US3767562A - Production of jet fuel - Google Patents

Production of jet fuel Download PDF

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Publication number
US3767562A
US3767562A US00177362A US3767562DA US3767562A US 3767562 A US3767562 A US 3767562A US 00177362 A US00177362 A US 00177362A US 3767562D A US3767562D A US 3767562DA US 3767562 A US3767562 A US 3767562A
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zone
feed
process according
hydrogenation
temperature
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US00177362A
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English (en)
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M Sze
J Reilly
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Lummus Technology LLC
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Lummus Co
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    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J8/00—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes
    • B01J8/02—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes with stationary particles, e.g. in fixed beds
    • B01J8/04—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes with stationary particles, e.g. in fixed beds the fluid passing successively through two or more beds
    • B01J8/0446—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes with stationary particles, e.g. in fixed beds the fluid passing successively through two or more beds the flow within the beds being predominantly vertical
    • B01J8/0449—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes with stationary particles, e.g. in fixed beds the fluid passing successively through two or more beds the flow within the beds being predominantly vertical in two or more cylindrical beds
    • B01J8/0453—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes with stationary particles, e.g. in fixed beds the fluid passing successively through two or more beds the flow within the beds being predominantly vertical in two or more cylindrical beds the beds being superimposed one above the other
    • 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
    • C10G2400/00—Products obtained by processes covered by groups C10G9/00 - C10G69/14
    • C10G2400/08—Jet fuel

Definitions

  • ABSTRACT Temperature in the reactor is also controlled by the operation of the process.
  • No. 3,147,210 discloses the production of jet fuel by catalytic hydrogenation of high boiling aromatic hydrocarbons, preceded by a hydrofining or hydrodesulfurization step.
  • the feedstock is desulfurized in cocurrent flow with added hydrogen in the first stage, hydrogen sulfide is stripped after the first stage; the stripped liquid is then subjected to catalytic hydrogenation in countercurrent flow with hydrogenin a second stage.
  • a yet further object of this invention is to provide a method for producing a jet fuel with a low aromatics content. Additionally, it is an object of this invention to provide a method for producing a jet fuel which exceeds the minimum IPT Smoke Point of mm. Other objects of this invention will become apparent from the specification, drawings and claims hereof.
  • the invention contemplates the production of jet fuel from an aromatics-containing petroleum feedstock having a boiling range within the temperature range of from about F. to about 550F. comprising the steps of: (a) passing the feed-stock in cocurrent contact with a hydrogen-rich gas through a first reaction zone operated at a temperature of from about 250F. to about 575F. at elevated pressure in contact witha hydrogenation catalyst; (b) removing from said first reaction zone a gas phase effluent comprising hydrogen and vaporized liquid materials, and a partially hydrogenated liquid effluent; (c) passing said liquid effluent into a second reaction zone operated at a temperature of from about 200F. to about 500F.
  • FIGURE is a diagrammatic illustration of the process of this invention.
  • the hydrogenation zones are preferably contained in one hydrogenation vessel, which has theform of a vertical cylinder having disched ends and pressure sustaining walls.
  • the interior of the vessel is divided by horizontal partitions 12, 14, and 24, which are preferably perforated or foraminous plates or the like, into a plurality of chambers or zones including an upper reaction chamber 16, an intermediate vapor-disengaging zone 20, and a lower reaction chamber 18.
  • the reaction chambers 16 and 18 are packed with a suitable hydrogenation catalyst 22, which may be of any of the well known hydrogenation-dehydrogenation catalysts, including such as Raney nickel, or nickel, platinum or palladium, preferably on a support such as alumina, silica, kieselguhr, diatomaceous earth, magnesia, zirconia or other inorganic oxides, alone or in combination.
  • a suitable hydrogenation catalyst 22 which may be of any of the well known hydrogenation-dehydrogenation catalysts, including such as Raney nickel, or nickel, platinum or palladium, preferably on a support such as alumina, silica, kieselguhr, diatomaceous earth, magnesia, zirconia or other inorganic oxides, alone or in combination.
  • the catalyst in zone 16 is supported on partition 12.
  • the catalyst in zone 18 is supported on a similar partition 24. Partition 24 is preferably spaced somewhat above the bottom of the converter, thus defining the upper boundary of an additional lower chamber or zone 26
  • Fresh aromatics-containing feed such as is hereinafter described, is introduced into the system at line 46, into a hydrogen stream in line 40, and the mixture proceeds in line 40 as indicated by the arrows until it joins line 44, from which is added a condensed recycle liquid from separator 34.
  • the resulting mixture then passes through line 42 into the top of the hydrogenation vessel, at a temperature of from about 250F. to about 575F. and a pressure of from about 400 to about 1,500 psi, depending on the boiling range of the feedstock and the severity of the hydrogenation.
  • the lower temperature and pressure correspond to lower boiling feeds and lower severity of treatment.
  • the mixture of feed recycle liquid and hydrogen passes downwardly through the catalyst bed in zone 16, under adiabatic reaction conditions in which a substantial amount of the aromatics present in the total liquid charge are hydrogenated to the corresponding naphthenic compounds.
  • the reaction mixture which passes out of zone 16 is a two-phase mixture.
  • the liquid phase is a mixture of paraffins, naphthenes and some unreacted aromatics.
  • the gas phase effluent is a mixture of hydrogen, inert gaseous impurities, and vaporized liquid hydrocarbons of a composition generally similar to that of the liquid phase effluent.
  • the liquid phase of the effluent passes downwardly through the vapor-disengaging zone 20 into the second hydrogenation zone 18 (through partition 14, which serves as a distributor plate).
  • reaction chamber 18 hydrogen introduced through line 48 and passing through chamber 26 contacts the liquid phase effluent countercurrently, hydrogenating the remaining aromatics to the corresponding naphthenes.
  • the hydrogen is introduced without being preheated, at a relatively low temperature, compared to that of the liquid phase effluent from zone 16; generally the hydrogen temperature is no higher than about l-l20F.
  • the liquid portion which emerges from hydrogenation zone 18 is briefly accumulated in chamber 26 of the reactor, permitting disengagement of vapors and sealing the outlet to line 50 to prevent escape of hydrogen.
  • the liquid product is collected in line 50 and contains a very minor portion, generally less than 1.5 volume per cent, of residual unhydrogenated aromatics.
  • the gas phase effluent from hydrogenation zone 18 contains excess hydrogen, inert gaseous impurities, and vaporized hydrocarbons of a composition similar to those contained in the gas phase effluent from hydrogenation zone 16.
  • the gas phase effluents from both the first hydrogenation zone 16 and the second hydrogenation zone 18 collect in vapor-disengaging zone 20.
  • the combined gas phase fraction is withdrawn through line 28, and first passed through heat exchanger or waste heat boiler 52, in which some of the heat is used to produce steam for use in other processing steps, or'in other processes, or for general purposes.
  • the still hot vapor mixture is then passed through line 54, then preferably through condenser 30 in which it is used to preheat the mixture fed to the reactor, then through condenser 32, where the vaporized liquid phase components remaining in the system are recondensed to liquids.
  • the resulting two-phase system consisting of gaseous hydrogen, inert gases, and reliquefied hydrocarbons, is passed into separator 34, where the liquid and gaseous phases are separated.
  • the liquid phase is passed through line 44 to be mixed with the feed to hydrogenation zone 16 as previously described.
  • the gaseous phase comprising hydrogen and inert gases, may be vented partially, as through line 56, to prevent build-up of inert inpurities in the system.
  • Fresh feed hydrogen gas may be supplied from line 48 through line 58 into the recycle gas, in the event that the recycle hydrogen is insufficient to supply the needs in the first hydrogenation zone.
  • An important feature of this invention is a built-in temperature control. Reactions of the type contemplated are exothermic. The production of the desired jet fuel is favored by low outlet temperatures. Furthermore, runaway reactions must be prevented or coke and undesirable side products will be formed. Accordingly, external temperature control means are usually necessitated in processes for hydrogenating aromatics for jet fuel production. The present process, however, provides an inherent temperature control, particularly in the second hydrogenation zone 18. As the hydrogen feed from line 58 passes upwardly through this zone, a portion of the heat present in that chamber is absorbed in the process of sensibly heating the hydrogen.
  • the vaporized hydrocarbons recovered from the vapor-disengaging zone 20 and used as recycle comprise partially hydrogenated feed containing up to about 5 percent aromatics. Because of the low concentration of aromatics, the ratio of recycle to fresh feed is less than 1:], generally in the range of about 0.05:1 to about 0.75:1, and depends on a number of factors, including hydrogen partial pressure and purity, desired temperature in the reactor, etc.
  • the feed to the process comprises a petroleum fraction having a boiling range within the temperature range of from about F. to about 550F.
  • Fractions, for example, with boiling ranges such as l35F.-480F., 350F.510F. and 300F.520F. are typical of those within this broad range which are suitable feedstocks for this process.
  • the feed can be either a straight run or other petroleum fraction; such fractions as kerosenes, light and heavy naphthas, catalytically cracked cycle oils and furnace oils can be utilized.
  • Particularly suitable is a feedstock generally boiling within the kerosene boiling range, that is, boiling within the range of from about 300F. to about 550F.
  • the first hydrogenation zone 16 is operated at a temperature of from about 300F. to about 575F. and the second zone at about 250F. to about 500F., within the pressure ranges previously mentioned.
  • the process of this invention does not accomplish desulfurization for practical purposes; consequently most feedstocks should be desulfurized prior to being introduced into the process, generally in a separate unit (not shown).
  • the feed is desulfurized just prior to its admission into the first hydrogenation zone, it will generally be sufficiently hot that no further heating is required to bring it up to reaction temperature. If, however, the feed has been obtained from a simple fractionation process or has been allowed to cool down prior to being passed into this process, or has been in storage, preheating is required. In any case, the hydrogen fed to the first hydrogenation zone 16 must be preheated prior to its introduction into this zone. The liquid recycle to this zone must also be preheated.
  • the preheating of the hydrogen, and feed if necessary can be accomplished in a number of ways, and can be performed separately or together.
  • a convenient method, in this process is to utilize the heat contained in the vapors in lines 28 and 54, which have been removed from the vapor-disengaging zone 20.
  • the combined hydrogen (and feed, if necessary) in stream 42, together with recycle liquid from line 44, is passed through heat exchanger 30, in which it is preheated to the desired inlet temperature by indirect heat exchange with the partially cooled vapors in line 54.
  • This heat exchange under some conditions, may have the additional effect of partially condensing some of the hydrocarbons in the combined vapor stream, facilitating the separation of hydrocarbons for recycle from the hydrogen and other gases, in separator 34.
  • the fresh feed is already sufficiently hot so as not to require preheating, it should be by-passed around the preheater to avoid overheating and undesirable side reactions.
  • the fresh feed will then enter the system, for example, through line 43 instead of through line 46, or by-pass can be accomplished in other ways known in the art. In this case, only the hydrogen and recycled liquid hydrocarbons will be preheated.
  • the preheating of the fresh feed, liquid recycle and hydrogen can be done in separate heat exchangers, and the heated materials mixed before being introduced into the reactor.
  • This separate preheating can be done using any source of available heat, including the hot vapor mixture in line 54.
  • the ratio of hydrogen to fresh feed in the mixture fed to reaction zone 16 may vary from a stoichiometric ratio of 1 mole of hydrogen per double bond to as much as about 300 percent of the stoichiometric requirement, and the ratio of hydrogen to the liquid material entering reaction zone 18 may vary from about 0.3 to about 1.0 moles/mole.
  • the L.H.S.V. in the first zone 16 is preferably maintained between about 0.5 and about 6.0, based on fresh feed only, while that in the second zone 18 is generally at a higher level
  • the overall L.H.S.V. is maintained, however, between 0.5 and 6.0.
  • the temperature conditions in the second zone should be adjusted to maintain the temperature of the liquid product at the outlet between about 300F. and about 500F., depending on the boiling range of the fresh feed, to provide optimum conditions favoring hydrogenation of the aromatics to naphthenes and close equilibrium approach.
  • a process for producing jet fuels by the two-stage hydrogenation of a hydrocarbon feed having a boiling range within the temperature range of about 300F to about 550F, and substantially free of sulfur-containing impurities comprising the steps of:

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  • Chemical & Material Sciences (AREA)
  • Organic Chemistry (AREA)
  • Chemical Kinetics & Catalysis (AREA)
  • Production Of Liquid Hydrocarbon Mixture For Refining Petroleum (AREA)
  • Liquid Carbonaceous Fuels (AREA)
US00177362A 1971-09-02 1971-09-02 Production of jet fuel Expired - Lifetime US3767562A (en)

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US17736271A 1971-09-02 1971-09-02

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JP (1) JPS5442005B2 (enExample)
AU (1) AU469380B2 (enExample)
CA (1) CA983420A (enExample)
DE (1) DE2242331A1 (enExample)
ES (1) ES406277A1 (enExample)
FI (1) FI59810C (enExample)
FR (1) FR2151059B1 (enExample)
GB (1) GB1399108A (enExample)
IT (2) IT962225B (enExample)
NL (1) NL176583C (enExample)

Cited By (33)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US4427534A (en) 1982-06-04 1984-01-24 Gulf Research & Development Company Production of jet and diesel fuels from highly aromatic oils
US4849093A (en) * 1987-02-02 1989-07-18 Union Oil Company Of California Catalytic aromatic saturation of hydrocarbons
US5183556A (en) * 1991-03-13 1993-02-02 Abb Lummus Crest Inc. Production of diesel fuel by hydrogenation of a diesel feed
JPH06507455A (ja) * 1991-11-05 1994-08-25 リーター、インゴルシュタット、シュピナライ　マシーネンバウ、アクチェンゲゼルシャフト 紡糸機紡糸単位におけるスプールの直径を確認するための方法および装置
US5393408A (en) * 1992-04-30 1995-02-28 Chevron Research And Technology Company Process for the stabilization of lubricating oil base stocks
US5882505A (en) * 1997-06-03 1999-03-16 Exxon Research And Engineering Company Conversion of fisher-tropsch waxes to lubricants by countercurrent processing
US5888376A (en) * 1996-08-23 1999-03-30 Exxon Research And Engineering Co. Conversion of fischer-tropsch light oil to jet fuel by countercurrent processing
US5906728A (en) * 1996-08-23 1999-05-25 Exxon Chemical Patents Inc. Process for increased olefin yields from heavy feedstocks
US5928497A (en) * 1997-08-22 1999-07-27 Exxon Chemical Pateuts Inc Heteroatom removal through countercurrent sorption
US5942197A (en) * 1996-08-23 1999-08-24 Exxon Research And Engineering Co Countercurrent reactor
US6241952B1 (en) 1997-09-26 2001-06-05 Exxon Research And Engineering Company Countercurrent reactor with interstage stripping of NH3 and H2S in gas/liquid contacting zones
US6274029B1 (en) 1995-10-17 2001-08-14 Exxon Research And Engineering Company Synthetic diesel fuel and process for its production
US6309432B1 (en) 1997-02-07 2001-10-30 Exxon Research And Engineering Company Synthetic jet fuel and process for its production
US6495029B1 (en) 1997-08-22 2002-12-17 Exxon Research And Engineering Company Countercurrent desulfurization process for refractory organosulfur heterocycles
US6497810B1 (en) 1998-12-07 2002-12-24 Larry L. Laccino Countercurrent hydroprocessing with feedstream quench to control temperature
US6514403B1 (en) * 2000-04-20 2003-02-04 Abb Lummus Global Inc. Hydrocracking of vacuum gas and other oils using a cocurrent/countercurrent reaction system and a post-treatment reactive distillation system
US6569314B1 (en) 1998-12-07 2003-05-27 Exxonmobil Research And Engineering Company Countercurrent hydroprocessing with trickle bed processing of vapor product stream
US6579443B1 (en) 1998-12-07 2003-06-17 Exxonmobil Research And Engineering Company Countercurrent hydroprocessing with treatment of feedstream to remove particulates and foulant precursors
US6623621B1 (en) 1998-12-07 2003-09-23 Exxonmobil Research And Engineering Company Control of flooding in a countercurrent flow reactor by use of temperature of liquid product stream
US20040085154A1 (en) * 2001-07-09 2004-05-06 Stark Donald C. Methods for bi-directional signaling
US6822131B1 (en) 1995-10-17 2004-11-23 Exxonmobil Reasearch And Engineering Company Synthetic diesel fuel and process for its production
US20040238409A1 (en) * 2003-05-30 2004-12-02 Harjeet Virdi Hydrogenation of middle distillate using a counter-current reactor
US6835301B1 (en) 1998-12-08 2004-12-28 Exxon Research And Engineering Company Production of low sulfur/low aromatics distillates
US20050077635A1 (en) * 2003-08-18 2005-04-14 Van Hasselt Bastiaan Willem Distribution device
JP2006104271A (ja) * 2004-10-01 2006-04-20 Nippon Oil Corp 水素化精製軽油の製造方法、水素化精製軽油及び軽油組成物
WO2006069406A1 (en) * 2004-12-23 2006-06-29 The Petroleum Oil And Gas Corporation Of South Africa (Pty) Ltd A process for catalytic conversion of fischer-tropsch derived olefins to distillates
US20110220546A1 (en) * 2010-03-15 2011-09-15 Omer Refa Koseoglu High quality middle distillate production process
KR101070519B1 (ko) 2003-05-30 2011-10-05 에이비이비이 러머스 글로벌 인코포레이티드 역류 반응장치를 이용하는 중간유분의 수소첨가 방법
WO2011061612A3 (en) * 2009-11-20 2012-01-05 Total Raffinage Marketing Process for the production of hydrocarbon fluids having a low aromatic content
US9315742B2 (en) 2009-11-20 2016-04-19 Total Marketing Services Process for the production of hydrocarbon fluids having a low aromatic content
EP1836284B1 (en) * 2004-12-23 2018-08-22 The Petroleum Oil and Gas Corporation of South Afr. Synthetically derived distillate kerosene and its use
US10246652B2 (en) 2013-12-23 2019-04-02 Total Marketing Services Process for the dearomatization of petroleum cuts
US10385278B2 (en) 2011-02-15 2019-08-20 Neste Oyj Use of renewable oil in hydrotreatment process

Families Citing this family (6)

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CA1033680A (en) * 1973-07-12 1978-06-27 Lummus Company (The) Production of jet fuel
US8022258B2 (en) 2005-07-05 2011-09-20 Neste Oil Oyj Process for the manufacture of diesel range hydrocarbons
US12203035B2 (en) 2005-07-05 2025-01-21 Neste Oyj Process for the manufacture of diesel range hydrocarbons
US8303804B2 (en) * 2008-10-06 2012-11-06 Exxonmobil Research And Engineering Company Process to improve jet fuels
CN108291152B (zh) * 2015-11-23 2021-01-08 国际壳牌研究有限公司 生物质向液态烃物质的转化
BR112018010526B1 (pt) * 2015-11-23 2021-10-05 Shell Internationale Research Maatschappij B.V. Processo para produzir produtos de hidrocarboneto líquidos a partir de pelo menos uma dentre uma matéria-prima que contém biomassa e uma matéria-prima derivada de biomassa

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US2952626A (en) * 1957-08-05 1960-09-13 Union Oil Co Mixed-phase hydrofining of hydrocarbon oils
US3147210A (en) * 1962-03-19 1964-09-01 Union Oil Co Two stage hydrogenation process
US3450784A (en) * 1966-09-22 1969-06-17 Lummus Co Hydrogenation of benzene to cyclohexane
US3484496A (en) * 1965-10-04 1969-12-16 British Petroleum Co Desulphurisation and hydrogenation of aromatic hydrocarbons
US3513085A (en) * 1967-09-06 1970-05-19 Arnold M Leas Producing isoparaffins and naphthenes from hydrocarbons
US3573198A (en) * 1969-02-10 1971-03-30 Universal Oil Prod Co Smoke point improvement of jet fuel kerosene fractions

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US2952626A (en) * 1957-08-05 1960-09-13 Union Oil Co Mixed-phase hydrofining of hydrocarbon oils
US3147210A (en) * 1962-03-19 1964-09-01 Union Oil Co Two stage hydrogenation process
US3484496A (en) * 1965-10-04 1969-12-16 British Petroleum Co Desulphurisation and hydrogenation of aromatic hydrocarbons
US3450784A (en) * 1966-09-22 1969-06-17 Lummus Co Hydrogenation of benzene to cyclohexane
US3513085A (en) * 1967-09-06 1970-05-19 Arnold M Leas Producing isoparaffins and naphthenes from hydrocarbons
US3573198A (en) * 1969-02-10 1971-03-30 Universal Oil Prod Co Smoke point improvement of jet fuel kerosene fractions

Cited By (56)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US4427534A (en) 1982-06-04 1984-01-24 Gulf Research & Development Company Production of jet and diesel fuels from highly aromatic oils
US4849093A (en) * 1987-02-02 1989-07-18 Union Oil Company Of California Catalytic aromatic saturation of hydrocarbons
JP2617158B2 (ja) 1991-03-13 1997-06-04 エービービー ルーマス クレスト インコーポレイテッド ディーゼル原料の水素化によるディーゼル燃料の製造
US5183556A (en) * 1991-03-13 1993-02-02 Abb Lummus Crest Inc. Production of diesel fuel by hydrogenation of a diesel feed
JPH06507455A (ja) * 1991-11-05 1994-08-25 リーター、インゴルシュタット、シュピナライ　マシーネンバウ、アクチェンゲゼルシャフト 紡糸機紡糸単位におけるスプールの直径を確認するための方法および装置
US5393408A (en) * 1992-04-30 1995-02-28 Chevron Research And Technology Company Process for the stabilization of lubricating oil base stocks
US6296757B1 (en) 1995-10-17 2001-10-02 Exxon Research And Engineering Company Synthetic diesel fuel and process for its production
US6822131B1 (en) 1995-10-17 2004-11-23 Exxonmobil Reasearch And Engineering Company Synthetic diesel fuel and process for its production
US6607568B2 (en) 1995-10-17 2003-08-19 Exxonmobil Research And Engineering Company Synthetic diesel fuel and process for its production (law3 1 1)
US6274029B1 (en) 1995-10-17 2001-08-14 Exxon Research And Engineering Company Synthetic diesel fuel and process for its production
US5888376A (en) * 1996-08-23 1999-03-30 Exxon Research And Engineering Co. Conversion of fischer-tropsch light oil to jet fuel by countercurrent processing
US5906728A (en) * 1996-08-23 1999-05-25 Exxon Chemical Patents Inc. Process for increased olefin yields from heavy feedstocks
US5942197A (en) * 1996-08-23 1999-08-24 Exxon Research And Engineering Co Countercurrent reactor
US6149800A (en) * 1996-08-23 2000-11-21 Exxon Chemical Patents Inc. Process for increased olefin yields from heavy feedstocks
US6309432B1 (en) 1997-02-07 2001-10-30 Exxon Research And Engineering Company Synthetic jet fuel and process for its production
US6669743B2 (en) 1997-02-07 2003-12-30 Exxonmobil Research And Engineering Company Synthetic jet fuel and process for its production (law724)
US5882505A (en) * 1997-06-03 1999-03-16 Exxon Research And Engineering Company Conversion of fisher-tropsch waxes to lubricants by countercurrent processing
US6495029B1 (en) 1997-08-22 2002-12-17 Exxon Research And Engineering Company Countercurrent desulfurization process for refractory organosulfur heterocycles
US5928497A (en) * 1997-08-22 1999-07-27 Exxon Chemical Pateuts Inc Heteroatom removal through countercurrent sorption
US6241952B1 (en) 1997-09-26 2001-06-05 Exxon Research And Engineering Company Countercurrent reactor with interstage stripping of NH3 and H2S in gas/liquid contacting zones
US6579443B1 (en) 1998-12-07 2003-06-17 Exxonmobil Research And Engineering Company Countercurrent hydroprocessing with treatment of feedstream to remove particulates and foulant precursors
US6569314B1 (en) 1998-12-07 2003-05-27 Exxonmobil Research And Engineering Company Countercurrent hydroprocessing with trickle bed processing of vapor product stream
US6623621B1 (en) 1998-12-07 2003-09-23 Exxonmobil Research And Engineering Company Control of flooding in a countercurrent flow reactor by use of temperature of liquid product stream
US6497810B1 (en) 1998-12-07 2002-12-24 Larry L. Laccino Countercurrent hydroprocessing with feedstream quench to control temperature
US6835301B1 (en) 1998-12-08 2004-12-28 Exxon Research And Engineering Company Production of low sulfur/low aromatics distillates
US6514403B1 (en) * 2000-04-20 2003-02-04 Abb Lummus Global Inc. Hydrocracking of vacuum gas and other oils using a cocurrent/countercurrent reaction system and a post-treatment reactive distillation system
US20040085154A1 (en) * 2001-07-09 2004-05-06 Stark Donald C. Methods for bi-directional signaling
WO2004108637A3 (en) * 2003-05-30 2005-04-14 Abb Lummus Global Inc Hydrogenation of middle distillate using a counter-current reactor
CN100540635C (zh) * 2003-05-30 2009-09-16 路慕斯技术有限公司 使用逆流反应器的中间馏出物氢化方法
US20040238409A1 (en) * 2003-05-30 2004-12-02 Harjeet Virdi Hydrogenation of middle distillate using a counter-current reactor
US7247235B2 (en) 2003-05-30 2007-07-24 Abb Lummus Global Inc, Hydrogenation of middle distillate using a counter-current reactor
KR101070519B1 (ko) 2003-05-30 2011-10-05 에이비이비이 러머스 글로벌 인코포레이티드 역류 반응장치를 이용하는 중간유분의 수소첨가 방법
US20050077635A1 (en) * 2003-08-18 2005-04-14 Van Hasselt Bastiaan Willem Distribution device
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ES406277A1 (es) 1975-07-01
AU4563572A (en) 1974-02-21
FI59810B (fi) 1981-06-30
AU469380B2 (en) 1976-02-12
JPS4834906A (enExample) 1973-05-23
GB1399108A (en) 1975-06-25
DE2242331A1 (de) 1973-03-08
NL176583C (nl) 1985-05-01
JPS5442005B2 (enExample) 1979-12-12
FI59810C (fi) 1981-10-12
FR2151059B1 (enExample) 1976-08-13
IT962255B (it) 1973-12-20
CA983420A (en) 1976-02-10
FR2151059A1 (enExample) 1973-04-13
NL7211962A (enExample) 1973-03-06
IT962225B (it) 1973-12-20

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