US3617502A - Desulfurization and conversion of hydrocarbonaceous black oils - Google Patents

Desulfurization and conversion of hydrocarbonaceous black oils Download PDF

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US3617502A
US3617502A US771250A US3617502DA US3617502A US 3617502 A US3617502 A US 3617502A US 771250 A US771250 A US 771250A US 3617502D A US3617502D A US 3617502DA US 3617502 A US3617502 A US 3617502A
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liquid phase
percent
boiling
temperature
reaction zone
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Frank Stolfa
Laurence O Stine
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Honeywell UOP LLC
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
    • 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
    • C10G49/22—Separation of effluents
    • 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

Definitions

  • the charge stock is initially subjected to fixed-bed catalytic desulfurization and hydrogenation, and a series of separation steps to concentrate that portion of the reaction zone product boiling at temperatures above the normal gasoline boiling range. This high-boiling concentrate is then subjected to a noncatalytic, thermal cracking reaction zone or coil.
  • the process described herein is primarily adaptable to the desulfurization of petroleum crude oil residuals having relatively low metals content i.e. containing less than l50 p.p.m. of total metals. More specifically, the present invention is directed toward a combination process for hydrogenating and desulfurizing hydrocarbonaceous charge stocks which are commonly referred to as black oils.
  • the various black oil charge stocks can be classified as (1) high metals" residuals, or (2) low metals" residuals.
  • the present invention is primarily directed to the processing of those hydrocarbonaceous black oils having low metals content i.e. less than about 150 p.p.m.
  • a black oil charge stock is generally characteried as a heavy carbonaceous material of which more than about 10.0 percent by volume boils above a temperature of l,050 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 percent by weight. With many stocks, the sulfur concentration can range as high as about 5.0 percent by weight. Conradson carbon residue factors generally exceed 1.0 percent by weight, and a great proportion of black oils indicate a Conradson residue factor above 10.0.
  • Exemplary of those 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 143 API and contaminated by the presence of about 3.0 percent by weight of sulfur, 3,830 p.p.m. oftotal nitrogen, 85 ppm. of total metals, about 1 1.0 percent by weight of insoluble asphaltenes, and about 41.0 percent nondistillables.
  • the present invention affords the conversion of such charge stocks into lower boiling, normally liquid hydrocarbon products, and further converts a considerable quantity of nondistillables. Additionally, the normally liquid portion of the product effluent has been substantially desulfurized to a level less than about 1.0 percent by weight.
  • the present invention is founded upon recognition of the fact that acceptable desulfurization of low metals-containing black oils is possible at relatively mild operating severities which favor extended catalyst life without simultaneously effecting asphaltene polymerization Hydrogenation reactions are enhanced at lower severities, particularly with respect to temperature.
  • an essential feature of the present invention is the subsequent processing of the hydrogenated and desulfurized product effluent from the fixed-bed catalytic reaction zone Therefore, as hereinafter set forth in greater detail, the desulfurized catalytic reaction effluent is separated to produce a hydrocarbon stream boiling substantially completely above the gasoline boiling range, which hydrocarbon stream is subsequently subjected to a noncatalytic thermal reaction zone or coil.
  • a principal object of our invention is to provide an economical process for effecting the desulfurization and conversion of asphaltene-containing black oils to distillable hydrocarbons of lower molecular weight.
  • a corollary objective is to extend the period of acceptable, economical catalyst life while desulfurizing and hydrogenating hydrocarbonaceous black oils containing less than about 150 p.p.m. of total metals.
  • Another object is to convert a sulfurous hydrocarbon charge stock, a significant quantity of which exhibits a boiling range above a temperature of 1,050 E, into lower boiling distillable hydrocarbons having a sulfur concentration less than about 1.0 percent by weight.
  • our invention relates to a process for the conversion of a sulfurous,.hydrocarbonaceous charge stock, of which at least about 10.0 percent boils above a temperature of about l,050 F., into lower boiling hydrocarbon products, which process comprises the steps of: (a) heating said charge stock to a temperature of from 500 F.
  • the fifth separation zone is a vacuum column which serves to concentrate the unconverted asphaltic residuum and to provide at least a heavy vacuum gas oil, a light vacuum gas oil and a slop-wax cut. Generally, the latter, with or without a portion of the heavy vacuum gas oil, is recycled to the thermal cracking coil.
  • a portion of the slop-wax cut may be recycled to combine with the fresh black oil charge to the fixed-bed reaction zone.
  • the total charge to the first, fixedbed catalytic hydrogenation zone including hydrogen recycle and makeup required to maintain pressure and supplant that which is consumed within the overall process, is heated to a temperature within the range of from about 650 to about 775 F.
  • the precise temperature, to which the charge to the catalytic reaction one is heated, is controlled within the aforesaid range by monitoring the temperature of the reaction zone product effluent. Since the principal reactions being effectedare exothermic, a temperature rise is experienced as the charge and hydrogen pass through the catalyst bed.
  • the maximum'catalyst temperature which is virtually the same as that of the product effluent, is maintained at a maximum level of about 800 F.
  • the first reaction zone emuent, being introduced into the first separation zone is at a temperature of from about 700 to about 775 F. in order that the heavier constituents of the reaction zone product eflluent are not carried over into the principally vaporous phase.
  • the principal function of the present invention resides in the production of maximum quantities of distillable hydrocarbons which have been substantially reduced with respect to sulfur concentration. Through the utilization of the present combination process, this is accomplished in a highly economical fashion while avoiding the difficulties of currently practiced processing techniques.
  • Paramount is the extension of the period of time during which the fixed-bed of the solid catalytic composite functions in an acceptable manner. With respect to the processing of high metals" black oils, being those containing more than about 150 ppm. of total metals, it has been found that a successful operation involves initially visbreaking 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 subjected to a series of separation steps in order to provide aliquid phase substantially free from gasoline boiling range hydrocarbons.
  • This liquid phase is utilized as the charge to a noncatalytic thermal reaction zone, or coil.
  • this particular process offers maximum production of distillable hydrocarbons accompanied by maximum desulfurization of the charge stock whose original metals content is less than about 150 p.p.m.
  • the total charge to the fixed-bed catalytic reaction zone includes the fresh hydrocarbon charge stock, a recycled hydrogen-rich gaseous phase, makeup hydrogen, and a recycled diluent, the source of the latter being hereinafter set forth.
  • This mixture is raised to a temperature offrom about 500 to about 775 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.
  • a certain measure of temperature control, within the fixed-bed of catalyst, is afforded through the conventional utilization of either a quench hydrogen stream, or quench liquid, or both, introduced at one or more intermediate loci of the catalyst bed.
  • the catalytic reaction zone is maintained under an imposed pressure of from about l,000 to about 4,000 p.s.i.g., and the hydrocarbon charge stock contacts the catalyst at a liquid hourly space velocity of from about 0.5 to 10.0, based upon the fresh hydrocarbon charge stock exclusive of recycled diluent and/or any quench streams employed for temperature control.
  • the hydrogen concentration will be in the range of from about 5,000 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 from about l.l:l to about 3.5 :l.
  • 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 percent by weight of alumina and 12.0 percent by weight of silica, or 63.0 percent by weight of alumina and 37.0 percent by weight of silica are generally preferred.
  • Suitable metallic components having hydrogenation activity are those selected from the group consisting of the metals of Groups Vl-B and Vlll of the Periodic Table, as set forth in The Periodic Table of The Elements, E. H. Sargent & Company, 1964.
  • the catalytic composite may comprise one or more metallic components selected from the group of molybdenum, tungs
  • the concentration of the catalytic metallic component, or components is primarily dependent upon the particular metal as well as the characteristics of the charge stock.
  • the metallic components of Group Vl-B are generally present in an amount within the range of from about l.0 percent to about 20.0 percent by weight, the iron-group metals in an amount within the range of about 0.2 percent to about 10.0 percent by weight, whereas the noble metals of Group Vlll are preferably present in an amount within the range of about 0.1 percent to about 5.0 percent by weight, all of which are calculated as if these compounds existed within the catalytic composite in the elemental state.
  • the refractory inorganic carrier material may comprise alumina, silica, zirconia, magnesia, Titania boria, strontia, hafnia, and mixtures of two or more including silica-alumina, alumina-silica-boron phosphate, silica-zirconia silica-magnesia, silica-titania, alumina-zirconia, alumina-magnesia, silica-alumina-titania, alumina-magnesia-zirconia, silica-alumina-boria, etc.
  • the phrase temperature substantially the same as is employed to indicate that the only reduction in temperature stems from normally experienced loss due to the flow of material from one piece of equipment to another, or from the conversion of sensible to latent heat by "flashing" where a pressure drop occurs.
  • hydrocarbons boiling within the gasoline boiling range is intended to connote those normally liquid hydrocarbons boiling at temperatures up to about 400 or about 450 F., including pentanes and heavier hydrocarbons, and, as in some localities, butanes.
  • a commonly referred to boiling range for gas oil is an initial boiling point of about 650 F. and an end boiling point of about l,050 F.
  • the heavy gas oil characteristically is considered having an initial boiling point of about 750 F. It is, of course, recognized that a light gas oil" can have an initial boiling point as low as about 350 F. and an end boiling point as high as about 800 F. Similarly, the heavy gas oil can have an initial boiling point as low as about 650 F.
  • the principal function served by the hot separator is to separate the mixed-phase product effluent into a principally vaporous phase rich in hydrogen and a principally liquid phase containing some 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 to about 775 F., and preferably below the 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 to about 140 F serves to concentrate the hydrogen in a second principally vaporous phase.
  • the hydrogen-rich vapor phase comprising about 82.5 mol percent hydrogen, and only about 2.3 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 in a second principally liquid phase.
  • the first liquid phase from the hot separator may be 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 1.121 to about 3.5:l.
  • the remaining portion of the principally liquid phase from the hot separator is introduced into a third separation one hereinafter referred to as the hot flash zone.
  • the hot flash zone functions at about the same temperature as the liquid phase withdrawn from the hot separator, but at a significantly reduced pressure of from about 150 to about 350 p.s.i.g.
  • the principally vaporous phase from the hot flash zone comprises primarily hydrocarbons boiling below a temperature of about 650 F and containing a relatively minor quantity of hydrocarbons normally considered to be within the heavy gas oil boiling range.
  • This principally vaporous stream may be combined with the liquid stream from the cold separator, and the mixture introduced into a cold flash zone at a pressure of from atmospheric to about 60 p.s.i.g. and a temperature of from 60 to about 140 F.
  • the principally liquid phase withdrawn from the hot flash zone is introduced into a thermal cracking reaction zone, or coil, at substantially the same temperature, and a pressure of from about 150 to about 350 p.s.i.g.
  • the thermally cracked product efi'luent, at a temperature of from about 875 to about 950 F., and a pressure of from about 40 to about 100 p.s.i.g., is cooled to a temperature of about 700 F., and introduced into a fourth separation zone hereinafter referred to as the flash fractionator."
  • the liquid phase from the flash fractionator is introduced into a vacuum column maintained at about 25 to about 75 mm. of Hg., absolute.
  • the vacuum column serves as the fifth separation zone, the principal function of which is the concentration and separate recovery of an asphaltic residuum, containing high molecular weight sulfurous compounds and being substantially free from distillable hydrocarbons.
  • gas oil streams are recovered from the vacuum column as a separate light vacuum gas oil (LVGO) having a boiling range of from about 320 to about 750 F., a medium vacuum gas oil (MVGO) boiling from about 750 to about 980 F., and a heavy vacuum gas oil containing the remainder of the distillable hydrocarbons.
  • a preferred technique is to separate a slop-wax cut, from the vacuum column, which contains primarily these distillables boiling above 980 F but may consist of up to about 30.0 percent by volume of the total distillables boiling above 750 F.
  • a portion of the slop-wax cut may be recycled to the catalytic hydrogenation/desul-v furization reaction zone, it is generally recycled to the thermal coil in order to increase the yield of the more desirable gas oils.
  • the amount of slop-wax so recycled is such that the combined feed ratio to the thermal reaction coil is above about 12:1 and generally not higher than about 3.0: l.
  • the drawing will be described in connection with the conversion ofa vacuum column bottoms product having a gravity of 6.0 AP] and an ASTM 20.0 percent volumetric distillation temperature of about l,055 F.
  • the charge stock contains 4,000 p.p.m. of nitrogen, 5.5 percent by weight of sulfur, p.p.m. of nickel and vanadium, 6.0 percent by weight of heptane-insoluble asphaltenes and has a Conradson carbon residue factor of 21.0 percent by weight.
  • the description will be directed toward a commercially scaled unit having a capacity of about 8,000 barrels per stream day.
  • the charge stock be converted into maximum distillable hydrocarbons which are recoverable by ordinary distillation techniques in commonly utilized fractionation systems.
  • the charge stock is processed in a fixed-bed catalytic desulfurization and desulfurization zone in admixture with about 10,000 s.c.f./bl. of hydrogen, based upon fresh feed exclusive of recycle streams, at a catalyst bed inlet temperature of about 700 F., and a pressure of about 3,105 p.s.i.g.
  • the liquid hourly space velocity, based upon fresh feed only, is about 0.5, and the combined liquid feed ratio is about 2.0: l.
  • Hot Separator Stream Analyses amount of about 7,678 bl./day (185.94 mols/hr.), is introduced into the system by way of line 1, and following heatexchange with various hot effluent streams, is passed into line 8 l 12 heater in admixture with a recycled hydrogen-rich stream 5 from line 3 and a hot separator bottoms liquid recycle in line Nitrogen l2.92 9.40 1.67 2.
  • Makeup hydrogen from a suitable external source, to mam Hydrogen 76792 736' Isa-o taln plant pressure, and to replace that hydrogen consumed in Hydrogen Sulfide 90643 356 ⁇ ; the overall process, is introduced by way of line 4.
  • the total charge to the heater is at a temperature of about 500 F.; this Mflhan, 88.95 mg, 2355 is increased to a level of about 700 F., as measured at the inlet Ethane 160.61 146.30 6.78 to the catalyst bed.
  • the thus-heated total charge passes Pmpan 8952 through line 6 into fixed-bed catalytic reaction one 7.
  • the catalyst disposed in reaction one 7 is a composite of 88.0 per- 5 I if: 1'2: cent by weight of alumina and 12.0 percent by weight of silica, Hem, 17:60 I 5 with which is combined 2.0 percent by weight of nickel and about 16.0 percent by weight of molybdenum, calculated as 3140 16,02 s the elemental metals. 320 520" F. 68.54 47.20 10.1 1
  • the stream in line 12 comprises about 19.9 mol percent bu- Nitrogen 11.07 1.85 12.93 tanes and lighter material, exclusive of hydrogen which is dis- "y l 843099 l75-52 8605-61 solved in the heavier hydrocarbons, and is considered, there- Hydrogen Sulfide 73902 26.36 765.38 fore a p p y liquid p That portion of the hot separator bottoms stream not diverted through line 2, continues through line 12 into hot Ethane 133.14 7.53 140.67 n h 13 A d r ff t d b r Propane use 7 7] as one re uc 1on 1n pressure is e co e y means 0 a reducmg valve not indicated 1n the drawmg, and the stream Bum" M49 L06 2655 40 enters hot flash one 13 at a pressure of about 250 p.s.i.g.
  • the "cranes 4.70 1.29 5.99 principal function of flash zone 13 is to concentrate the heavier components in a liquid phase which serves as the charge to c,-' F. 2.00 3.19 5.19 thermal cracking coil 17.
  • the Z- vaporous phase in line 14 comprises about 89.2 mol percent of 228328, 5:: material boiling below about 520 F. exclusive of hydrogen,
  • Line No. 14 16 The conversion product effluent, in mixed phase in line 8, at 1 temperature of about 800 F., is utilized as a heat-exchange medium, and is introduced into hot separator 9 at a temperature of 775' F. and a pressure of about 3,040 p.s.i.g. A prin- Nitrogen 1.45 0.12 cipally vaporous phase is withdrawn by way of line 10, and a principally liquid phase via line 12.
  • the liquid phase in line 16 is admixed with about 1.0 wt. percent of 230 p.s.i.g. steam, and the mixture enters thermal coil 17 at a temperature of about 740 F. and a pressure of about 170 p.s.i.g.
  • the thermally cracked produce effluent at a 7 pressure of about 55 p.s.i.g. and a temperature of about 930 F., and, after being cooled, passesvia line 18 into a rectified flash fractionator 19 at a temperature of about 700 F. and a pressure of about 55 p.s.i.g.
  • a vapor phase is withdrawn from flash fractionator 19 through line 20, and a liquid phase through line 24.
  • the latter is introduced into vacuum flash column 25 at a temperature of about 750 F.
  • the vacuum column is functioning at about 25 mm. of Hg., absolute through the utilization of standard vacuum jets which are not illustrated in the drawing.
  • the separation effected in flash fractionator 18 is presented in the following table 1V, along with the component analysis of the thermally cracked product effluent in line 18, exclusive of water.
  • the principally vaporous phase withdrawn from hot separator 9 through line 10 is cooled to a temperature of about 120 F., and is introduced into cold separator 11 at a pressure of about 3,000 p.s.1.g.
  • a hydrogen-rich gaseous phase is withdrawn through line 3, and is recycled therethrough to combine with the fresh hydrocarbon charge stock in line 1.
  • a principally liquid phase is withdrawn from cold separator l 1 through line 15.
  • the separation effected in cold separator 11, exclusive of makeup hydrogen, is presented in the following table V.
  • Vacuum flash column 25 serves to concentrate the residuum, 39.41 mols/hr. leaving via line 28, and also to separate a light vacuum gas oil (LVGO), line 26 and a heavy vacuum gas oil (HVGO), line 27.
  • the HVGO having a boiling range of 750 to 1,050 F., is in an amount of about 66.62 mols/hr., and the LVGO is in a amount of 58.45 mols/hr.
  • Lighter material boiling below 320 F. is removed from vacuum flash column 25 by the jets which are not indicated in the drawing.
  • Cold flash zone 21 has been illustrated for the sake of completeness, indicating the separation of the mixture of the hot flash vapors (line 14), the flash fractionator vapors (line 20) and the cold separator liquid (line 15). In a commercial installation, the vapors from the flash fractionator would be recovered separately due to the relatively high degree of olefinicity thereof. Component analyses indicating the separation effected in cold flash zone 21 are presented in the following table V1.
  • the sulfur concentration of the distillable hydrocarbon products is about 0.83 percent by weight.
  • the fixed-bed catalytic reaction one must necessarily be operated at a significantly higher severity level in order to produce the maximum quantity of distillables.
  • the use of the present process affords a reduction in operating severity, as measured by the catalyst bed inlet temperature, of from 50 to as much as 100 F.
  • the present process increases catalyst life (expressed as barrels per pound) from 50 to'80 percent, resulting in longer on-stream cycles.
  • a reduction in the size of the vacuum column, from a nominal diameter of 11.0 to 8.6 feet, is made possible. This, as will be noted by those skilled in the art of petroleum processing techniques, affords a significant reduction in capital outlay for equipment.
  • a process for the conversion of a sulfurous, hydrocarbonaceous charge stock, of which at least about 10.0 percent boils above a temperature of about l,050 F., into desulfurized lower-boiling hydrocarbon products comprises the steps of:

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US771250A 1968-10-28 1968-10-28 Desulfurization and conversion of hydrocarbonaceous black oils Expired - Lifetime US3617502A (en)

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US3920538A (en) * 1973-11-30 1975-11-18 Shell Oil Co Demetallation with nickel-vanadium on silica in a hydrocarbon conversion process
US4005006A (en) * 1975-07-18 1977-01-25 Gulf Research & Development Company Combination residue hydrodesulfurization and thermal cracking process
US4061562A (en) * 1976-07-12 1977-12-06 Gulf Research & Development Company Thermal cracking of hydrodesulfurized residual petroleum oils
US4097363A (en) * 1976-07-12 1978-06-27 Gulf Research & Development Company Thermal cracking of light gas oil at high severity to ethylene
US4606812A (en) * 1980-04-15 1986-08-19 Chemroll Enterprises, Inc. Hydrotreating of carbonaceous materials
US4925573A (en) * 1988-03-31 1990-05-15 Shell Internationale Research Maatschappij, B.V. Process for separating hydroprocessed effluent streams
US4961839A (en) * 1988-05-23 1990-10-09 Uop High conversion hydrocracking process
US5110444A (en) * 1990-08-03 1992-05-05 Uop Multi-stage hydrodesulfurization and hydrogenation process for distillate hydrocarbons
US5114562A (en) * 1990-08-03 1992-05-19 Uop Two-stage hydrodesulfurization and hydrogenation process for distillate hydrocarbons
US6605208B2 (en) * 2000-11-03 2003-08-12 Sanford P. Brass Process for reduction of emissions in asphalt production

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US9822317B2 (en) 2014-10-10 2017-11-21 Uop Llc Process and apparatus for selectively hydrogenating naphtha
US9393538B2 (en) 2014-10-10 2016-07-19 Uop Llc Process and apparatus for selectively hydrogenating naphtha

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US2339918A (en) * 1940-09-13 1944-01-25 Universal Oil Prod Co Hydrocarbon conversion
US2355366A (en) * 1942-01-12 1944-08-08 Phillips Petroleum Co Process for catalytically desulphurizing hydrocarbon oil
US3409538A (en) * 1967-04-24 1968-11-05 Universal Oil Prod Co Multiple-stage cascade conversion of black oil

Cited By (13)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
DE2304450A1 (de) * 1973-01-30 1974-08-01 Linde Ag Verfahren zur gewinnung von schwefelarmen spaltprodukten aus schwefelreichem rohoel oder schwefelreichen rohoelfraktionen
US3920538A (en) * 1973-11-30 1975-11-18 Shell Oil Co Demetallation with nickel-vanadium on silica in a hydrocarbon conversion process
US4005006A (en) * 1975-07-18 1977-01-25 Gulf Research & Development Company Combination residue hydrodesulfurization and thermal cracking process
US4061562A (en) * 1976-07-12 1977-12-06 Gulf Research & Development Company Thermal cracking of hydrodesulfurized residual petroleum oils
US4097363A (en) * 1976-07-12 1978-06-27 Gulf Research & Development Company Thermal cracking of light gas oil at high severity to ethylene
US4606812A (en) * 1980-04-15 1986-08-19 Chemroll Enterprises, Inc. Hydrotreating of carbonaceous materials
US4925573A (en) * 1988-03-31 1990-05-15 Shell Internationale Research Maatschappij, B.V. Process for separating hydroprocessed effluent streams
US4961839A (en) * 1988-05-23 1990-10-09 Uop High conversion hydrocracking process
US5110444A (en) * 1990-08-03 1992-05-05 Uop Multi-stage hydrodesulfurization and hydrogenation process for distillate hydrocarbons
US5114562A (en) * 1990-08-03 1992-05-19 Uop Two-stage hydrodesulfurization and hydrogenation process for distillate hydrocarbons
US6605208B2 (en) * 2000-11-03 2003-08-12 Sanford P. Brass Process for reduction of emissions in asphalt production
US20030205506A1 (en) * 2000-11-03 2003-11-06 Kenneth Hucker Process for reduction of emissions in asphalt production
AU2002220080B2 (en) * 2000-11-03 2005-03-17 Sanford P. Brass Process for reduction of emissions in asphalt production

Also Published As

Publication number Publication date
YU33886B (en) 1978-06-30
FR2030066B1 (nl) 1973-11-16
GB1276920A (en) 1972-06-07
YU269469A (en) 1977-12-31
NO126921B (nl) 1973-04-09
AT303940B (de) 1972-12-11
NL162128C (nl) 1980-04-15
CS167269B2 (nl) 1976-04-29
FR2030066A1 (nl) 1970-10-30
ES372923A1 (es) 1971-11-16
DE1954004A1 (de) 1970-06-18
BR6913725D0 (pt) 1973-02-27
NL6916220A (nl) 1970-05-01
SE367213B (nl) 1974-05-20
NL162128B (nl) 1979-11-15
JPS4925403B1 (nl) 1974-06-29
OA03159A (fr) 1970-12-15

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