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EP2379674A2 - Procédé adiabatique à plusieurs niveaux pour réaliser la synthèse fischer-tropsch - Google Patents

Procédé adiabatique à plusieurs niveaux pour réaliser la synthèse fischer-tropsch

Info

Publication number
EP2379674A2
EP2379674A2 EP09767977A EP09767977A EP2379674A2 EP 2379674 A2 EP2379674 A2 EP 2379674A2 EP 09767977 A EP09767977 A EP 09767977A EP 09767977 A EP09767977 A EP 09767977A EP 2379674 A2 EP2379674 A2 EP 2379674A2
Authority
EP
European Patent Office
Prior art keywords
reaction
reaction zone
zone
process gas
reaction zones
Prior art date
Legal status (The legal status is an assumption and is not a legal conclusion. Google has not performed a legal analysis and makes no representation as to the accuracy of the status listed.)
Withdrawn
Application number
EP09767977A
Other languages
German (de)
English (en)
Inventor
Ralph Schellen
Leslaw Mleczko
Evin Hizaler Hoffmann
Stephan Schubert
Rushikesh Apte
Current Assignee (The listed assignees may be inaccurate. Google has not performed a legal analysis and makes no representation or warranty as to the accuracy of the list.)
Bayer Intellectual Property GmbH
Original Assignee
Bayer Technology Services GmbH
Priority date (The priority date is an assumption and is not a legal conclusion. Google has not performed a legal analysis and makes no representation as to the accuracy of the date listed.)
Filing date
Publication date
Application filed by Bayer Technology Services GmbH filed Critical Bayer Technology Services GmbH
Publication of EP2379674A2 publication Critical patent/EP2379674A2/fr
Withdrawn legal-status Critical Current

Links

Classifications

    • 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
    • C10G2/00—Production of liquid hydrocarbon mixtures of undefined composition from oxides of carbon
    • C10G2/30—Production of liquid hydrocarbon mixtures of undefined composition from oxides of carbon from carbon monoxide with hydrogen
    • C10G2/32—Production of liquid hydrocarbon mixtures of undefined composition from oxides of carbon from carbon monoxide with hydrogen with the use of catalysts
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C1/00—Preparation of hydrocarbons from one or more compounds, none of them being a hydrocarbon
    • C07C1/02—Preparation of hydrocarbons from one or more compounds, none of them being a hydrocarbon from oxides of a carbon
    • C07C1/04—Preparation of hydrocarbons from one or more compounds, none of them being a hydrocarbon from oxides of a carbon from carbon monoxide with hydrogen
    • C07C1/0455—Reaction conditions
    • C07C1/046—Numerical values of parameters

Definitions

  • the present invention relates to a punctstuf ⁇ g adiabates process for carrying out the Fischer-Tropsch synthesis at low temperatures, wherein the synthesis is carried out in 5 to 40 successive reaction zones under adiabatic conditions.
  • Fischer-Tropsch synthesis is now a well-known chemical synthetic route that allows the production of hydrocarbons from carbon monoxide and hydrogen. More recently, this synthetic route has become increasingly important as the world's resources of natural hydrocarbons, such as petroleum or wax, are becoming increasingly scarce. However, carbon monoxide and hydrogen are available in greater quantities, or are by-products in connection with other processes, or can be selectively produced without necessarily resorting to the aforementioned natural resources.
  • the present invention relates to the low-temperature process variant. Differences between the two process variants in addition to the temperature, especially on the basis of the products obtained from the process variants. While predominantly shorter-chain, liquid or gaseous hydrocarbons are obtained in the high-temperature process variant, in the low-temperature process variant predominantly long-chain, liquid, sometimes highly viscous hydrocarbons are obtained.
  • the low-temperature process variant is generally advantageous in that the shorter-chain hydrocarbons of the high-temperature process variant can be obtained from the longer-chain hydrocarbons by the "cracking" process known to those skilled in the art, but at the same time the longer-chain hydrocarbons are starting materials for higher-value products, such as waxes , form.
  • a fundamental problem of the two process variants of the Fischer-Tropsch synthesis is the highly exothermic nature of the synthesis, so that high process requirements are placed on the removal of heat from the synthesis, if one wants to control the synthesis exactly.
  • the temperature has a strong influence on the product properties of the Fischer-Tropsch synthesis.
  • the process according to the disclosure of US 6,558,634 is disadvantageous in that the reaction zones are cooled over the walls of the chamber in which the reaction zone is located, thereby attempting to provide the necessary temperature control of the process. Cooling via a chamber wall causes a temperature profile to be established along the direction of flow of the process gases carbon monoxide and hydrogen in a reaction zone due to the exothermic nature of the reaction, since a constant heat flow is removed from the reaction zone via the surface of the chamber wall, which is above the temperature and the heat capacity, and possibly on the flow rate of the heat transfer medium must be controlled in the cooling chamber.
  • Cooled circulation stream Only part of the cycle stream is taken out of the process as a product.
  • Carbon monoxide is not disclosed. Furthermore, the process is one-stage with regard to the process gases.
  • WO 2005/075606 An alternative process variant, which also allows a certain temperature control, is disclosed in WO 2005/075606.
  • the temperature control is achieved by a large number of parallel microchannels and high surface to volume ratios. Again, the reaction chambers are in contact with cooling chambers through which a heat transfer medium flows. This results in principle from the analog disadvantages, as they are already in Related to the disclosure of US 6,558,634 have been set forth in this regard.
  • the temperature control of the method according to WO 2005/075606 is advantageous because, as just stated, the surface in contact with the process gases hydrogen and carbon monoxide is considerably larger compared to the volume of the respective reaction zone, so that the formation of a profile is not prevented, the dimensions of which can certainly be significantly reduced.
  • EP 1 251 951 discloses a device and the possibility of carrying out chemical reactions in the device, wherein the device is characterized by a cascade of reaction zones in contact with one another and heat exchanger devices which are arranged in a composite with one another. The method to be carried out here is thus characterized by the contact of the various reaction zones with a respective heat exchanger device in the form of a cascade.
  • EP 1 251 951 (B1) discloses multi-stage processes in cascades of reaction zones from which heat in an undefined amount is removed by heat conduction. Accordingly, the disclosed method is disadvantageous in that accurate temperature control of the process gases of the reaction is not possible.
  • Carbon monoxide in the context of the present invention refers to a process gas comprising essentially carbon monoxide.
  • the proportion of carbon monoxide in the process gas supplied to the process, called carbon monoxide is usually between 70 and 100% by weight, preferably between 80 and 100% by weight.
  • Hydrogen in the context of the present invention, denotes a process gas which essentially comprises hydrogen.
  • the proportion of hydrogen at the, the Method supplied process gas, called hydrogen between 90 and 100 wt .-%, preferably between 95 and 100 wt .-%.
  • this may also include secondary components.
  • secondary components Non-exhaustive examples of minor components which may be included in the process gas include argon, nitrogen and / or carbon dioxide. The same applies in the case of the process gas carbon monoxide.
  • the two process gases hydrogen and carbon monoxide are together with their respective minor components in the context of the present invention also collectively referred to as the process gas.
  • the process gas is thus understood as a gas mixture comprising hydrogen, carbon monoxide and secondary components.
  • Liquid hydrocarbons in the context of the present invention refer to aliphatic hydrocarbons which are present in the liquid phase reaction zones under the conditions of the process. Usually these are aliphatic hydrocarbons comprising at least nine carbon atoms, preferably comprising at least twelve carbon atoms.
  • adiabat means that no heat supply or removal measures are taken.
  • An advantage of the adiabatic driving method according to the invention of the 5 to 40 reaction zones connected in series with respect to a non-adiabatic mode of operation is that no means for heat removal must be provided in the reaction zones, which entails a considerable simplification of the construction. This results in particular simplifications in the manufacture of the reactor and in the scalability of the process and an increase in reaction conversions.
  • the heat generated in the course of the exothermic reaction progress can be utilized in the single reaction zone to increase the conversion in a controlled manner.
  • Another advantage of the method according to the invention is the possibility of very accurate temperature control by the close staggering of adiabatic reaction zones. It can thus be set and controlled in each reaction zone advantageous in the reaction progress temperature.
  • the catalysts used in the process according to the invention are usually catalysts which consist of a material which, in addition to its catalytic activity for the reaction of the formula (I), is characterized by sufficient chemical resistance under the conditions of the process and by a high specific surface area.
  • Catalyst materials characterized by such chemical resistance under the conditions of the process include, for example, catalysts comprising oxides of aluminum, titanium, zirconium and / or silicon and / or oxides of lanthanides. In most cases, these materials are carriers of the catalyst materials on which the active ingredients of the catalyst are applied.
  • Suitable active constituents of the catalysts are, for example, cobalt compounds, as are already known to the person skilled in the art and / or compounds which contain nickel, platinum and / or palladium. In preferred embodiments, these may also be doped with magnesium oxide as a promoter.
  • a high specific surface area is a specific surface area of at least 1 m 2 / g, preferably of at least 10 m 2 / g.
  • the catalysts of the invention are each in the reaction zones and can be used in all known forms, e.g. Fixed bed, fluidized bed, present.
  • the appearance is fixed bed.
  • the fixed bed arrangement comprises a catalyst bed in the strict sense, ie loose, supported or unsupported catalyst in any form and in the form of suitable packings.
  • catalyst bed as used herein also encompasses contiguous areas of suitable packages on a support material or structured catalyst supports. These would be, for example, to be coated ceramic honeycomb carrier or foams with comparatively high geometric surfaces or corrugated layers of metal wire mesh on which, for example, catalyst granules is immobilized.
  • a Special form of packing is considered in the context of the present invention, the presence of the catalyst in monolithic form.
  • the catalyst is preferably present in beds of particles having mean particle sizes of 1 to 10 mm, preferably 1.5 to 8 mm, particularly preferably 2 to 6 mm.
  • the catalyst is in a fixed bed arrangement in monolithic form.
  • the monolithic catalyst is provided with channels through which the process gases flow.
  • the channels have a diameter of 0.1 to 3 mm, preferably a diameter of 0.2 to 2 mm, more preferably from 0.5 to 1 mm.
  • the catalyst is preferably present in loose beds of particles, as have already been described in connection with the fixed bed arrangement.
  • Beds of such particles are advantageous because the particles have a high specific surface area of the catalyst material compared to the process gases and thus a high conversion rate can be achieved.
  • the mass transport limitation of the reaction by diffusion can be kept low.
  • the particles are not yet so small that disproportionately high pressure losses occur when the fixed bed flows through.
  • the ranges of the particle sizes given in the preferred embodiment of the process, comprising a reaction in a fixed bed, are thus an optimum between the achievable conversion and the pressure loss generated when carrying out the process. Pressure loss is coupled in a direct manner with the necessary energy in the form of pump and / or compressor performance, so that a disproportionate increase in the same would result in an uneconomical operation of the method.
  • the conversion takes place in 7 to 30, more preferably 10 to 20 reaction zones connected in series.
  • a preferred further embodiment of the method is characterized in that at least the process gas emerging from at least one reaction zone is subsequently passed through at least one heat exchange zone downstream of said reaction zone.
  • the process gas and the liquid hydrocarbons are passed through at least one of these reaction zone downstream heat exchange zone.
  • after each reaction zone there is at least one, preferably exactly one, heat exchange zone, through which at least the process gas leaving the reaction zone is passed, preferably together with the liquid hydrocarbons.
  • the reaction zones can either be arranged in a reactor or arranged divided into several reactors.
  • the arrangement of the reaction zones in a reactor leads to a reduction in the number of apparatuses used.
  • the individual reaction zones and heat exchange zones can also be arranged together in a reactor or in any combination of reaction zones with heat exchange zones in several reactors.
  • reaction zones and heat exchange zones are present in a reactor, then in an alternative embodiment of the invention there is a heat insulation zone between them, in order to be able to obtain the adiabatic operation of the reaction zone.
  • each of the series-connected reaction zones can be replaced or supplemented independently of one another by one or more reaction zones connected in parallel.
  • the use of reaction zones connected in parallel allows in particular their replacement or supplementation during ongoing continuous operation of the process.
  • Parallel and successive reaction zones may in particular also be combined with one another.
  • the process according to the invention particularly preferably has exclusively reaction zones connected in series.
  • the reactors preferably used in the process according to the invention can consist of simple containers with one or more reaction zones, as e.g. in Ullmann's Encyclopedia of Industrial Chemistry (Fifth, Completely Revised Edition, VoI B4, page 95-104, page 210-216), wherein in each case between the individual reaction zones and / or heat exchange zones heat insulation zones can be additionally provided.
  • the catalysts or the fixed beds thereof are in a conventional manner on or between gas and liquid permeable walls comprising the reaction zone of
  • Catalyst beds technical devices for uniform gas distribution can be attached. These can be perforated plates, bubble-cap trays, valve trays or other internals which, by producing a small but uniform pressure loss, cause a uniform entry of the process gas into the fixed bed.
  • the inlet temperature of the air entering the first reaction zone process gas from 10 to 260 0 C, preferably from 50 to 250 0 C, more preferably from 150 to 235 ° C.
  • the absolute pressure at the inlet of the first reaction zone is between 10 and 70 bar, preferably between 20 and 50 bar, more preferably between 25 and 40 bar.
  • the residence time of the process gas in all reaction zones together is between 10 and 700 s, preferably between 20 and 500 s, particularly preferably between 150 and 250 s.
  • the residence time of the liquid hydrocarbons in all reaction zones together is between 50 and 2500 s, preferably between 100 and 1500 s, more preferably between 600 and 900 s.
  • the process gas is preferably fed only before the first reaction zone. This has the advantage that the entire process gas can be used for the absorption and removal of the heat of reaction in all reaction zones. In addition, by such a procedure, the space-time yield can be increased, or the necessary catalyst mass can be reduced. However, it is also possible, before one or more of the reaction zones following the first reaction zone, to replenish the process gas as required. In addition, the temperature and the conversion can be controlled via the supply of the process gas hydrogen between the reaction zones.
  • the process gas is cooled after at least one of the reaction zones used, more preferably after each reaction zone.
  • the process gas is passed after exiting a reaction zone through one or more of the above-mentioned heat exchange zones, which are located behind the respective reaction zones.
  • These can be embodied as heat exchange zones in the form of heat exchangers known to the person skilled in the art, such as, for example, tube bundle, plate, annular groove, spiral, finned tube and / or microstructured heat exchangers.
  • the heat exchangers are preferably microstructured heat exchangers.
  • microstructured means that the heat exchanger for the purpose of heat transfer comprises fluid-carrying channels, which are characterized in that they have a hydraulic diameter between 50 ⁇ m and 5 mm. The hydraulic diameter is calculated as four times the flow cross-sectional area of the fluid-conducting channel divided by the circumference of the channel.
  • steam is generated during cooling of the process gas and / or liquid hydrocarbon in the heat exchange zones through the heat exchanger.
  • the heat exchangers including the heat exchange zones it is preferable in the heat exchangers including the heat exchange zones to carry out evaporation on the side of the cooling medium, preferably partial evaporation.
  • Partial evaporation referred to in the context of the present invention, an evaporation in which a gas / liquid mixture of a substance is used as a cooling medium and in which there is still a gas / liquid mixture of a substance after heat transfer in the heat exchanger.
  • the carrying out of evaporation is particularly advantageous because in this way the achievable heat transfer coefficient from / to the process gas on / from the cooling / heating medium becomes particularly high and thus efficient cooling can be achieved.
  • the execution of a partial evaporation is particularly advantageous because the absorption / release of heat by the cooling medium thereby no longer in a temperature change of the
  • Cooling medium results, but only the gas / liquid balance is shifted. This has the consequence that over the entire heat exchange zone, the process gas is cooled to a constant temperature. This in turn safely prevents the occurrence of radial
  • a mixing zone may also be provided upstream of the inlet of a reaction zone in order to standardize the radial temperature profiles possibly formed during cooling in the flow of the process gas by mixing transversely to the main flow direction.
  • reaction zones connected in series are operated at an average temperature increasing or decreasing from reaction zone to reaction zone.
  • the temperature of reaction zone to reaction zone can both rise and fall. This can be adjusted, for example, via the control of the heat exchange zones connected between the reaction zone. Further options for setting the average temperature are described below.
  • the thickness of the flow-through reaction zones can be chosen to be the same or different and results according to laws well-known in the art from the above-described residence time and the respectively enforced in the process amounts of process gas.
  • the mass flows of liquid hydrocarbons which can be used in the process according to the invention and which also give the amounts of the process gases hydrogen and carbon monoxide to be used are usually between 100 and 200 t / h, preferably between 10 and 170 t / h, particularly preferably between 70 and 100 t / h.
  • the maximum outlet temperature of the process gas from the reaction zones is usually in a range from 220 ° C. to 300 ° C., preferably from 240 ° C. to 280 ° C., more preferably from 250 ° C. to 260 ° C.
  • the control of the temperature in the reaction zones is preferably carried out by at least one of the following measures: dimensioning of the adiabatic reaction zone, control of heat dissipation between the reaction zones, addition of process gas between the reaction zones, molar ratio of the reactants / excess of hydrogen used, addition of inert gases, in particular nitrogen, carbon dioxide, before and / or between the reaction zones.
  • the composition of the catalysts in the reaction zones according to the invention may be identical or different. In a preferred embodiment, the same catalysts are used in each reaction zone. However, it is also advantageous to use different catalysts in the individual reaction zones. Thus, especially in the first reaction zone, when the concentration of hydrogen and carbon monoxide is still high, a less active catalyst can be used and in the further reaction zones the activity of the catalyst can be increased from reaction zone to reaction zone.
  • the control of the catalyst activity can also be carried out by dilution with inert materials or carrier material. Also advantageous is the use of a catalyst in the first and / or second reaction zone, which is particularly stable against deactivation at the temperatures of the process in these reaction zones.
  • FIG. 1 shows reactor temperature (T) and conversion of carbon monoxide (U) over a number of 12 reaction zones (S) with downstream heat exchange zones (according to Example 1).
  • FIG. 2 shows reactor temperature (T) and conversion of carbon monoxide (U) over a number of 16 reaction zones (S) with downstream heat exchange zones (according to Example 2).
  • a process gas consisting of hydrogen and carbon monoxide at a molar ratio of 2.15 to 1 flows into a first of a total of 12 packed catalyst beds from an alumina-supported monolithic cobalt catalyst having a channel diameter of 0.75 mm.
  • the process thus comprises 12 reaction zones.
  • the catalyst is in the reaction zones in each case to a proportion of 25 vol .-%. In each case there is a void fraction of 75% by volume per reaction zone in order to allow a free outflow of the liquid hydrocarbons formed.
  • Each after a reaction zone is a heat exchange zone in which the process gas and the liquid hydrocarbons formed are cooled at the same time before they enter the next reaction zone.
  • the absolute inlet pressure of the process gas directly in front of the first reaction zone is 30 bar.
  • the length of the fixed catalyst beds, ie the reaction zones, is chosen so that a uniform temperature profile over the reaction zones is achieved. The exact values are listed in Table 1.
  • the results are shown in FIG.
  • the individual reaction zones are listed on the x-axis, so that a spatial course of developments in the process is visible.
  • the left-hand y-axis indicates the temperature of the process gas, which is substantially identical to that of the liquid hydrocarbons.
  • the temperature profile across the individual reaction zones is shown as a thick, solid line.
  • On the right y-axis the total conversion of carbon monoxide is indicated.
  • the course of sales Carbon monoxide across the individual reaction zones is shown as a thin dashed line.
  • the inlet temperature of the process gas before the first reaction zone is about 220 ° C. Due to the exothermic reaction to liquid hydrocarbons under adiabatic conditions, the temperature in the first reaction zone rises to 255 ° C, before the process gas and the liquid hydrocarbons are cooled in the downstream heat exchange zone. The inlet temperature before the next reaction zone is again about 220 0 C. By exothermic adiabatic reaction, it rises again to about 255 ° C. The sequence of heating and cooling continues identically to the 12th reaction zone.
  • a process gas consisting of hydrogen and carbon monoxide having a molar ratio of 2.15 to 1 flows into a first of a total of 16 fixed catalyst beds of catalyst analogous to that of Example 1.
  • the process thus comprises 16 reaction zones.
  • the catalyst is analogous to that of Example 1 in the reaction zones in each case to a proportion of 25 vol .-%.
  • Each after a reaction zone is a heat exchange zone in which the process gas and the liquid hydrocarbons formed are cooled at the same time before they enter the next reaction zone.
  • the absolute inlet pressure of the process gas directly in front of the first reaction zone is again 30 bar.
  • the length of the fixed catalyst beds, that is to say the reaction zones, is chosen such that a substantially uniform temperature profile over the reaction zones is achieved. The exact values are listed in Table 2.
  • Table 2 Lengths of the reaction zones according to Example 2
  • the results are shown in FIG.
  • the individual reaction zones are listed on the x-axis, so that a spatial course of developments in the process is visible.
  • the left-hand y-axis indicates the temperature of the process gas, which is substantially identical to that of the liquid hydrocarbons.
  • the temperature profile across the individual reaction zones is shown as a thick, solid line.
  • On the right y-axis the total conversion of carbon monoxide is indicated.
  • the course of the conversion of carbon monoxide across the individual reaction zones is shown as a thin dashed line.
  • the inlet temperature of the process gas upstream of the first reaction zone is 230 ° C. Due to the exothermic reaction to liquid hydrocarbons under adiabatic conditions, the temperature in the first reaction zone rises to 255 ° C, before the process gas and the liquid hydrocarbons are cooled in the downstream heat exchange zone. The inlet temperature before the next reaction zone is again 230 0 C. By exothermic, adiabatic reaction, it rises again to about 255 ° C. The sequence of heating and cooling continues identically until the 15th reaction zone. After this, a somewhat lower cooling to only about 235 ° C before the 16th reaction zone is provided, since the small amount of residual hydrogen and carbon monoxide present can expect no more so strong adiabatic temperature increase.

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  • Chemical & Material Sciences (AREA)
  • Organic Chemistry (AREA)
  • Chemical Kinetics & Catalysis (AREA)
  • Oil, Petroleum & Natural Gas (AREA)
  • Engineering & Computer Science (AREA)
  • General Chemical & Material Sciences (AREA)
  • Organic Low-Molecular-Weight Compounds And Preparation Thereof (AREA)
  • Production Of Liquid Hydrocarbon Mixture For Refining Petroleum (AREA)
  • Catalysts (AREA)
  • Low-Molecular Organic Synthesis Reactions Using Catalysts (AREA)

Abstract

L'invention concerne un procédé adiabatique à plusieurs niveaux pour réaliser la synthèse Fischer-Tropsch à basses températures, la synthèse étant réalisée dans des conditions adiabatiques dans 5 à 40 zones de réaction couplées en série.
EP09767977A 2008-12-20 2009-12-04 Procédé adiabatique à plusieurs niveaux pour réaliser la synthèse fischer-tropsch Withdrawn EP2379674A2 (fr)

Applications Claiming Priority (2)

Application Number Priority Date Filing Date Title
DE102008064282A DE102008064282A1 (de) 2008-12-20 2008-12-20 Vielstufig adiabates Verfahren zur Durchführung der Fischer-Tropsch-Synthese
PCT/EP2009/008671 WO2010069486A2 (fr) 2008-12-20 2009-12-04 Procédé adiabatique à plusieurs niveaux pour réaliser la synthèse fischer-tropsch

Publications (1)

Publication Number Publication Date
EP2379674A2 true EP2379674A2 (fr) 2011-10-26

Family

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Application Number Title Priority Date Filing Date
EP09767977A Withdrawn EP2379674A2 (fr) 2008-12-20 2009-12-04 Procédé adiabatique à plusieurs niveaux pour réaliser la synthèse fischer-tropsch

Country Status (4)

Country Link
US (1) US8557880B2 (fr)
EP (1) EP2379674A2 (fr)
DE (1) DE102008064282A1 (fr)
WO (1) WO2010069486A2 (fr)

Families Citing this family (2)

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Publication number Priority date Publication date Assignee Title
GB201107070D0 (en) 2011-04-27 2011-06-08 Davy Process Techn Ltd FT process using can reactor
EP3218448A1 (fr) 2014-11-14 2017-09-20 SABIC Global Technologies B.V. Réacteur de fischer-tropsch à plusieurs étages et procédés de production d'hydrocarbures

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US2503356A (en) 1946-07-01 1950-04-11 Texas Co Method of synthesizing hydrocarbons and the like
US2680125A (en) 1948-11-02 1954-06-01 Metallgesellschaft Ag Catalytic hydrogenation of carbon monoxide
IT1283774B1 (it) 1996-08-07 1998-04-30 Agip Petroli Processo di fischer-tropsch con reattore a colonna a bolle multistadio
GB2322633A (en) 1997-02-28 1998-09-02 Norske Stats Oljeselskap Fischer-Tropsch reactor
US6451864B1 (en) 1999-08-17 2002-09-17 Battelle Memorial Institute Catalyst structure and method of Fischer-Tropsch synthesis
JP5224627B2 (ja) 2000-01-25 2013-07-03 メギット （ユーケー） リミテッド 熱交換器付き化学反応器
GB0027575D0 (en) 2000-11-10 2000-12-27 Sasol Tech Pty Ltd Production of liquid hydrocarbon roducts
US7084180B2 (en) 2004-01-28 2006-08-01 Velocys, Inc. Fischer-tropsch synthesis using microchannel technology and novel catalyst and microchannel reactor
GB2444055B (en) 2006-11-23 2011-11-23 Gtl F1 Ag Gas to liquids plant with consecutive Fischer-Tropsch reactors and hydrogen make-up
WO2008080357A1 (fr) 2006-12-29 2008-07-10 Accelergy Shanghai R & D Center Co., Ltd. Procédé d'analyse de transfert de matière de processus catalytique à haut rendement

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Title
See references of WO2010069486A3 *

Also Published As

Publication number Publication date
WO2010069486A8 (fr) 2011-06-16
US20120136076A1 (en) 2012-05-31
WO2010069486A3 (fr) 2011-04-21
WO2010069486A2 (fr) 2010-06-24
DE102008064282A1 (de) 2010-06-24
US8557880B2 (en) 2013-10-15

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