (pt) 1976-11-01
DK133474B (da) 1976-05-24
CS166719B2 (pt) 1976-03-29
NO126919B (pt) 1973-04-09
SE369197B (pt) 1974-08-12
DE1954002A1 (de) 1970-05-21
OA03157A (fr) 1970-12-15
SU374838A3 (pt) 1973-03-20
ES372921A1 (es) 1971-11-16
NL6916218A (pt) 1970-05-01

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