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US3371029A - Mixed-phase conversion product separation process - Google Patents

Mixed-phase conversion product separation process Download PDF

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Publication number
US3371029A
US3371029A US598067A US59806766A US3371029A US 3371029 A US3371029 A US 3371029A US 598067 A US598067 A US 598067A US 59806766 A US59806766 A US 59806766A US 3371029 A US3371029 A US 3371029A
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phase
vapor phase
line
hydrocarbons
liquid
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US598067A
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Inventor
Jack N Weiland
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Universal Oil Products Co
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Universal Oil Products Co
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Priority to US598067A priority Critical patent/US3371029A/en
Application filed by Universal Oil Products Co filed Critical Universal Oil Products Co
Priority to ES347843A priority patent/ES347843A1/es
Priority to DE19671645825 priority patent/DE1645825A1/de
Priority to BE707354D priority patent/BE707354A/xx
Priority to BR195153/67A priority patent/BR6795153D0/pt
Priority to TR16585A priority patent/TR16585A/xx
Priority to GR670134824A priority patent/GR34824B/el
Priority to FR130317A priority patent/FR1547873A/fr
Priority to CH1685567A priority patent/CH542276A/de
Priority to SE16469/67A priority patent/SE345870B/xx
Priority to GB54558/67A priority patent/GB1197469A/en
Priority to OA53119A priority patent/OA02549A/xx
Priority to YU2395/67A priority patent/YU32372B/xx
Priority to NL6717013A priority patent/NL6717013A/xx
Application granted granted Critical
Publication of US3371029A publication Critical patent/US3371029A/en
Anticipated expiration legal-status Critical
Expired - Lifetime legal-status Critical Current

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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
    • C10G5/00—Recovery of liquid hydrocarbon mixtures from gases, e.g. natural gas

Definitions

  • the present invention involves a particular scheme for separating a mixed-phase hydrocarbonaceous product eilluent resulting from the conversion of a heavier-than-gasoline (boiling substantially, completely above a temperature of 400 F.) hydrocarbon charge stock, which product eflluent contains hydrogen, normally liquid hydrocarbons and normally gaseous hydrocarbons.
  • the mixed-phase separation process hereinafter described in detail is applicable to a hydrocarbon conversion process which may be classified as hydrogen-consuming, and in which processing techniques dictate the recycle of a hydrogen-rich gaseous phase and, in many instances, the recycle of at least a portion of the normally liquid product eiliuent.
  • Such hydrogen-consuming processes include the hydroreiining, or hydrotreating of kerosene fractions, middle-distillate fractions, light and heavy vacuum gas oils, light and heavy cycle stocks, etc., for the primary purpose of reducing the concentration of various contaminating intluences contained therein.
  • Another typical hydrogen-consuming hydrocarbon conversion process is known in the petroleum rening art as hydrocracking.
  • hydrocracking techniques are utilized to convert relatively heavy hydrocarbonaceous material into lower-boiling hydrocarbon products such as gasoline and fuel oil.
  • desired end result is the production of liquefied petroleum gas.
  • Relatively recent developments in the area of petroleum technology have indicated that the hydrocracking reactions can be applied successfully to residual stocks, or so-called black oils.
  • Exemplary of the material classified as black oils are atmospheric tower bottoms products, vacuum tower bottoms products (vacuum residuum), crude oil residuum, topped crude oils, crude oils extracted from tar sands, etc.
  • Black oils particularly the heavy oils extracted from tar sands, topped or reduced crudes, and vacuum residuum, etc., contain high molecular weight sulturous cornpounds in exceedingly large quantities.
  • these black oils contain excessive quantities of nitrogenous compounds, high molecular weight organo-metallic cornplexes comprising nickel and vanadium, and a considerable amount of asphaltic material.
  • a signicant proportion of which has a gravity less than 10.0 This material is generally further characterized by a boiling range indicating that 10.0% or more, by volume, boils above a temperature of about 1050" F. Although the amount is not known accurately, a significant quantity of the available black oils are further characterized in that more than 50.0% by volume thereof boils above a temperature of about 1050 F.
  • the utilization of these high molecular weight black oils as .a source of more valuable liquid hydrocarbon products,I such as gasoline and fuel oil is precluded by present-day rening techniques, due especially to the exceedingly high sulfur and asphaltic concentrations. The conversion of a portion of such material into distillable hydrocarbons-ie.
  • black oils examples include a vacuum tower bottoms product having a gravity of 7.1 API at 60 F., and containing 4.1% by weight of sulfur and 23.7% by weight of asphaltics; a topped Middle East Kuwait crude oil, having a gravity of 11.0 API at 60 F., and containing 10.1% by weight of asphaltics and Iabout 5.2% by weight of sulfur; and a vacuum residuum having a gravity of 8.8 API at 60 F., and containing 3.0% by weight of sulfur and 4300 ppm. of nitrogen, and having a 20.0% volumetric distillation point at 1055 F.
  • the asphaltic material is found to be colloidally dispersed within the black oil, and, when subjected to elevated temperatures, has the tendency to occulate and polymerize, whereby the conversion thereof to more valuable oil-soluble products becomes extremely ditlicult.
  • the heavy bottoms from a crude oil vacuum distillation column indicates a Conradson Carbon Residue factor of, for instance, 16.0% by Weight.
  • Such a material is useful only as road asphalt, or as an extremely low grade fuel when cut-back with distillate hydrocarbons such as kerosene, light gas oil, etc.
  • the principal object of the present invention is to provide an improved process for separating a mixed-phase hydrocarbonaceous reaction product eluent, which product etlluent contains hydrogen, normally liquid hydrocarbons and normally gaseous hydrocarbons.
  • Another object of this invention is to afford a mixedphase conversion product separation process, which product effluent contains hydrogen, normally liquid hydrocarbons, a portion of both of which is intended to be recycled to a conversion process, and normally gaseous hydrocarbons including methane, ethane, and propane.
  • Another object is to convert sulfur-contaminated black oils having a gravity, at 60 F., of less than about 20.0 API, and a boiling range indicating that a substantial portion thereof is non-distillable, into lower boiling distillable hydrocarbon products of significantly reduced sulfur concentration.
  • the present invention alfords a process for separating a mixed-phase hydrocarbonaceous reaction product eluent, resulting from the conversion of a hydrocarbon charge stock boiling above a temperature of about 400 F., said product efliuent containing hydrogen to be recycled, normally liquid hydrocarbons, and normally gaseous hydrocarbons, which process comprises the steps of: (a) separating said etlluent in a first separation zone under substantially the same pressure as said eiuent, to provide a first liquid phase and a first vapor phase; (b) cooling said vapor phase to a temperature within the range of from about 60 F.
  • the latter includes recycle of a portion of said fourth liquid phase to combine with the cooled first vapor phase, prior to separation of the latter in said second separation zone. Also, a portion of said first liquid phase is recycled to combine with the charge stock to the conversion zone.
  • the first separation zone referred to as a hot separator, is maintained at essentially the sarne pressure as the reaction product effluent being initially separated therein, and, for the various hydrogen-consuming conversion processes hereinbefore described, such pressure is in the range of from about 1000 p.s.i.g. to about 3000 p.s.i.g.
  • the temperature of the reaction product eiuent, as it enters this hot separator is below about 750 P. At temperatures above 750 F., the heavier normally liquid hydrocarbons are carried over in the first vapor phase, whereas at temperatures below about 700 F., ammonium salts, resulting from the conversion of nitrogenous compounds contained within the hydrocarbonaceous charge stock, tend to fall into the liquid phase.
  • the second separation zone although maintained under essentially the same pressure as the reaction product effluent and the hot separator, is at a temperature of from about 60 F. to about 140 F., and is referred to as a cold separator.
  • the third and fourth separation zones may be maintained at substantially the same pressure, the pressure therein is substantially reduced from the pressure under which the hot and cold separators are maintained.
  • the pressure of the third and fourth separation zones will generally be superatmospheric, the maximum pressure will be about 200 p.s.i.g.
  • the third separation zone referred toas a hot ash zone, will, however, operate at an elevated temperature somewhat 'less than the temperature of the first liquid phase emanating from the first separation zone, and generally above about 700 F.
  • the fourth separation zone referred to as a cold flash zone, will operate at a significantly reduced temperature within the range of about 60 F. to about 140 F.
  • a first product stream is essentially a gaseous phase rich in hydrogen, and generally containing at least about 80.0 mol percent thereof, and less than about 0.1% of normally Iliquid hydrocarbons; it is therefore, extremely well suited as a hydrogen-rich recycle gaseous phase.
  • a second product stream also essentially a gaseous phase, contains about 97.5 mol percent propane and lighter gaseous components, including substantial quantities of hydrogen sulde resulting from the conversion of sulfurous compounds.
  • the third product stream consists essentially of normally liquid hydrocarbon products which may be subjected to fractionation for the purpose of obtaining particularly desired selected fractions thereof.
  • gasoline boiling range up to about 380 F.
  • middle-distillate hydrocarbons in an amount of about 10.0% by volume ⁇ and about 88.0% by volume of a fuel oil containing less than 1.0% by weight of sulfur.
  • the present invention comprises a series of integrated steps for the separation of a mixed-phase reaction product effluent in an easy and economical manner.
  • the present invention is uniquely adaptable to processes designed for the conversion of black oils.
  • the novel mixedphase separation process of the present invention is equally applicable to various reaction product effluent streams which may be obtained from sources other than the conversion of such hydrocarbon black oils.
  • the conversion of the previously described black oils will be employed.
  • the conversion of black oils is intended to accomplish primarily two objects: first, to desulfurize the black oil to the extent dictated by the desired end result, whether maximizing fuel oil, or gasoline boiling range hydrocarbons; secondly, it is intended to produce distillable hydrocarbons, being those normally liquid hydrocarbons including pentanes, having boiling points below about 1050 F.
  • the conversion conditions are those conditions imposed upon a conversion zone for the purpose of achieving both desulfurization and conversion into lower-boiling hydrocarbon products. It Will be noted by those skilled in the art of petroleum refining techniques, that the conversion conditions hereinafter enumerated are significantly less severe than those being currently commercially employed in processing similar charge stocks.
  • the distinct economic advantages, over and above those normally stemming from the production of the more valuable distillable hydrocarbons, will be recognized.
  • the conversion conditions are intended to include temperatures above 700 F., with an upper limit of about 800 F., as measured at the inlet to the fixed-bed of catalyst disposed within the reaction zone. Since the bulk of the reactions being effected are exothermic, the reaction zone effluent will be at a higher temperature. In order that catalyst stability be preserved, it is preferred to control the inlet temperature at a level such that the temperature of the reaction product eflluent does not exceed 900 F.
  • Hydrogen is admixed with the black oil charge stock, by means of compressive recycle, in an amount usually less than about 10,000 s.c.f./bbl., at the selected operating pressure; the hydrogen is present in the recycle gaseous phase preferably in an amount of about 80.0% or more.
  • a preferred range of the quantity of hydrogen being admixed with the fresh black oil charge stock is from about 3000 to about 6000 s.c.f./bbl.
  • the conversion reaction zone will be maintained at a pressure greater than about 1000 p.s.i.g., and generally in the range of about 1500 p.s.i.g. to about 3000 p.s.i.g.
  • the point of pressure measurement is generally either the discharge of the compressive means, the inlet to the catalyst bed, or the pressure in the cold separator.
  • the black oil passes through the catalyst at a liquid hourly space velocity (defined as volumes of liquid hydrocarbon charge per hour, as measured at 60 F., per volume of catalyst disposed within the reaction zone) of from about 0.25 to about 2.0. Notwithstanding that the conversion of black oils may be conducted in a batchwise fashion, it readily lends itself to the more economical continuous processing in an enclosed vessel. When conducted as a continuous process, it is preferred to introduce the hydrogen-hydrocarbon mixture into the vessel in such a manner that the same passes therethrough in downward ow.
  • the internals of the vessel may be constructed in any suitable manner capable of providing the required contact between the liquid charge stock, the gaseous mixture and the catalyst.
  • hydrogen is employed in admixture with the charge stock, and preferably in an amount of from about 3000 to about 6000 s.c.f./bbl.
  • the hydrogen-containing gaseous phase herein sometimes designated as recycle hydrogen since it is conveniently recycled externally of the conversion zone, fulfills a number of various functions; it serves as a hydrogenating agent, a heat carrier, and particularly a means for stripping converted material from the catalytic composite, thereby creating still more available catalytically active sites for the incoming, unconverted hydrocarbon charge stock.
  • recycle hydrogen serves as a hydrogenating agent, a heat carrier, and particularly a means for stripping converted material from the catalytic composite, thereby creating still more available catalytically active sites for the incoming, unconverted hydrocarbon charge stock.
  • the catalytic composite disposed within the reaction zone can be characterized as comprising a metallic component possessing hydrogenation activity, which component is composited with a refractory inorganic oxide carrier material which may be of either synthetic or natural origin.
  • a refractory inorganic oxide carrier material which may be of either synthetic or natural origin.
  • the precise composition and method of manufacturing the carrier material is not considered to be an essential element of the present process, although a siliceous carrier, such as 88.0% by weight of alumina and 12.0% by weight of silica, or 63.0% alumina and 37% silica, are generally preferred for processes designated to convert black oils.
  • Suitable metallic components, having hydrogenation activity are those selected from the -group consisting of the metals of Group VI-B and VII of the Periodic Table, as indicated in the Periodic Chart of the Elements, Fisher Scientific Company (1953).
  • the catalytic composite may comprise one or more metallic components from the group of molybdenum, tungsten, chromium, iron, cobalt, nickel, platinum, palladium, iridium, osrnium, rhodium, ruthenium, and mixtures thereof.
  • concentration of the catalytically active metallic cornponent, or components is dictated by the particular metal as well as the physical and chemical characteristics of the black oil charge stock.
  • the metallic components of Group Vl-B are ⁇ generally present in an amount within the range of 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 platinum-group metals are preferably present in an amount within the range of about 0.1% to about 5.0% by weight, all of which are calculated as if the components existed within the finished catalytic composite as the elemental metal.
  • the refractory inorganic oxide 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, alumina-titania, magnesia-zirconia, titania-zirconia, magnesia-titania, silica-alumina-zirconia, silica-alumina-magnesia, silica-alumina-titania, silica-magnesia-zirconia, silica-alumina-boria, etc. It is preferred to utilize a carrier material containing at least a portion of silica, and preferably a composite of alumina and si
  • the drawing will be described in connection with the conversion of a Middle East reduced crude oil having a gravity of 16.6 API at 60 F., and an ASTM 65.0% volumetric distillation temperature 0f 1034 F.
  • the reduced crude oil contains about 3.8% by weight of sulfur, 2032 p.p.m. of nitrogen, 6.5% by weight of pentane-insoluble asphaltics, a Conradson Carbon Residue factor of 8.0 weight percent and about p.p.m. of metals, the latter being principally nickel and vanadium.
  • this object be accomplished with minimal production of propane and lighter hydrocarbons. That is, with respect to that portion of the conversion product efliuent boiling at temperatures below about 650 F., and constituting normally gaseous hydrocarbons, gasoline boiling range hydrocarbons, and middle-distillate hydrocarbons, it is intended that there be minimum production of the light, normally gaseous hydrocarbons, with accompanying maximum production of the normally liquid hydrocarbons.
  • the 40,000 bbl./ day of reduced crude oil, 557,620 lbs./ hr., entering the process via line 1, is admixed with makeup hydrogen of about 97.5 mol percent purity in an amount of 9440 lbs/hr., from an external source indicated by line 2. It has been found appropriate in some instances to add water to the reaction zone in admixture with the charge stock. When this is deemed advisable, it takes place via line 3; for the present illustrative purposes, it is presumed that such water addition is not being effected.
  • the hydrogen-crude oii mixture continues through line 1, being further admixed with 159,420 lbs/hr.
  • Heater 5 is employed to raise the temperature of the charge to about 705 F., and the heated mixture in line 6 is admixed with 527,190 lbs/hr. of a hot recycle stream (750 F.) in line 7; the reactor charge at 720 F. continues through line 6 into conversion reactor S at a pressure of about 2125 p.s.i.g.
  • the catalyst disposed in conversion zone 8 is a composite of 2.0% ⁇ by weight of nickel, 16.0% by weight of molybdenum and a carrier material of 68.0% by weight of alumina, 22.0% by weight of boron phosphate and 10.0% by weight of silica.
  • the reduced crude oil contacts the catalyst at a liquid hourly space velocity of 0.8, and the combined feed ratio, based only on normally liquid charge, is 2.0.
  • the total conversion product efiiuent leaves reactor 8 via line 9, and passes therethrough into hot separator 10. Since the conversion efiiuent product is at a temperature of about 780 F. and a pressure of 2075 p.s.i.g., it is employed as a heat-exchange medium in order that its temperature be lowered to 750 P., prior to entering hot separator 10.
  • the pressure within hot separator 19 is about 2060 p.s.i.g., lower than reactor 8 inlet pressure due to pressure drop through the system.
  • a first liquid is withdrawn from separator 10 through line 11, in an amount of 975,600 lbs./ hr., and of this amount 527,190 lbs/hr. are diverted through line 7 t0 combine with the heated mixture in line 6.
  • the remaining portion, 448,410 lbs/hr., continues through line 11 into hot flash zone 24.
  • a first vapor phase in an amount of 278,070 lbs/hr. is removed from hot separator 10 through line 12, passes through condenser 13 whereby the temperature is lowered to 120 F.; at this point, due to pressure drop through the system, the pressure is about 2005 p.s.i.g.
  • the cooled first vapor phase passes through line 14, is admixed with a portion, 275,780 lbs./hr., of a fourth liquid phase in line 23 hereinafter described, and the mixture is introduced into cold separator 15.
  • a second vapor phase containing about 800 mol percent hydrogen, in an amount of 159,420 lbs/hr., is removed via line 16, is raised to a pressure of about 2245 p.s.i.g.
  • the first liquid phase in line 11 entering hot flash zone 24 is at a temperature of about 745 F. and a substantially reduced pressure of about 220 p.s.i.g.
  • a third liquid phase is removed via line 27 in an amount of 425,610 lbs./ hr., to be combined with a fourth liquid phase, hereinafter described, as the major product stream.
  • a third vapor phase is removed through line 25 in an amount of 22,800 lbs/hr., and is cooled to about 105 F. in condenser 26, prior to continuing through line 19 into cold ash separator 20.
  • the cooled third vapor phase is combined with a second liquid phase in line 18 from cold separator 15, the latter in an amount of 393,930 lbs/hr., the total charge to the cold flash zone thus being 416,730 lbs./ hr.
  • the material entering cold flash separator is at a pressure of about 200 p.s.i.g. and a temperature of 105 F.
  • a fourth vapor phase in an amount of 19,450 lbs/hr. (97.5 lmol percent, propane and lighter normally gaseous components), is removed from separator 20 through line 21. Since the material contains a considerable quantity of hydrogen sulfide, it is generally subjected to a suitable treating process prior to being vented and/or burned as flue gas. The particular economic aspects to be considered will dictate whether the fourth vapor phase is suitably treated to recover the small quantity of C4-plus normally liquid hydrocarbons contained therein.
  • a fourth liquid phase in an amount of 397,280 lbs./ hr., is removed from cold flash zone 20y through line 22. Of this amount, 275,780 lbs/hr.
  • Fractionator 39 will be operated ⁇ at conditions of Itemperature and pressure, and will be designed in accordance with the desired fractions to be recovered there- F from. With respect to the commercially-scaled unit being described, as previously stated, the primary objective was to maximize the production of fuel-oil (650 F.plus) having a sulfur concentration not ⁇ greater than 1.0% by weight. This product, in an amount of 480,550 lbs/hr., is indicated as leaving fractionator 30 via line 33. A second, middle-distillate fraction (380 F.-650 F.) is removed via line 32 in an amount of 42,440 lbs/hr. The gasoline boiling range material, having an end boiling point Iof 380 F., is removed via line 31 in an amount of 24,120 lbs/hr.
  • the arnmonia and/or ammonium salts contained in conversion zone 8 effluent may be removed fro-m the process by being adsorbed in water which is injected into the reaction product efiiuent before the same is passed into hot separator 10.
  • the water and ammonia are removed with the first vapor phase, introduced into cold separator 15, and subsequently withdrawn via line 34.
  • the water may be injected into the rst vapor phase leaving hot separator 10.
  • Table I illustrates the composition of the conversion zone effluent (line 9), the first vapor phase (line 12) and the first liquid phase (line 11) prior to diverting a portion to the conversion zone via line 7.
  • the 19.36 mols/hr. of water in the reaction zone effluent is a consequence of the water of saturation in the recycle gas. As hereinafter indicated, a minor portion of this water finds its way into the hydrogen-rich recycled gaseous phase (line 16 in the drawing).
  • the function of hot separator 10 is indicated by virtue of the fact that 63.3 mol percent of the first liquid phase (line 11) consists of 380 F.plus hydrocarbons, While the first vapor phase contains 98.6 mol percent of material boiling -below about 380 F. However, the first liquid phase comprises about 24.7 mol percent of hydrogen and a total of 36.7% material boiling below about 380 F.
  • Hot ash zone 24 serves to ⁇ separate light gaseous material from the normally liquid hydrocarbons ultimately fractionated in fractionator 30. Were these gaseous components not removed at this point, the presence of the same in fractionator 30 would make overhead condensation for reflux purposes extremely diiiicult, and would unnecessarily adversely aiiect the recovery of gasoline hydrocarbons. A number of compressive stages and/or absorption zones would be required to fractionate to recover the desired product streams. ln the following Table III, the attainment of this objective is clearly illustrated by the component analyses of that portion of the first liquid phase not diverted through line 7, (termed line 11a for convenience in the table), the third vapor phase from the hot flash Zone (line 25) and the third liquid phase (line 27).
  • the third liquid phase (line 27) contains 94.7 molpercent of 380 F.plus hydrocarbons, and only 2.5 mol percent of hydrogen. Taking into account the C75-380 F. gasoline portion, the third liquid phase constitutes only 4.3 mol percent of material boiling below hexane.
  • Table 1V the function of cold flash zone 20 is illustrated by the component analyses of the feed thereto in line 19, being the mixture of the second liquid phase and the cooled third vapor phase, the fourth vapor phase (line 21) and the fourth liquid phase (line 22).
  • Table IV includes the component analysis of the combined third (line 27) and fourth (line 22) liquid phases, after a portion of the latter has been diverted via line 23 to combine with the rst vapor phase in line 14. This would be the material continuing through line 22 into heater 28 for product separation in fractionator 30.
  • Table V indicates the overall yields on both a volumetric and Weight basis.
  • the desired product of the black oil conversion process described was the maximum quantity of fuel oil (650 F.plus) having a sulfur content of less than 1.0% by Weight. As seen from the following table, this product was obtained in an amount of 87.9% by volume.
  • the desired product was recovered at the expense of the production of only 0.7% by weight of light gaseous waste products, methane, ethane and propane.
  • 1510 bbl./day of gasoline boiling range hydrocarbons were produced, and about 4000 bbL/day of .middle-distillate, the combined sulfur content of these two streams being only about 0.3% by weight.
  • a butane-pentane concentrate may be recovered, where desired as a motor fuel blending component, in an amount of more than 320 bbl./day.

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  • Chemical & Material Sciences (AREA)
  • Oil, Petroleum & Natural Gas (AREA)
  • Engineering & Computer Science (AREA)
  • Chemical Kinetics & Catalysis (AREA)
  • General Chemical & Material Sciences (AREA)
  • Organic Chemistry (AREA)
  • Production Of Liquid Hydrocarbon Mixture For Refining Petroleum (AREA)
US598067A 1966-11-30 1966-11-30 Mixed-phase conversion product separation process Expired - Lifetime US3371029A (en)

Priority Applications (14)

Application Number Priority Date Filing Date Title
US598067A US3371029A (en) 1966-11-30 1966-11-30 Mixed-phase conversion product separation process
GB54558/67A GB1197469A (en) 1966-11-30 1967-11-30 Mixed-phase conversion product separation process
BE707354D BE707354A (es) 1966-11-30 1967-11-30
BR195153/67A BR6795153D0 (pt) 1966-11-30 1967-11-30 Processo para separacao de um produto de conversao de fase mista
TR16585A TR16585A (tr) 1966-11-30 1967-11-30 Karisik-faz konversiyon ueruenue ayirma usulue
GR670134824A GR34824B (el) 1966-11-30 1967-11-30 Μεθοδος διαχωρισμου προιοντος μετατροπης μικτης φασεως.
ES347843A ES347843A1 (es) 1966-11-30 1967-11-30 Un procedimiento para separar un fluido saliente de pro- ducto de conversion, en fase mixta y caliente.
CH1685567A CH542276A (de) 1966-11-30 1967-11-30 Verfahren zur Trennung eines heissen gemischtphasigen Konversionsproduktes
SE16469/67A SE345870B (es) 1966-11-30 1967-11-30
DE19671645825 DE1645825A1 (de) 1966-11-30 1967-11-30 Verfahren zur Trennung und Aufarbeitung eines heissen mischphasigen Kohlenwasserstoffumwandlungsprodukts
FR130317A FR1547873A (fr) 1966-11-30 1967-11-30 Procédé de séparation d'un produit de conversion à phases mélangées
OA53119A OA02549A (fr) 1966-11-30 1967-12-06 Procédé de séparation d'un produit de conversion à phases mélangées.
YU2395/67A YU32372B (en) 1966-11-30 1967-12-07 Postupak za izdvajanje koncentrovanog vodonika iz smese tecnih i gasovitih ugljovodonika
NL6717013A NL6717013A (es) 1966-11-30 1967-12-14

Applications Claiming Priority (3)

Application Number Priority Date Filing Date Title
US598067A US3371029A (en) 1966-11-30 1966-11-30 Mixed-phase conversion product separation process
OA53119A OA02549A (fr) 1966-11-30 1967-12-06 Procédé de séparation d'un produit de conversion à phases mélangées.
NL6717013A NL6717013A (es) 1966-11-30 1967-12-14

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BE (1) BE707354A (es)
BR (1) BR6795153D0 (es)
CH (1) CH542276A (es)
DE (1) DE1645825A1 (es)
ES (1) ES347843A1 (es)
FR (1) FR1547873A (es)
GB (1) GB1197469A (es)
NL (1) NL6717013A (es)
OA (1) OA02549A (es)
SE (1) SE345870B (es)
TR (1) TR16585A (es)
YU (1) YU32372B (es)

Cited By (18)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US3437707A (en) * 1967-09-12 1969-04-08 Universal Oil Prod Co Process for aromatic alkylation and olefinic oligomerization
US3437584A (en) * 1967-08-09 1969-04-08 Universal Oil Prod Co Method for converting heavy carbonaceous materials
US3437708A (en) * 1967-09-12 1969-04-08 Universal Oil Prod Co Separation process for alkylated aromatic compounds and olefinic oligomerizaton products
US3437706A (en) * 1967-09-12 1969-04-08 Universal Oil Prod Co Process for aromatic alkylation and olefinic oligomerization
FR2021706A1 (es) * 1968-10-28 1970-07-24 Universal Oil Prod Co
FR2021705A1 (es) * 1968-10-28 1970-07-24 Universal Oil Prod Co
FR2030066A1 (es) * 1968-10-28 1970-10-30 Universal Oil Prod Co
US3546099A (en) * 1969-02-26 1970-12-08 Universal Oil Prod Co Method for separating the effluent from a hydrocarbon conversion process reaction zone
US3549519A (en) * 1968-10-28 1970-12-22 Universal Oil Prod Co Mixed-phase thermal cracking process
US3666658A (en) * 1970-11-23 1972-05-30 Universal Oil Prod Co Hydroprocessing product separation
US4159937A (en) * 1978-08-30 1979-07-03 Uop Inc. Mixed-phase reaction product effluent separation process
US4159935A (en) * 1978-08-30 1979-07-03 Uop Inc. Conversion of hydrocarbonaceous black oils
EP0336484A1 (en) * 1988-03-31 1989-10-11 Shell Internationale Researchmaatschappij B.V. Process for separating hydroprocessed effluent streams
EP0665281A2 (en) * 1994-01-27 1995-08-02 The M.W. Kellogg Company Integrated distillate recovery process
EP3184607A1 (fr) 2015-12-23 2017-06-28 Axens Procédé d'hydrotraîtement ou d'hydroconversion avec striper et ballon séparateur basse pression sur la section de fractionnement
US20170183584A1 (en) * 2015-12-23 2017-06-29 Axens Installation and process for jointly implementing compression of the acid gases from the hydroconversion or hydrotreatment unit and that of the gaseous effluents from the catalytic cracking unit
US20170183581A1 (en) * 2015-12-29 2017-06-29 Uop Llc Process and apparatus for recovering hydrogen from hydroprocessed hot flash liquid
US10711205B2 (en) 2017-06-22 2020-07-14 Uop Llc Process for recovering hydroprocessed effluent with improved hydrogen recovery

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US7915457B2 (en) * 2005-09-26 2011-03-29 Symrise Gmbh & Co. Kg Intramolecular Prins reaction and catalysts suitable therefor

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US7915457B2 (en) * 2005-09-26 2011-03-29 Symrise Gmbh & Co. Kg Intramolecular Prins reaction and catalysts suitable therefor

Cited By (32)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US3437584A (en) * 1967-08-09 1969-04-08 Universal Oil Prod Co Method for converting heavy carbonaceous materials
US3437707A (en) * 1967-09-12 1969-04-08 Universal Oil Prod Co Process for aromatic alkylation and olefinic oligomerization
US3437708A (en) * 1967-09-12 1969-04-08 Universal Oil Prod Co Separation process for alkylated aromatic compounds and olefinic oligomerizaton products
US3437706A (en) * 1967-09-12 1969-04-08 Universal Oil Prod Co Process for aromatic alkylation and olefinic oligomerization
FR2030066A1 (es) * 1968-10-28 1970-10-30 Universal Oil Prod Co
FR2021705A1 (es) * 1968-10-28 1970-07-24 Universal Oil Prod Co
FR2021706A1 (es) * 1968-10-28 1970-07-24 Universal Oil Prod Co
US3549519A (en) * 1968-10-28 1970-12-22 Universal Oil Prod Co Mixed-phase thermal cracking process
US3546099A (en) * 1969-02-26 1970-12-08 Universal Oil Prod Co Method for separating the effluent from a hydrocarbon conversion process reaction zone
US3666658A (en) * 1970-11-23 1972-05-30 Universal Oil Prod Co Hydroprocessing product separation
JPS5130558B1 (es) * 1970-11-23 1976-09-01
US4159937A (en) * 1978-08-30 1979-07-03 Uop Inc. Mixed-phase reaction product effluent separation process
US4159935A (en) * 1978-08-30 1979-07-03 Uop Inc. Conversion of hydrocarbonaceous black oils
FR2434859A1 (fr) * 1978-08-30 1980-03-28 Uop Inc Procede de separation d'un effluent de produit de reaction en phase mixte
AU608961B2 (en) * 1988-03-31 1991-04-18 Shell Internationale Research Maatschappij B.V. Process for separating hydroprocessed effluent streams
US4925573A (en) * 1988-03-31 1990-05-15 Shell Internationale Research Maatschappij, B.V. Process for separating hydroprocessed effluent streams
EP0336484A1 (en) * 1988-03-31 1989-10-11 Shell Internationale Researchmaatschappij B.V. Process for separating hydroprocessed effluent streams
EP0665281A2 (en) * 1994-01-27 1995-08-02 The M.W. Kellogg Company Integrated distillate recovery process
US5453177A (en) * 1994-01-27 1995-09-26 The M. W. Kellogg Company Integrated distillate recovery process
EP0665281A3 (en) * 1994-01-27 1995-12-20 Kellogg M W Co Integrated process for the production of distillates.
RU2726528C2 (ru) * 2015-12-23 2020-07-14 Аксенс Процесс гидроочистки или гидроконверсии с использованием отпарной колонны и барабана-сепаратора низкого давления на участке фракционирования
US20170183574A1 (en) * 2015-12-23 2017-06-29 Axens Hydrotreatment of hydroconversion process with a stripper and a low pressure separator drum in the fractionation section
US20170183584A1 (en) * 2015-12-23 2017-06-29 Axens Installation and process for jointly implementing compression of the acid gases from the hydroconversion or hydrotreatment unit and that of the gaseous effluents from the catalytic cracking unit
CN106906002A (zh) * 2015-12-23 2017-06-30 阿克森斯公司 在分馏段中使用汽提塔和低压分离器鼓的加氢处理或加氢转化方法
US10563140B2 (en) * 2015-12-23 2020-02-18 Axens Installation and process for jointly implementing compression of the acid gases from the hydroconversion or hydrotreatment unit and that of the gaseous effluents from the catalytic cracking unit
EP3184607A1 (fr) 2015-12-23 2017-06-28 Axens Procédé d'hydrotraîtement ou d'hydroconversion avec striper et ballon séparateur basse pression sur la section de fractionnement
RU2730019C2 (ru) * 2015-12-23 2020-08-14 Аксенс Устройство и способ, осуществляющие совместное сжатие кислых газов с установки гидроконверсии или гидрообработки и газовых потоков с установки каталитического крекинга
CN106906002B (zh) * 2015-12-23 2021-04-02 阿克森斯公司 在分馏段中使用汽提塔和低压分离器鼓的加氢处理或加氢转化方法
US11028330B2 (en) * 2015-12-23 2021-06-08 Axens Hydrotreatment or hydroconversion process with a stripper and a low pressure separator drum in the fractionation section
US20170183581A1 (en) * 2015-12-29 2017-06-29 Uop Llc Process and apparatus for recovering hydrogen from hydroprocessed hot flash liquid
US10781380B2 (en) * 2015-12-29 2020-09-22 Uop Llc Process and apparatus for recovering hydrogen from hydroprocessed hot flash liquid
US10711205B2 (en) 2017-06-22 2020-07-14 Uop Llc Process for recovering hydroprocessed effluent with improved hydrogen recovery

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Publication number Publication date
SE345870B (es) 1972-06-12
YU32372B (en) 1974-10-31
OA02549A (fr) 1970-05-05
YU239567A (en) 1974-04-30
TR16585A (tr) 1973-01-01
FR1547873A (fr) 1968-11-29
DE1645825A1 (de) 1970-07-16
GB1197469A (en) 1970-07-08
CH542276A (de) 1973-11-15
BR6795153D0 (pt) 1973-02-20
NL6717013A (es) 1969-06-17
ES347843A1 (es) 1969-02-16
BE707354A (es) 1968-04-01

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