WO1990006297A1 - Production de methanol a partir d'une charge d'hydrocarbures - Google Patents

Production de methanol a partir d'une charge d'hydrocarbures Download PDF

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Publication number
WO1990006297A1
WO1990006297A1 PCT/US1989/005370 US8905370W WO9006297A1 WO 1990006297 A1 WO1990006297 A1 WO 1990006297A1 US 8905370 W US8905370 W US 8905370W WO 9006297 A1 WO9006297 A1 WO 9006297A1
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WIPO (PCT)
Prior art keywords
methanol
carbon
gas
steam
partial oxidation
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PCT/US1989/005370
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English (en)
Inventor
Joseph D. Korchnak
Michael Dunster
Alan English
Original Assignee
Davy Mckee Corporation
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Filing date
Publication date
Application filed by Davy Mckee Corporation filed Critical Davy Mckee Corporation
Publication of WO1990006297A1 publication Critical patent/WO1990006297A1/fr

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    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C29/00—Preparation of compounds having hydroxy or O-metal groups bound to a carbon atom not belonging to a six-membered aromatic ring
    • C07C29/15—Preparation of compounds having hydroxy or O-metal groups bound to a carbon atom not belonging to a six-membered aromatic ring by reduction of oxides of carbon exclusively
    • C07C29/151—Preparation of compounds having hydroxy or O-metal groups bound to a carbon atom not belonging to a six-membered aromatic ring by reduction of oxides of carbon exclusively with hydrogen or hydrogen-containing gases
    • C07C29/1516—Multisteps
    • C07C29/1518—Multisteps one step being the formation of initial mixture of carbon oxides and hydrogen for synthesis
    • C—CHEMISTRY; METALLURGY
    • C12—BIOCHEMISTRY; BEER; SPIRITS; WINE; VINEGAR; MICROBIOLOGY; ENZYMOLOGY; MUTATION OR GENETIC ENGINEERING
    • C12C—BEER; PREPARATION OF BEER BY FERMENTATION; PREPARATION OF MALT FOR MAKING BEER; PREPARATION OF HOPS FOR MAKING BEER
    • C12C11/00—Fermentation processes for beer
    • C12C11/02—Pitching yeast
    • Y—GENERAL TAGGING OF NEW TECHNOLOGICAL DEVELOPMENTS; GENERAL TAGGING OF CROSS-SECTIONAL TECHNOLOGIES SPANNING OVER SEVERAL SECTIONS OF THE IPC; TECHNICAL SUBJECTS COVERED BY FORMER USPC CROSS-REFERENCE ART COLLECTIONS [XRACs] AND DIGESTS
    • Y02—TECHNOLOGIES OR APPLICATIONS FOR MITIGATION OR ADAPTATION AGAINST CLIMATE CHANGE
    • Y02P—CLIMATE CHANGE MITIGATION TECHNOLOGIES IN THE PRODUCTION OR PROCESSING OF GOODS
    • Y02P20/00—Technologies relating to chemical industry
    • Y02P20/10—Process efficiency
    • Y02P20/129—Energy recovery, e.g. by cogeneration, H2recovery or pressure recovery turbines

Definitions

  • the present invention relates to the production of methanol from hydrocarbonaceous feedstock by a process which includes the partial oxidation of hydrocarbonaceous
  • feedstock to produce a synthesis gas containing hydroyen, carbon monoxide and carbon dioxide, which is further
  • Methanol has long been produced by reacting hydrogen with carbon monoxide and/or carbon dioxide in the presence of a catalyst according to the equation:
  • Hydrocarbonaceous feedstock such as natural gases recovered from sites near petroleum deposits, are convenient starting materials for the production of methanol.
  • natural gases contain, as their principal constituent, methane, with minor amounts of ethane, propane and butane. Also included in the conversion in some instances, may be low-boiling liquid hydrocarbons.
  • the feedstock is first converted into a synthesis gas containing hydrogen, carbon monoxide and carbon dioxide.
  • the synthesis gas can be treated to adjust the ratio of hydrogen to carbon monoxide and carbon dioxide in order to obtain the proper stoichiometric proportions for methanol synthesis.
  • the treated gas stream is compressed and fed to the methanol synthesis loop, where the carbon monoxide, carbon dioxide and hydrogen are reacted in conta with a catalyst to- produce methanol.
  • the methanol is then purified by any of a number of conventional techniques. Methods that have been employed to convert
  • hydrocarbonaceous feedstock to synthesis gas in prior art methanol production processes include steam reforming, combined steam reforming/autothermal reforming, and partial oxidation.
  • Steam reforming involves an endothermic reaction exemplified for methane by the equations
  • reforming furnace is dependent on the steam-to-feedstock ratio at the inlet and the temperature and. pressure at the outlet of the reforming furnace.
  • the endothermic heat of reaction is supplied by firing burners adjacent the catalyst-filled tubes in the refractory lined reforming furnace. Waste heat in the flue gas is recovered by raising and superheating steam and preheating combustion air. After heat recovery and final cooling, the synthesis gas is compressed in a centrifugal compressor to between 50 and 100 bar .
  • the fresh synthesis gas joins the gases circulating in the methanol synthesis loop at the suction the circulation compressor. From the discharge of the circulation compressor, the bulk of the circulating gases are preheated to the reaction temperature of 210° to 270° and fed to a catalytic methanol synthesis reactor.
  • Gases leaving the reactor are used to preheat the circulating gas being fed to the reactor before crude methanol is condensed in a cooler and separated in a
  • This excess hydrogen has to be compressed to methanol loop pressure and then purged from the loop to be used as fuel in the reforming furnace.
  • Steam reforming can be supplemented by autother reforming to produce a stoichiometric synthesis gas
  • Desulfurized natural gas feedstock is split into two
  • the first stream is mixed with steam and
  • the primary reformer As excess methane is oxidized in the downstream secondary reformer, the primary reformer can be operated at higher pressure, lower temperature and lower steam to gas ratios than in the case of the single primary reforming furnace route.
  • the synthesis gas from the primary steam reformer is mixed with the second stream of natural gas feedstock and introduced into the secondary reformer with preheated oxygen. Heat of reaction in the secondary reformer is supplied by combustion of methane, hydrogen and carbon monoxide.
  • temperature of the secondary reformer is typically
  • Waste heat in the hot synthesis gas is
  • the size of the primary roforming furnace can be reduced by as much as 75%.
  • the total natural gas usage (feedstock plus fuel) is reduced by approxiinately 6% due to reduced fuel requirement in the primary reforming furnace and reduced feedstock requirements as a result of more efficient utilizatior. of the carbon contained in the feedstock.
  • the capital ccst is also reduced when compared to the conventional route as a result of the reduction in primary reforming furnace size, gas volumetric flow rates and compression power.
  • the autothermal reforming step requires a relative large volume of catalyst. Typically, space velocity
  • space velocity means the volumetric hourly
  • Partial oxidaticn of hydrocarbonaceous feedstock represents one alternative to steam reforming in the
  • the feedstock is compressed to approximately 30-80 bar, heated and introduced to a partial oxidation generator. Preheated oxygen is injected into the generator burner. It has been reported that the feedstock is converted to carbon rich synthesis gas
  • Heat contained in the synthesis gas leaving the generator at approximately 1400°C is recovered by raising steam before the gas is passed to a carbon scrubber where free carbon is removed.
  • Conversion efficiency of oxidation processes can generally be improved by the use of catalysts; but where the oxidation process in only partial, i.e. with insufficien oxygen to completely oxidize the hydrocarbon, then the catalyst is subject to carbon deposit and blockage. Carbon deposits can be avoided by using expensive catalyst
  • rhodium catalyst enables partial oxidation without causing deposition of carbon, but at temperatures greater than 900°C, thermal decomposition ensues producing ethylene or acetylene
  • a "LHSV" Liquid Hourly Space Velocity from 0.5 to 25 1/hour is disclosed; particularly, a high yield from partial oxidation of gasoline vapor, without steam, is produced at a temperature of 725°C and a LHSV of 20, and with steam, is produced at temperatures of 700°C and 800°C and at a LHSV of 2.
  • hydrocarbonaceous feedstock with a relatively low oxygen demand, thereby increasing throughput of hydrocarbonaceous feed.
  • the invention provides a process for producing methanol from hydrocarbonaceous feedstock in a manner which uses relatively smaller and less costly equipment and operates at a relatively higher level of efficiency, in terms of feedstock conversion, than prior art processes for methanol production.
  • This invention provides a process for the production of methanol in which synthesis gas is generated by the catalytic partial oxidation of a hydrocarbonaceous feedstock, such as natural gas, with an oxidant stream under temperature and steam conditions producing essential no free carbon at a space velocity in the range from 20,000 hour -1 to 500,000 hour -1 ; hydrogen, carbon monoxide and carbon dioxide in the synthesis gas are reacted under
  • the ratio of hydrogen to carbon monoxide and carbon dioxide in the synthesis gas is adjusted by removal of carbon monoxide and/or carbon dioxide to provide the proper stoichiometric amounts of reactants for the methanol production reaction.
  • the invention provides a process for producing methanol from hydrocarbonaceous feedstock which comprises:
  • FIG. 1 is an elevated cross-section view of a catalytic partial oxidation reactor having at its input a mixer and distributor suitable for introducing the reactants to the catalyst bed.
  • FIG. 2 is an enlarged elevational cross-section view of broken-away portion of the mixer and distributor o f FIG . 1 .
  • FIG. 3 is a top view of a broken-away quarter section of the mixer and distributor of FIG. 1.
  • FIG. 4 is a bottom view of a broken-away quarter section of the mixer and distributor of FIG. 1.
  • FIG. 5 is a diagrammatic elevational
  • FIG. 6 is a graph plotting ⁇ xygen-to-carbon molar ratio vs. steam-to-carbon molar ratio in the catalytic partial oxidation reaction for three different operating temperatures at an operating pressure of 400 psig (2760 KPa).
  • FIG. 7 is a graph plotting the hydrogen-to-carbon monoxide molar ratio in the catalytic partial oxidation reaction product vs. the steam-to-carbon molar ratio for three different operating temperatures at an operating pressure of 400 psig (2760 KPa).
  • FIG. 8 is the graph plotting the volume % methane in the catalytic partial oxidation product vs. the
  • FIG. 9 is a graph plotting the volume % carbon dioxide in the catalytic partial oxidation product vs.
  • FIG. 10 is a graph plotting the molar ratio of total hydrogen and carbon monoxide in the catalytic partial oxidation product to total hydrogen and carbon vs. steam- to-carbon molar ratio in the feedstock for three different: operating temperatures at an operating pressure of 400 psig (2760 KPa).
  • FIG. 11 is a flow diagram of a process for producing methanol from a hydrocarbonaceous feedstock in accordance with the invention
  • FIG. 12 is a sectional view of a tube cooled converter for use in producing methanol in accordance with the invention.
  • FIG. 13 is a process flow diagram of a first portion of a large methanol plant in accordance with the invention.
  • FIG. 14 is a process flow diagram of a second portion of the large methanol plant of FIG. 13.
  • FIG. 15 is a process flow diagram of a third portion of the large methanol plant of FIG. 13.
  • FIG. 16 is a process flow diagram of a first portion of a small methanol plant in accordance with the invention.
  • FIG. 17 is a process flow diagram of a second portion of the small methanol plant of FIG. 16. DESCRIPTION OF THE PREFERRED EMBODIMENTS
  • the process of the invention involves three steps: conversion of the hydrocarbonaceous feedstock into synthesis gas containing hydrogen, carbon dioxide and carbon monoxide by catalytic partial oxidation under
  • hydrocarbonaceous feedstock is carried out according to a process described in copending, commonly assigned U.S.
  • One particular aspect of the invention is the substantial capital cost savings and/or advantageous
  • catalytic partial oxidation performed at a temperature, as measured at the exit of the catalytic reaction zone, equal to or greater than a minimum non- carbon-forming temperature equal to or greater than a minimum temperature selected as a linear function which includes a range from 1600°F (870°C) to 1900°F (1040°C) corresponding to a range of the steam-to-carbon molar ratio from 0.4:1 to 0:1 and at a space velocity in the range from 20,000 hour -1 to 500,000 hour- 1 produces essentially no free carbon deposits on the catalyst.
  • products of the partial catalytic oxidation in the process of the invention consist essentially of hydrogen, carbon monoxide and carbon dioxide at oxidation temperatures equal to or greater than the minimum temperature, rhodium
  • catalysts are not required to prevent carbon formation.
  • dotted line 25 represents a generally linear function which, at a steam/carbon ratio of 0, corresponds to a minimum partial oxidation temperature of about 1900°F (1040°C), and at a steam/carbon ratio of 0.4 corresponds to a minimum partial oxidation temperature of about 1600°F (870°C); favorable catalytic partial oxidation without producing free carbon occurs at temperatures and steam/carbon ratios equal to or greater than points on the line. Further, lower minimum temperatures at corresponding steam/carbon ratios greater than 0.4 can be extrapolated from the linear function represented by line 25.
  • the catalytic partial oxidation is performed at a temperature, as measured at the exit of the catalyst, in the range from 1400°F (760°C) to 2300°F
  • the catalytic partial oxidation temperature is in the range from 1600°F (870°C) to 2000°F (1090°C). At temperatures below about 1400°F (760°C), uneconomic quantities of methane are left unconverted, and at temperatures above 2300°F (1260°C), excessive amounts of oxygen are used.
  • the pressure at which the partial oxidation takes place is generally above 150 psig (1030 KPa) and preferably above 300 psig (2060 KPa).
  • the pressure can be up to or above the methanol synthesis loop pressure.
  • the pressure does not exceed the methanol synthesis loop pressure by more than that necessary to provide synthesis gas flow through the processing equipment into the methanol synthesis loop.
  • the process of the invention can be employed to convert hydrocarbonaceous feedstock to synthesis gas with very low levels of hydrocarbon slippage (unreacted
  • the process of the invention can be carried out efficiently using relatively small volumes and relatively inexpensive catalyst materials, provided that the surface area-to-volume requirements of the invention are met.
  • the reactant gases are introduced to the reaction zone, i.e. the catalyst bed, at an inlet temperature not lower than 200°F (93°C) less than the autoignition temperature of the feed mixture.
  • the autoignition temperature of the feed depends on the composition and conditions of the feed mixture and on the catalyst employed.
  • the reactant gases are introduced at a temperature at or above the autoignition temperature of the mixture.
  • reactants shall be completely mixed prior to the reaction taking place. Introducing the thoroughly mixed reactant gases at the proper temperature ensures that the partial oxidation reactions will be mass transfer controlled.
  • reaction rate is relatively independent of catalyst activity, but dependent on the
  • the catalytic partial oxidation step can be understood with reference to the figures.
  • the catalytic partial oxidation zone is typically the catalyst bed of a reactor such as that illustrated in FIG. 1. As shown in FIG. 1, a reactor for partially oxidizing a gaseous
  • feedstock includes an input mixing and distributor section indicated generally at 30.
  • the mixer and distributor 30 mixes the feedstock with an oxidant and. distributes the mixture to the entrance of a catalytic reactor section indicated generally at 32 wherein the feedstock is partially oxidized to produce a product which is then passed through the exit section indicated generally at 34.
  • the reactor includes an outer shell 40 of
  • ⁇ layer 43 of insulation such as 2300°F (1260°C) BPCF ceramic fiber insulation, is secured to the inside of the upper portion of the shell 40 including the top 42.
  • layers 46, 48 and 50 are secured on the inside of the shell.
  • the layer 46 is a castable or
  • the layer 48 is also a castable or equivalent layer of insulation but containing 60% alumina for
  • the internal layer 50 is a refractory or equivalent layer such as 97% alumina with ceramic anchors or 97% alumina brick for withstanding the interior environment of the reactor section.
  • the catalytic reactor section 32 contains one or more catalyst discs 54. As shown, the reactor contains a sequence of discs 54 separated by high alumina rings 58 between each adjacent pair of discs. The stack is support by a grill with high alumina bars 56.
  • a sample port 60 is formed in the lower end of the reaction section and has a tube, such as type 309 stainless steel tube 62 extending below the bottom refractory disc 54 for withdrawing samples of the product.
  • the outlet section 34 is suitably formed for being connected to a downstream heat recovery boiler (not shown) and/or other processing equipment.
  • the catalyst comprises a high surface area material capable of catalyzing the partial oxidation of the hydrocarbonaceous feedstock.
  • the catalyst is in a
  • the catalyst has a geometric surface area to volume ratio of at least 20 cm 2 /cm 3 . While there is no strict upper limit of surface area to volume ratio, it normally does not exceed about 40 cm 2 /cm 3 .
  • the catalyst configuration normally considered to have catalytic activity, provided that the catalyst configuration has the desired surface area to volume ratio.
  • the catalyst disc 54 can be, for example, a monolithic structure having a honeycomb type cross-sectional configuration. Suitable monolithic structures of this type are produced commercially, in sizes smaller than those used in the process of the invention, as structural substrates for use in the catalytic conversion of automobile exhausts and as catalytic combustion chambers of gas turbines or for catalytic oxidation of waste streams.
  • the monolithic structure is an extruded material containing a plurality of closely packed channels running through the length of the structure to form a honeycomb structure. The channels are typically square and may be packed in a density as high as 1,200 per square inch of cross section.
  • the monolithic structure can be constructed of any of a variety of materials, including cordierite (MgO/Al 2 O 3 /SiO 2 ), Mn/Mgo cordierite (Mn-MgO/Al 2 O 3 /SiO 2 ), mullite (Al 2 O 3 /SiO 2 ), mullite aluminum titanate (AlaO 3 /SiO 2 -(Al,Fe) 2 O 3 /TiO 2 ), zirconia spinel (ZrO 2 /MgO/Al 2 O 3 ), spinel (MgO/Al 2 O 3 ), alumina (Al 2 O 3 ) and high nickel alloys.
  • the monolithic catalyst may consist solely of any of these structural materials, even though these materials are not normally considered to have catalytic activity by themselves.
  • the monolithic substrate can be coated with any of the metals or metal oxides known to have activity as oxidation catalysts. These include, for example, palladium, platinum, rhodium, iridium, osmium, ruthenium, nickel, chromium, cobalt, cerium, lanthanum and mixtures thereof.
  • Other metals which can be used tc coat the catalyst disc 54 include noble metals and metals of groups IA, IIA, III, IV, VB, VIB, or VIIB of the periodic table of elements.
  • the catalyst discs 54 may also consist of structural packing materials, such as that used in packing absorption columns. These packing materials generally comprise thin sheets of corrugated metal tightly packed together to form elongate channels running therethrough.
  • the structural packing materials may consist of corrugated sheets of metals such as high temperature alloys, stainless steels, chromium, manganese, molybdenum and refractory materials. These materials can, if desired, be coated with metals or metal oxides known to have catalytic activity for the oxidation reaction, such as palladium, platinum, rhodium, iridium, osmium, ruthenium, nickel, chromium, cobalt, cerium, lanthanum and mixtures thereof.
  • metals or metal oxides known to have catalytic activity for the oxidation reaction, such as palladium, platinum, rhodium, iridium, osmium, ruthenium, nickel, chromium, cobalt, cerium, lanthanum and mixtures thereof.
  • the catalyst discs 54 can also consist of dense wire mesh, such as high temperature alloys or platinum mesh. If desired, the wire mesh can also be coated with a metal or metal oxide having catalytic activity for the oxidation reaction, including palladium, platinum, rhodium, iridium, osmium, ruthenium, nickel, chromium cobalt, cesium, lanthanum and mixtures thereof.
  • a metal or metal oxide having catalytic activity for the oxidation reaction including palladium, platinum, rhodium, iridium, osmium, ruthenium, nickel, chromium cobalt, cesium, lanthanum and mixtures thereof.
  • the surface area to volume ratio of any of the aforementioned catalyst configurations can be increased coating the surfaces thereof with an aqueous slurry
  • particulate metal or metal oxide such as alumina, or metals of groups 1A, III, IV VB, VIB and VIIB and firing the coated surface at high temperature to adhere the particulate metal to the surface, but not so high as to cause sintering of the surface.
  • the particles employed should have a BET (Brunnauer-Emmett-Teller) surface area greater than about 10 m 2 /gram, preferably greater than about 200 m 2 /gram.
  • steam is introduced into the catalytic partia] oxidation zone at a temperature not lower than 200°F (93°C) below its autoignition temperature.
  • the gaseous mixture enters the catalytic partial oxidation zone at a temperature equal to or greater than its autoignition temperature. It is possible to operate the reactor in a mass transfer controlled mode with the reactants entering the reaction zone at a temperature somewhat below the autoignition temperature since the heat of reaction will provide the necessary energy to raise the reactant
  • reaction zone a temperature within the reaction zone.
  • the reactant mixture When the reactant mixture enters the catalytic partial oxidation zone at a temperature exceeding its autoignition temperature, it is necessary to introduce the mixture to the catalyst bed immediately after mixing; that is, the mixture of hydrocarbonaceous feedstock and oxidant should preferably be introduced to the catalyst bed before the autoignition delay time elapses. It is also essential that the gaseous reactants be thoroughly mixed. Failure to mix the reactants thoroughly reduces the quality of the product and can lead to overheating.
  • one of the feed gases i.e. hydrocarbonaceous gas or oxygen-containing gas.
  • the input section 30 is introduced into the input section 30 through a first inlet port 66 through the top 42 which communicates to an upper feed cone 68 which forms a first chamber.
  • the cone 68 is fastened by supports 69 in the top 42.
  • the other feed gas is introduced into the input section 30 through second inlets 70 extending through side ports of the shell 40 and communicating to a second chamber 72 which is interposed between the upper chamber 68 and the inlet of the catalyst reaction section 32.
  • a ring 73 mounted on the central portion of an upper wall 75 of the chamber 72 sealingly engages the lower edge of the cone 68 so that the wall 82 forms a common wall between the upper chamber 68 and lower chamber 72.
  • the chamber 72 has an upper outer annular portion 74, see also Figs.
  • a lower portion of the chamber 72 has a tubular wall 76 which extends downward in the refractory sleeve 50.
  • the bottom of the chamber 76 is formed by a cast member 78.
  • steam can be introduced into either both of the hydrocarbonaceous feedstock and oxygen
  • oxygen-containing gas The gases are fed to the reactor in relative proportions such that the steam-to-carbon molar ratio is in the range from 0:1 to 3.0:1, preferably from 0.8:1 to 2.0:1.
  • the oxygen-to-carbon ratio is from 0.4:1 to 0.8:1, preferably from 0.45 to 0.65.
  • air may be used as the oxidant in the process of the invention, it is preferred to use oxygen or an oxygen-rich gas, in order to minimize the inert ingredients such as nitrogen that must be carried through the system.
  • oxygen-rich is meant a ga s containing at least 70 mole.% oxygen, preferabJy at loast 30 mole.% oxygen.
  • the reactant mixture preferably enters the
  • this will generally be between about 550°F (290°C) and 1,100°F (590°C).
  • hydrocarbonaceous feedstock and steam are admixed and heated to a temperature from 650°F (340°C) to
  • Oxygen or oxygen-containing gas such as air, is heated to a temperature from 150°F (65°C) to 1200°F (650°C) and passes through the other inlet port(s) 66 or 70.
  • the mixing and distributing means comprises a plurality of elongated tubes 80 having upper ends mounted in the upper wall 75 of the chamber 72.
  • the lumens of the tubes at the upper end communicate with the upper chamber 68.
  • the bottom ends of the tubes 80 are secured to the member 78 with the lumens of the tubes communicating with the upper ends of passageways 84 formed vertically through the member 78.
  • Orifices 86 are formed in the walls of the tvibes 80 for directing streams of gas from the chamber 72 into the lumens of the tubes 80.
  • the number of tubes 80, the internal diameter 90 (see FIG. 5) of the tubes 80, the size and number of the orifices 86 in each tube are selected relative to the gas input velocities and pressures through inlets 66 and 70 so as to produce turbulent flow within the tubes 80 at a
  • passageways 84 is selected to be equal to or greater than that required for providing substantially complete mixing of the gas streams from chambers 68 and 72 under the conditions of turbulence therein.
  • the size of the internal diameter 90 of the tubes 80 as well as the length 94 of the tubes is designed to produce a sufficient pressure drop in the gas passing from the chamber 68 to the reaction chamber so as to provide for substantially uniform gas flow through the tubes 80 from the chamber 68.
  • the size of the orifices 86 is selected to provide sufficient pressure drop between the chamber 72 and the interior of the tubes 80 relative to the velocity and pressures of the gas entering through inlets 70 so as to provide substantially uniform volumes of gas flows through the orifices 86 into the tubes 80.
  • the diverging passageways in the member 78 are formed in a manner to provide for reduction of the velocity of the gas to produce uniform gas distribution over the inlet of the catalyst. The rate of increase of the
  • the angle 98 that the wall of the passageway 04 makes with the straight wall of the tubes 80 must generally be equal to or less than about 15° and preferably equal to or less than 7° in order to minimize or avoid creating vortices within the passageways 84. This assures that the
  • the gas exiting the catalytic partial oxidation reactor contains hydrogen, carbon dioxide, carbon monoxide and some methane.
  • the synthesis gas leaving the catalyst zone is first cooled by heat exchange, either by heating the hydrocarbon and steam feed stock, by heating the oxidant stream, by super heating steam, by raising steam in a boiler, by preheating boiler feed water or any combination thereof.
  • the gas stream that is converted to methanol in the methanol synthesis loop contains hydrogen and carbon oxides, i.e. carbon dioxide and carbon monoxide.
  • the molar ratio of hydrogen to carbon oxides in the gas to be
  • MSSGR methanol stoichiometric synthesis gas ratio
  • the MSSGR should have a value of at least 0.8 and preferably from about 0.95 to 1.1 for methanol synthesis.
  • the synthesis gas from the catalytic partial oxidation reactor will be slightly carbon-rich, that is, the MSSGK is somewhat lower than that ideally desired.
  • carbon is removed, either from the synthesis gas passing to the methanol synthesis loop, the methanol loop or the purge gas from the methanol synthesis loop.
  • Carbon dioxide can be removed from the synthesis gas stream by any know method.
  • the gas stream can be passed through a countercurrent liquid stream of a carbon dioxide absorbing medium.
  • Commercial processing units for carbon dioxide removal a re available, for example, under the trademarks Selexol, Amine Guard, and Benfieid
  • any known method for carbon monoxide removal can be employed.
  • One suitable method involves converting at least a portion of the carbon monoxide to carbon dioxide by water gas shift reaction and then removing the carbon dioxide from the gas stream.
  • the water gas shift reaction is known, and suitable equipment for carrying out the reaction is commercially available.
  • An alternative method for removing carbon oxides is pressure swing absorption. This procedure can be employed not only to remove the desired amount of carbon monoxide and/or carbon dioxide, but also to remove
  • Pressure swing adsorption involves the adsorption of components to be removed at high pressure followed by their desorption at low pressure.
  • the process operates on a repeated cycle having two basic steps, adsorption and regeneration. Not all the hydrogen is recovered as some is lost in the waste gas during the regeneration stage, but by careful selection of the frequency and sequence of steps within the cycle the recovery of hydrogen is maximized.
  • Regeneration of the adsorbent is carried out in three basic steps.
  • the adsorber is depressurized to the low pressure. Some of the waste components are desorbed during this step. (b) The adsorbent is purged at low pressure, with the product hydrogen removing the remaining waste components.
  • Pressure swing adsorption is most effectively employed to remove carbon oxides and inert matnriaJs, for example, methane and nitrogen, from the methanol synthesis loop.
  • the hydrogen, carbon dioxide and carbon monoxide are reacted under methanol-producing conditions. Any known procedure for reacting hydrogen, carbon dioxide and carbon monoxide to produce methanol can be used. Preferably, they are reacted in a medium pressure process in a circulatin catalytic reactor at a pressure from about 50 atm (5070 to 120 atin (12160 KPa), more preferably from about 70 at (7090 KPa) to 100 atm (10130 KPa). This type of synthes and equipment for carrying it out are known in the art.
  • the synthesis gas is compressed to between 50
  • the gas enters methanol synthesis vessel in which it is admixed with recycle gas.
  • the gas then passes into a catalytic methanol converter.
  • the methanol converter typically comprises a pressure vessel containing a catalyst bed and facilities for moderating the exothermic reaction of hydrogen with the carbon oxides to produce methanol, for example by in jecting cold gas at intervals within the catalyst bed.
  • Any commercially available catalyst which is capable of catalyzing the reaction of hydrogen, carbon monoxide and carbon dioxide to produce methanol can be employed.
  • Such catalysts are manufactured, for example, ly Imperial Chemical Industries, Katalco, and Haldor Topsoe, Inc.
  • Preferred catalysts for methanol synthesis are
  • the methanol synthesis reaction generally takes place at a temperature in the range from about 410°F (210°C) to 570°F (300°C),
  • the exit gas from the methanol converter is passed through a condenser in which it is cooled with water and then through a separator.
  • the bottoms product of the separator contains methanol condensate.
  • An amount of gas is purged from the methanol synthesis loop in order to maintain the concentration of inert gases circulating within the loop at acceptable levels.
  • the remaining gas is admixed with incoming synthesis gas to be recycled to the methanol converter.
  • carbon monoxide and/or carbon dioxide can be removed either from the circulating gas in the methanol synthesis loop or from the purge gas in order to maintain the desired MSSGR, preferably between 0.96 and 1.10, entering the converter.
  • the desired MSSGR preferably between 0.96 and 1.10
  • the crude methanol containing water is then purified by conventional means to obtain essentially pure methanol.
  • the methanol condensate is purified in one or more distillation columns.
  • the catalytic partial oxidation process offers the following advantages.
  • Hydrocarbonaceous feedstock such as natural gas stream 200, FIG. 14, is optionally desulfurized using conventional methods.
  • Desulfurization may, for example, be conveniently carried out by preheating the hydrocarbonaceous feedstock to a temperature between about 250°F (120°C) and 750°F (400oC) and absorbing the sulfur compounds into zinc oxide contained in one or several desulfurization vessels 202.
  • desulfurization vessel 202 is located upstream of a
  • the desulfurization vessel 202 can be located downstream of the synthesis gas compressor, as shown in the embodiment in FIG. 16.
  • the desulfurized hydrocarbonaceous feedstock is saturated with water vapor in the forced film saturator 204.
  • the forced film saturator 204 is a distinctive feature of the process and results in reduced capital cost and power requirements when compared with the more conventional pacled tower type of saturator presently employed in commercial installations.
  • the forced film saturator 204 consists of a vertical shell and tube exchanger and a water circulation system 206.
  • the desulfurized hydrocarbonaceous feedstock and water enter the top head 208 of the exchanger and flow vertically downward through the tubes 210.
  • the feedstock is heated as it flows though the tubes, for example to a temperature of about 400°F (206°C), and as it is heated, its water vapor content increases due to vaporization of the circulating water.
  • the heating medium in this case
  • methanol reactor effluent gas stream 330 passes to the shell side of the tubes 210 in the forced film saturator at the bottom and flows countercurrent to the feedstock flow to emerge at a lower temperature at the top outlet of the shell.
  • the desulfurized hydrocarbonaceous feedstock and unvaporized water emerge from the exchanger tubes 210 at the bottom and are separated in the bottom head 212.
  • Recovered water is recirculated by the pump 206 to the top 208 of the forced film saturator and the saturated hydrocarbonaceous feedstock passes to saturated feedstock line 220.
  • Makeup water 216 is added to the recycle water stream, and blowdown line 218 for the saturator water recycle stream is ptovided
  • the major advantages of the forced film saturator 204 over the conventional packed bed saturator are: (1) The forced film saturator has a considerably lower water circulation rate and hence lower power consumption. (2) The forced film saturator has a simpler design and
  • the saturated feedstock passes through line 220 to coil 222 of a fired heater 224 where the saturated feedstock is preheated to a temperature in the range from 650°F (340°C) to 1200°F (650°C), for example to 1100°F (595°C), by combustion of fuel in the fired heater with air.
  • the fuel 225 can be one or more waste gas streams from the methanol plant, such as fusel oil stream 226 from distillation, flash gas stream 228 from the ammonia synthesis loop, PS ⁇ purge gas stream 230, and light ends stream 232 from distillation which are combusted with air 227.
  • the heated saturated feedstock on line 234 from the heater 224 is fed to the catalytic partial oxidation
  • CPO CPO
  • the oxygen stream 236 is obtained from an air separation plant 230 and preheated to a temperature in the range from 150°F (65°C) to 1200°F (650°C), for example, 300°F before being mixed with the natural gas and steam feedstock passing to the CPO catalyst.
  • High pressure steam is used to preheat the oxygen in a heat exchanger 240.
  • The. main overall reactions taking place within the CPO reactor 28 are the partial oxidation reactions:
  • FIG . 6 shows oxygen consumption for the above process, as a function of the steam-to-carbon molar ratio of natural gas feedstock, for reaction temperatures of 1,600oF (870°C), 1,750°F (950°C) and 1,900°F (1040°C) and an
  • the dashed line 25 in FIG. 6 represents the minimum temperature and steam conditions, i.e. the minimum temperature is ⁇ linear function of the steam-to-carbon ratio, which have been discovered to prevent formation of carbon deposits on the catalyst.
  • lower temperature of reaction is preferred, such as an exit temperature of 1700°F (925oC) at a pressure of about 415 psig (2860 KPa).
  • the saturator 204 provides the total quantity of steam necessary to achieve a steam-to- carbon molar ratio of approximately 1.3 to 1.0 and
  • FIG. 7 shows the molar ratio of hydrogen, as H 2 , to carbon monoxide in the product as a function of
  • FIGS. 8 and 9, respectively, show the amounts of methane and carbon dioxide, as volume %, in the product as a function of steam-to-carbon ratio for reaction temperatures of 1,600°F (870°C), 1,750°F (950°C) and 1,900°F (1040°C).
  • FIG. 10 shows the effective H 2 production of the process, expressed as total moles of H 2 and carbon monoxid in the product divided by total moles of H 2 and carbon in the feedstock as a function of steam-to-carbon ratio for reaction temperatures of 1,600°F (870°C), 1,750°F (950o C) and 1,900°F (1040°C).
  • the reactor effluent 244 is first cooled by generating steam in a boiler 246 which receives water from a high pressure steam drum 248 and returns steam to the drum
  • the drum 248 operates at a high pressure of, for example, about 1550 psig (10700 KPa).
  • Water supply 250 for the steam drum is first heated in coils 252 of the heater 224 and then fed to the drum through line 254.
  • Steam output. 256 from the drum 248 is further heated in coils 258 to produce
  • superheated steam 260 which can be used to operate si eam turbines or provide heating for process steps.
  • Line 262 to a blow down drum (not shown) provides for blow down of the drum 248.
  • the reactor effluent is passed by line 264 to distillation column reboiler 266, by line 268 to
  • distillation column reboiler 270 by line 272 to
  • the synthesis gas in line 282 is mixed with hydrogen 204 from PSA unit 286, FIG. 14, and passes to m ⁇ ke-up compressor unit 288 where the synthesis gas is compressed to the methanol loop pressure, for
  • the make-up compressor unit of FIG. 13 includes compressor 290, heat exchanger 292, knockout drum 294, and compressor 296 to produce the make-up gas stream 290 which is passed to the methanol synthesis loop in FIG. 14.
  • the compressors 290 and 296 can
  • the make-up synthesis gas 298 from the compresso unit 280 joins the gas circulating in the methanol synthesis loop at the discharge 304 of the circulator 302.
  • the loop stream 306 is preheated in loop heat interchanyer 300 before passing in line 310 to tube cooled methanol converter 3.1
  • Start-up heater 314 is provided so that, upon start up of the methanol converter, the incoming process stream can be heated to the reaction temperature until the methanol converter becomes hot enough due to the heat generated by the reaction to heat the incoming gases, whereupon oporatien of the start-up heater 314 can be discontinued.
  • the tube cooled converter shown in detail in FIG . 12, is a distinctive feature of the process and includes a simple gas-to-gas heat exchanger with no high differential pressure tubesheets.
  • the vessel shell 314 is designed to retain the process pressure.
  • Inlet 316 connects through distributor 310 to intermediate branch tubes 320 which, in turn, are connected through distributors 322 to a plurality of tubes 324 extending vertically upward through the
  • the catalyst bed 326 is supported on a bed of ceramic balls 320 which are separated from the outlet 330 by screen support 332.
  • the reactant gas is fed though inlet 316 into the bottom of the converter where it is distributed by distributors 310, intermediate branch tubes 320, and distributors 322 to the vertical tubes 324.
  • the gas is heated as it flows upward by heat exchange through the tube and distributor walls. From the exits 334 of tubes 324 at the top of the vessel, the reactant gas pass downward, through the catalyst, which is packed in the space between the tubes 324.
  • the gas flow through the tubes 324 is countercurrent to the gas flow through the catalyst bed.
  • the temperature profile against methanol concentration is quasi-isothermal and is very favorable in terms of reaction rates and conversion.
  • the main features of the tube cooled converter are: (1) The design is mechanically simple with no high differential pressure across the tube material and no tubesheet construction problems. (2) Catalyst loading is simple. (3) A single reactor can be constructed capable of producing methanol at high capacity of over 2000 tons (1,815,000 kg) per day. (4) Methanol synthesis loop
  • the gas is cooled, for example to a temperature of about 340°F (170°C), by heat transfer to the circulating water and incoming natural gas stream flowing through the tubes 210.
  • the converter effluent gas emerges from the saturator in line 336 and passes to the heat exchanger 300 where the effluent gas is further cooled by heat exchange with the loop feed for the converter. Then the gas from the converter is fed over line 338 to loop condenser 340 which is cooled by water to condense the methanol in the methanol loop stream. Condensed methanol is separated from the methanol loop in separator 342, and passed over line 344 to pressure letdown vessel 346 before being passed in crude methanol stream 340 to the
  • the gas stream 354 from the separator is split into a recycle stream 356 passing to the suction inlet of the circulator 302 and a purge stream 350 taken through valve 360 to maintain the concentration of inert material in the methanol loop at acceptable levels.
  • Hydrogen is
  • the crude methanol 340 is passed to a distillation column 364 where light end materials such as absorbed gases are
  • the noncondensed light ends 232 are passed as fuel to the fired heater.
  • Bottom streams frcm the column 364 are recycled through reboilers 266 and 374 heated by the synthesis gas stream and steam, respectively.
  • the methanol stream 378 is passed by pump 380 to distillation column 382 where the stream is separated into the fusel oil stream 226, a water stream 384, and the product methanol stream 386.
  • the distillation column has retailers 270 and 388 heated by the synthesis gas stream and steam, respectively.
  • the overhead 389 is withdrawn under vacuum produced by vacuum pump system 390.
  • the vacuum pump system 390 includes pump 398 circulating water from collector 400 through water cooled heat exchanger 402. Excess water is passed by line 404 from the collector 400 back to the bottom of the distillation column 382.
  • the water stream 304, withdrawn by pump 406 from the bottoms of column 302, is further cooled by heat exchanger before discharge in line 410.
  • FIGS. 16 and 17 Hydrocarbonaceous feedstock, such as natural gas stream 420, FIG. 17, is passed to a feedstock saturator 422 which is similar to the saturator 204 in the embodiment shown in FIG. 14. Water is recirculated by pump 424 to the top of the forced film saturator where it flows with the feedstock stream through the tubes of the saturator. The saturated feedstock is separated from the water in the bottom of the saturator to form saturated feedstock stream 425 and the water recycle stream. Makeup water for the recycle water stream comes from synthesis gas condensate 426 and
  • Methanol converter effluent stream is fed to the shell side of the saturator to heat the water and feedstock stream.
  • Line 430 provides Cor blowdown of the saturator water.
  • the saturated feedstock 425 passes through coil 434 of a fired heater 436 where the saturated feedstock is preheated by combustion of waste fuel 438 with air 439.
  • the heated saturated feedstock on line 440 from the heater 436 is fed to the catalytic partial oxidation (CPO) reactor 28 where it is mixed with an oxygen or oxygen-containing stream 442 and subjected to catalytic partial oxidation to produce synthesis gas as has been described hereinbefore.
  • CPO catalytic partial oxidation
  • the reactor effluent 444 is first cooled by generating steam in a boiler 446 which receives water from a steam drum 448 and returns steam to the drum.
  • a second steam drum 450 has water flow boiled in coils 452 of the fired heater 436. Steam outputs 454 and 456 of the steam drums 448 and 450 are combined to produce the steam 458 used in the process.
  • Lines 460 and 462 provide for boiler blowdown from the drums 448 and 450.
  • Boiler feedwater 464 branches into streams 466 and 468 feeding the respective drums 448 and 450.
  • the branch 466 is preheated in heat exchanger 472 by the reactor effluent in line 470 from the boiler 446 before passing in line 474 to the drum 448.
  • the synthesis gas in line 476 from the heat exchanger 477 is further cooled by water in condenser 478 and then passed through line 480 to synthesis gas separator 402 where water condensate 484 is removed by pump 486 to provide the
  • the synthesis cjas in line 490 is compressed to the methanol loop pressure by compressor 492 to produce feed 494 to the desulfurization vessel 202.
  • the output 496 of the desulfurization vessel 202 forms the make-up gas for the methanol loop of FIG. 17.
  • the make-up synthesis gas 496 joins the gas circulating in the methanol synthesis loop at the discharge 498 of the methanol loop circulator 500.
  • the loop stream 502 is preheated in heat exchanger 504 by steam before passing in line 506 to tube cooled methanol converter 508 which substantially similar to tube cooled methanol converter 312 of FIG. 14.
  • Line 510 from the loop heater 504 provides heated loop gas to the top of the converter 508 during initial start-up of the methanol converter.
  • methanol loop stream line 518 Condensed methanol is separated from the methanol loop in separator 520 and passed over line 522 and valve 524 to pressure letdown vessel 57.
  • the gas stream 528 from the separator 520 is splid into a recycle stream 530 passing to the suction inlet of the circulator 500 and a purge stream 532 taken through valve 534 to maintain the concentration of inert material in the methanol loop at acceptable levels. Flash gas 536 through valve 530 from the let-down vessel 526 is combined with the purge stream 532 to form the waste fuel stream 488 to the fired heater 436.
  • crude methanol 540 is passed to a distillation column 542 where the stream is separated into a light ends stream 544, a water stream 546, and the product methanol stream 540.
  • the distillation column has a reboiler 550 heated by steam.
  • the overhead 552 is cooled by water cooled condenser 554 producing a
  • condensed methanol stream 556 a portion of which is used as a reflux stream through pump 558.
  • the remaining portion of the condensed methanol stream 556 is combined with the fused oil fraction 560 to form the product stream 548.
  • the water stream 546 from the bottoms of column 542, is cooled in cooler 562 and passed by pump 564 to the stream 428 feeding the saturator 422.
  • a methanol plant in accordance with FIGS. 13, 1 and 15 is operated to produce 2000 tons (1,800,000 kg) of methanol per day.
  • TABLES I, II, and III se forth moles/hour, mole percent, and parameters of pressure, temperature, water/steam, and heat transfer for the plant.
  • the moles/hour are lb moles/hour (0.4536 kg moles/hour).
  • a methanol plant in accordance with FIGS. 16 and 17 is operated to produce 102,000 1bs. (46,400 kg) of methanol per day.
  • the following TABLES IV, V, and VI set forth moles/hour, mole percent, and parameters of pressure, temperature, water/steam, and heat transfer for the plant.
  • the moles/hour are 1b moles/hour (0.4536 kg moles/hour).

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Abstract

Procédé de production permettant de produire du méthanol à partir d'hydrocarbures, qui consiste à oxyder partiellement par catalyse les hydrocarbures à une température et une vapeur propres à produire un gaz de synthèse contenant de l'hydrogène, du monoxyde de carbone et du dioxyde de carbone mais non pas du carbone libre, à faire réagir ensuite l'hydrogène, le monoxyde de carbone et le dioxyde de carbone dans des conditions de production de méthanol et à récupérer enfin le méthanol.
PCT/US1989/005370 1988-11-30 1989-11-30 Production de methanol a partir d'une charge d'hydrocarbures WO1990006297A1 (fr)

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Cited By (12)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US5648582A (en) * 1993-08-20 1997-07-15 Regents Of The University Of Minnesota Stable, ultra-low residence time partial oxidation
US5654491A (en) * 1996-02-09 1997-08-05 Regents Of The University Of Minnesota Process for the partial oxidation of alkanes
US5886056A (en) * 1997-04-25 1999-03-23 Exxon Research And Engineering Company Rapid injection process and apparatus for producing synthesis gas (law 560)
US5905180A (en) * 1996-01-22 1999-05-18 Regents Of The University Of Minnesota Catalytic oxidative dehydrogenation process and catalyst
US5935489A (en) * 1997-04-25 1999-08-10 Exxon Research And Engineering Co. Distributed injection process and apparatus for producing synthesis gas
US5980782A (en) * 1997-04-25 1999-11-09 Exxon Research And Engineering Co. Face-mixing fluid bed process and apparatus for producing synthesis gas
US5980596A (en) * 1997-04-25 1999-11-09 Exxon Research And Engineering Co. Multi-injector autothermal reforming process and apparatus for producing synthesis gas (law 565).
US6254807B1 (en) 1998-01-12 2001-07-03 Regents Of The University Of Minnesota Control of H2 and CO produced in partial oxidation process
US6267912B1 (en) 1997-04-25 2001-07-31 Exxon Research And Engineering Co. Distributed injection catalytic partial oxidation process and apparatus for producing synthesis gas
US6333294B1 (en) 1998-05-22 2001-12-25 Conoco Inc. Fischer-tropsch processes and catalysts with promoters
WO2017121978A1 (fr) * 2016-01-11 2017-07-20 Johnson Matthey Public Limited Company Procédé de synthèse du méthanol
WO2020163169A1 (fr) * 2019-02-05 2020-08-13 Saudi Arabian Oil Company Production de gaz synthétique

Families Citing this family (2)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
CA2357527C (fr) 2001-10-01 2009-12-01 Technology Convergence Inc. Procede de production de methanol
DE102007040707B4 (de) * 2007-08-29 2012-05-16 Lurgi Gmbh Verfahren und Anlage zur Herstellung von Methanol

Citations (2)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
EP0067491A2 (fr) * 1981-06-12 1982-12-22 Stamicarbon B.V. Procédé pour la préparation de méthanol
EP0111376A2 (fr) * 1982-12-14 1984-06-20 Stamicarbon B.V. Procédé de préparation de méthanol

Patent Citations (2)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
EP0067491A2 (fr) * 1981-06-12 1982-12-22 Stamicarbon B.V. Procédé pour la préparation de méthanol
EP0111376A2 (fr) * 1982-12-14 1984-06-20 Stamicarbon B.V. Procédé de préparation de méthanol

Cited By (16)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US5648582A (en) * 1993-08-20 1997-07-15 Regents Of The University Of Minnesota Stable, ultra-low residence time partial oxidation
US6072097A (en) * 1996-01-22 2000-06-06 Regents Of The University Of Minnesota Catalytic oxidative dehydrogenation process and catalyst
US6846773B1 (en) 1996-01-22 2005-01-25 Regents Of The University Of Minnesota Catalytic oxidative dehydrogenation process and catalyst
US5905180A (en) * 1996-01-22 1999-05-18 Regents Of The University Of Minnesota Catalytic oxidative dehydrogenation process and catalyst
US6548447B1 (en) 1996-01-22 2003-04-15 Regents Of The University Of Minnesota Catalytic oxidative dehydrogenation process and catalyst
US5654491A (en) * 1996-02-09 1997-08-05 Regents Of The University Of Minnesota Process for the partial oxidation of alkanes
US5935489A (en) * 1997-04-25 1999-08-10 Exxon Research And Engineering Co. Distributed injection process and apparatus for producing synthesis gas
US5980596A (en) * 1997-04-25 1999-11-09 Exxon Research And Engineering Co. Multi-injector autothermal reforming process and apparatus for producing synthesis gas (law 565).
US6267912B1 (en) 1997-04-25 2001-07-31 Exxon Research And Engineering Co. Distributed injection catalytic partial oxidation process and apparatus for producing synthesis gas
US5980782A (en) * 1997-04-25 1999-11-09 Exxon Research And Engineering Co. Face-mixing fluid bed process and apparatus for producing synthesis gas
US5886056A (en) * 1997-04-25 1999-03-23 Exxon Research And Engineering Company Rapid injection process and apparatus for producing synthesis gas (law 560)
US6254807B1 (en) 1998-01-12 2001-07-03 Regents Of The University Of Minnesota Control of H2 and CO produced in partial oxidation process
US6333294B1 (en) 1998-05-22 2001-12-25 Conoco Inc. Fischer-tropsch processes and catalysts with promoters
WO2017121978A1 (fr) * 2016-01-11 2017-07-20 Johnson Matthey Public Limited Company Procédé de synthèse du méthanol
WO2020163169A1 (fr) * 2019-02-05 2020-08-13 Saudi Arabian Oil Company Production de gaz synthétique
US11807822B2 (en) 2019-02-05 2023-11-07 Saudi Arabian Oil Company Producing synthetic gas

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AR245685A1 (es) 1994-02-28
AU4665189A (en) 1990-06-26
CN1043493A (zh) 1990-07-04

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