EP0209665A1 - Procédé d'hydrogénation de charbon en phase liquide et en lit fixe avec catalyseur - Google Patents

Procédé d'hydrogénation de charbon en phase liquide et en lit fixe avec catalyseur Download PDF

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Publication number
EP0209665A1
EP0209665A1 EP86106751A EP86106751A EP0209665A1 EP 0209665 A1 EP0209665 A1 EP 0209665A1 EP 86106751 A EP86106751 A EP 86106751A EP 86106751 A EP86106751 A EP 86106751A EP 0209665 A1 EP0209665 A1 EP 0209665A1
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Prior art keywords
oil
separator
hydrogenation
coal
reactor
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Granted
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EP86106751A
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German (de)
English (en)
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EP0209665B1 (fr
Inventor
Bernd Strobel
Frank Dr. Dipl.-Chem. Friedrich
Eckhard Wolowski
Rainer Löring
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RAG AG
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Ruhrkohle AG
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    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G1/00—Production of liquid hydrocarbon mixtures from oil-shale, oil-sand, or non-melting solid carbonaceous or similar materials, e.g. wood, coal
    • C10G1/002—Production of liquid hydrocarbon mixtures from oil-shale, oil-sand, or non-melting solid carbonaceous or similar materials, e.g. wood, coal in combination with oil conversion- or refining processes

Definitions

  • the invention relates to a process for coal hydrogenation by means of bottom phase and catalyst fixed bed hydrogenation according to the preamble of claim 1.
  • the coal When extracting oils from coal by pressure hydrogenation, the coal is mixed with a distillate-like solvent (grinding oil, mashing oil, return oil) that is continuously recovered in the process and mixed into a pumpable mixture and, after adding hydrogen-containing gas, at pressures above 100 bar and temperatures between 450 and 500 ° C in one or more reactors in the liquid phase (bottom phase) converted.
  • a finely divided catalyst is used and continuously passed through the reaction space as part of the coal suspension, or a flowing catalyst bed in which the catalyst material is in the form of a piece in the reaction space.
  • the reaction products are first separated into two partial streams, the bottom and the top product, in a hot separator a little below the reaction temperature.
  • the sump contains the solid and non-liquefiable constituents of the coal and optionally the catalyst, as well as the asphalt-containing, non-distillable Products of coal hydrogenation as well as small to medium quantities of distillable oil.
  • This oil is recovered from the sludge by distillation or by other measures such as smoldering or solvent extraction and used to produce the return oil.
  • the low oil residue is commonly used as a feedstock to generate hydrogen in suitable gasification processes.
  • the top product of the hot separator can be passed over a second hot separator in order to separate any entrained solids and asphalt parts.
  • the sump of the second hot separator is then processed together with the sludge from the first hot separator or used directly to produce the return oil.
  • the top product of the hot separator (s) which is free of solids and asphaltene, consists of vaporous oils: heavy oil, boiling above 325 ° C at normal pressure; Medium oil, boiling at 200 - 325 ° C, light oil, boiling up to 200 ° C.
  • the top product also contains hydrogen and hydrocarbon gases, as well as small amounts of other products such as water vapor and inorganic gases.
  • the subset of the vapors and gases from the hot separator is chosen to be substantially larger than corresponds to the newly formed oil; A large part of the oil returned to the production of coal pulp is thus also hydrogenated refined on the fixed bed catalyst. Unless additional reaction space with catalyst is provided for this, the degree of refining of the product oil is reduced even further. By replacing light and medium oil with heavy oil, the boiling point of the product oil can be reduced with this mode of operation; however, this requires a significant increase in equipment and energy for the distillation steps and still does not improve the degree of refining of the oil fractions.
  • Both embodiments of the process according to DE-OS 26 54 635 also have the disadvantage that oils from all boiling points are only treated together on the same catalyst and therefore with moderate refining performance.
  • a particularly intensive refining and hydrogenation, especially of the lighter fractions, is usually desirable, e.g. B. to be able to use them in reforming processes without further hydrogenation stage; at the same time, it is usually sufficient to hydrogenate the return oil for the pulp production and possibly also the medium oil content of the product only moderately intensively, so that hydrogen can be saved.
  • the object of the invention is to avoid the disadvantages inherent in the generic method and to obtain a hydrocarbon oil which is essentially free of oxygen, nitrogen and sulfur compounds in high yield.
  • a return oil is obtained which is lighter in terms of density and boiling point than unrefined return oil. Only a small proportion of the total evaporable oil remains in the sludge and is extracted from it.
  • the temperature at the head of the hot separator is at least 440 ° C; less than 25% of the return oil is obtained from the sludge, which is less than 20% of the total vaporizable oil (product oil plus return oil).
  • the oil obtained from the blowdown is pumped back directly into the high pressure position, whereby both the injection into the hot top products of the hot separator can be useful to control the inlet temperature of the first fixed bed reactor and the injection upstream of the (first) hot separator, the sensible heat of the products from the bottom phase reactor is sufficient to evaporate the blowdown oil.
  • the oil from the sludge can be hydrogenated in a separate reactor outside the gas cycle if a particularly gentle procedure is desired for the first fixed bed reactor system; the pre-refined oil can then either be injected into the top product of the optionally second hot separator and thus further refined on the first fixed bed reactor system, or it is used directly as a return oil fraction.
  • the pre-refined oil is advantageously passed immediately after leaving the separate reactor and together with the accompanying gases and vapors before the preheater of the sump phase.
  • the tangible Heat and the excess hydrogen are used, and the coal pulp can be made available in a correspondingly higher concentration.
  • a selective refining for producing a light oil which is suitable directly as a reformer and with a relatively low hydrogen consumption will be achieved according to a further development of the method if, after the return oil has been extracted from the intermediate separator, a second intermediate separation is carried out at a further reduced temperature.
  • a high-pressure fractionation column can be attached to the second intermediate separator to improve the selectivity.
  • the separating cut is placed at about 185 ° C, so that a moderately refined, but storage-stable medium oil is produced, which, for. B. can be used as heating oil.
  • the mixture of vapors and gases leaving the fractionation column overhead is finely cleaned in the second fixed bed reactor system after appropriate preheating. After cooling, a light oil that is practically completely free of heteroatoms (boiling up to 185 ° C.) is obtained from the cold separator.
  • the method according to the invention has the following advantages over the known methods: Despite the removal of oxygen and nitrogen from the product oil, due to the simultaneously hydrogenating refining of the return oil, there is no reduction in yield, but on the contrary, surprisingly, there is a substantial increase in yield.
  • the reactor for the return oil hydrogenation and the reactor for the fine refining of the product oil in a common high-pressure gas circuit saves a considerable number of apparatus and machines and a lot of energy compared to the operation of separate high-pressure systems. Since the oils are relaxed less frequently, the loss of solubility in hydrogen pressure is lower. The higher molecular weight compounds formed during the intermediate condensation of low-refined oils have a severe damaging effect on fixed bed catalysts, which is why such oils generally require an additional intermediate treatment, e.g. B. by distillation. In contrast, intermediate condensation is avoided in the process according to the invention; the oils get mainly in vapor form and under hydrogen pressure directly onto the refining catalysts, which therefore have a long service life. As catalysts, the compounds of Fe, Co, Ni, W, Mo, Zn or Sn with oxygen or sulfur that are customary in the hydrogenation of carbon are mostly used on a support and also in combinations of several of these compounds.
  • FIG. 1 shows the operation of the method according to the invention with an intermediate separator arranged between two fixed bed reactors.
  • ground coal and optionally catalyst mass are mixed with the oil from an intermediate separator 14 to form a slurry.
  • the ratio of coal (anhydrous) to oil can be about 1: 0.8 to 1: 3, ratios between 1: 1 and 1: 1.5 are advantageous.
  • the coal pulp is conveyed with a pump 2 against the operating pressure of the hydrogenation system, which is more than 100 bar, preferably 150 to 400 bar.
  • hydrogenation gas is supplied, which consists of recycle gas (line 21) and hydrogen (line 22).
  • the hydrogen content in the cycle gas (hydrogenation gas) should be more than 50% by volume.
  • - Recycle gas is also blown into the hydrogenation reactors 5, 12 and 16 in the required amount for temperature control at different heights.
  • the total amount of circulating gas, measured on the compressor 20, is between 1 and 8 normal cubic meters per kg of coal (water and ash free); 3 to 5 cubic meters per kg of pure coal are preferred.
  • the amount of fresh hydrogen is, depending on the hydrogen consumption, 700 to 1500 normal cubic meters per kg of coal used.
  • Coal slurry and hydrogenation gas are heated in a preheater 4 and reacted in a bottom phase reactor 5 at temperatures between 450 and 500 ° C.
  • the reactor 5 can consist of a single or of several vessels. If it is equipped with a flowing catalyst bed, the coal slurry need not contain any catalyst mass.
  • a hot separator 6 vapors and gases are separated from the liquid and solid substances (sludge) at 440 to 480 ° C and passed on overhead of the separator 6. The blowdown is decompressed and flashed to extract the oils it contains. The flash residue is used in the usual way to generate hydrogen.
  • the vapors from the flash process are condensed in a heat exchanger 8 and passed into a receiver 9 as flash oil.
  • the flash oil is either used directly via a line 10 A for slurry production, or a high pressure pump 10 delivers the flash oil into the vapors / gases from the hot separator 6.
  • the temperature of the mixture is regulated by a heat exchanger 11 so that the inlet temperature in the reactor 12 den desired value (between 350 and 420 ° C).
  • Hydrogenation and refining catalysts of the type which are customary in the processing of coal oils and petroleum are used as catalysts in the reactors 12 and 16; it can be the same or below in the reactors 12 or 16 Different catalysts are used in order to improve the degree of refinement, saturation, cleavage and hydrogen consumption. B. to achieve the best results for the respective coal or the respective end product.
  • the vapors / gases from the reactor 12 are cooled in a heat exchanger 13 to such an extent that such an amount of oil is continuously condensed as is consumed in the production of the coal pulp.
  • This return oil is expanded from an intermediate separator 14 and returned to the mixing plant 1.
  • the temperatures required before the intermediate separator 14 are between 250 and 350 ° C.
  • the vapors and gases that leave the intermediate separator 14 are raised to the inlet temperature of the fixed bed reactor 16 (350 to 420 ° C.) by heat exchange and possibly additional temperature control (cooler / heater 15). By cooling to temperatures below 50 ° C in a heat exchanger 17, the product oil is separated from the mixture of vapors and gases; In addition, hydrogenation water, which contains ammonia and hydrogen sulfide, condenses at this point. These liquids are expanded from a cold separator 18 and fed to further processing or use.
  • a gas mixture is drawn off, which essentially consists of hydrogen and hydrocarbon gases, but also contains hydrogen sulfide, ammonia and small amounts of carbon oxides.
  • this gas is cleaned to the required extent and enriched with hydrogen.
  • a cycle gas compressor 20 conveys the cycle gas back to the hydrogenation reactors.
  • Fig. 2 shows an embodiment of the method according to the invention with two intermediate separators 14 and 15 A between the two fixed bed reactors 12 and 16.
  • This method of operation is advantageous if only the light oil portion of the product oil has to be refined to a very large extent, but the middle oil portion can still be used as a storage-stable, moderately refined product.
  • the return oil is obtained from the vapors / gases mixture after the reactor 12.
  • the top product from the intermediate separator 14 is cooled in the heat exchanger 15 to such an extent that essentially middle oil (boiling at 185 to 325 ° C.) can be removed from the second separator 15 A.
  • This intermediate separator 15 A can be equipped in the manner of a distillation column with packing or other internals to improve the selectivity.
  • the vapors and gases withdrawing at the top of the separator or the column are brought to the inlet temperature of the fixed bed reactor (350 to 420 ° C.) in a heat exchanger 15 B.
  • an oil is obtained from the cold separator 18 which mainly consists of light oil (boiling end 185 ° C.) and which has reformer use quality.
  • the distillate from the flash system 7 is injected with the aid of the pump 10 via a line 26 into the hot products of the bottom phase reactor 5, before they enter the hot separator 6.
  • the heat required to evaporate the flash oil is removed from the products of the bottom phase reactor 5.
  • FIG. 4 shows an embodiment of the method according to the invention in which an additional reactor 25 is arranged outside the common gas circuit.
  • the flash oil is heated via a line 23 in a heat exchanger / preheater 24 after the addition of hydrogen or hydrogen-containing gas and in the fixed bed reactor 25 at 350 to 420 ° C. under approximately the same pressure as in the other reactors hydrated.
  • the addition of hydrogen to flash oil is in the range of 0.5 to 5 m3 / kg.
  • the entire outlet products of the reactor 25 are fed through line 26 to the pulp which is already under pressure upstream of the preheater 4.
  • the solids content of the coal pulp produced in the mixing plant 1 is accordingly set higher than in the other working methods.
  • the hydrogen not consumed in the reactor 25 is completely available for the bottom phase hydrogenation; the fresh hydrogen supply via line 22 can therefore be reduced accordingly.
  • Example 1 In a system according to Example 1, 126 kg of water-free gas coal (120 kg / h free of water and ash) are mixed with 5 kg of dried red mass and 134 kg / h of return oil to form a slurry and together with 650 m3 / h of hydrogenation gas (gas quantities under normal conditions ) consisting of 150 m3 / h fresh hydrogen and 500 m3 / h recycle gas (contains 60 vol .-% hydrogen) passed through a bottom phase reactor 5.
  • the pressure in the reactor 5 is 400 bar, the temperature 470 ° C.
  • the temperature in The vapor space of the hot separator 6 is kept at 440 ° C.
  • the sludge from the hot separator 6 is subjected to flash evaporation (flashing); this results in 24 kg / h of flash oil, which is used without further treatment to produce the return oil.
  • the entire top products of the hot separator 6 are passed through the fixed bed reactor 12, which contains 80 kg of a commercially available catalyst made of sulfidic nickel and molybdenum on Al2O3-SiO2 carrier.
  • the average catalyst temperature is 380 ° C, the pressure 400 bar.
  • the emerging products are cooled to 275 ° C. 110 kg / h of oil are obtained in liquid form, which are removed from the intermediate separator 14 and combined with the flash oil from the line 10 A.
  • the return oil produced in this way contains 38% heavy oil (boiling above 325 ° C) and 62% medium oil.
  • the top products of the intermediate separator 14 are passed through the fixed bed reactor 16, which is filled with 80 kg of a commercially available catalyst consisting of molybdenum sulfide and nickel sulfide on alumina carrier.
  • the mean catalyst temperature is 390 ° C, the pressure 400 bar.
  • 65 kg / h (54% of the waf coal) of water-clear product oil are condensed from the reaction products and are discharged from the cold separator 18.
  • the product oil contains 20 mg / kg basic nitrogen and 50 mg / kg phenolic oxygen. After storage for 1 month under the exclusion of air and light, the oil is slightly yellowish.
  • the oil yield is higher and at the same time the oil quality is considerably better than with coal hydrogenation without integrated refining stages.
  • the blowdown from the hot separator 6 is subjected to flash evaporation; 21 kg / h of distillate are obtained, which are pumped in front of the inlet of the fixed bed reactor 12.
  • the entire top product of the hot separator 6 is also passed through this reactor 12.
  • the reactor 12 contains 80 kg of a commercially available refining catalyst based on nickel, molybdenum and alumina; the average catalyst temperature is 390 ° C.
  • the return oil consists of 30% heavy oil boiling above 325 ° C and 70% medium oil boiling up to 325 ° C.
  • the vapors and gases that withdraw at the top of the intermediate separator 14 are passed through the fixed bed reactor 16, which also contains 80 kg of a commercially available refining catalyst based on Ni, Mo and Al2O3.
  • the average catalyst temperature is 390 ° C.
  • 18 55 kg of water-light product oil are obtained in the cold separator, which consists of 40% light oil ( ⁇ 185 ° C) and 60% medium oil (185 - 325 ° C). After a month, the oil is still colorless.
  • the oil contains only 6 mg / kg basic nitrogen and less than 15 mg / kg phenolic oxygen.
  • the light oil alone contains less than 2 mg / kg nitrogen. So the oil gain is 55% and the oil quality is good.
  • Example 2 An experiment is carried out under the same conditions as in Example 2, but the vapors and gases carried at 290 ° C overhead of the intermediate separator 14 are cooled to 170 ° C and in the stripping section of a high pressure-resistant packed column 15 A with about 25 theoretical plates ( at 20 l / h liquid load) initiated. 33 kg / h of medium oil with a boiling range of 175-325 ° C. are released from the bottom of the column. The vapors and gases withdrawing at the top of the column at 160 ° C. are heated and passed over the fixed bed reactor 16.
  • the reactor 16 contains 50 kg of a commercially available Ni-Mo-Al2O3 refining catalyst. The average temperature of the catalyst bed is kept at 375 ° C. The reactor outlet products are cooled to 20 ° C. From the cold separator, 22 kg of light oil with a boiling end of 185 ° C. are obtained per hour
  • the medium oil contains 0.06% basic nitrogen and ⁇ 0.1% oxygen. After storage for 1 month under the exclusion of air and light, the oil is straw yellow; it has not formed any sediments.
  • the light oil contains ⁇ 1 mg / kg titratable nitrogen and oxygen. After 1 month of storage, it remains water-light.
  • Flash oil is distilled off from the sludge of the hot separator 6 in a system 7 for flash evaporation in a vacuum. This is combined with the sludge (2 kg / h) of the second hot separator 9, which consists mainly of oil, and is pumped via line 26 into the reactor outlet products before they reach the hot separator 6.
  • the top products of the second hot separator 9 are passed at 380 ° C. over 80 l of a Ni-Mo alumina catalyst in the reactor 12. After cooling to 280 ° C 14 154 kg / h of oil are obtained in the intermediate separator, which are used for the production of slurry.
  • the top product stream from the intermediate separator 14 is heated to 390 ° C and passed at 400 ° C over 80 l of a Co-Mo-Al2O3 catalyst in the reactor 16.
  • 54 kg / h of product oil with a basic nitrogen content of 10 mg / kg and a phenolic oxygen of 15 mg / kg are obtained from the cold separator 18.
  • the oil consists of 45% light oil; the rest is medium oil. After 1 month of storage there was a slight yellow discolouration of the originally water-light oil.
  • a hydrogenation test is carried out with a sub-bituminous coal.
  • 109 kg / h of the anhydrous coal (corresponding to 100 kg / h coal, free of water and ash) are mixed with 4 kg / h red mass and 87 kg / h return oil to a suspension, which is continuously pumped to the preheater 4 using the high-pressure pump 2 is promoted.
  • cycle gas containing 85% by volume of hydrogen is supplied in an amount of 150 m3 / h.
  • the entire flash oil (25 kg / h) is treated with 125 m3 / h fresh hydrogen at a temperature of 385 ° C and at 152 bar pressure.
  • the bottom phase reactor 5 has a volume of 200 l; it is operated at a temperature of 458 ° C and a pressure of 150 bar.
  • the products are broken down in the hot separator 6 at 450 ° C.
  • the catalyst is a commercially available hydrogenation catalyst which consists of tungsten and nickel sulfide on alumina support; the pressure in the reactor is 150 bar, the temperature is 390 ° C.
  • the products of reactor 12 are cooled to 330 ° C; 87 kg / h of a medium oil / heavy oil mixture can be taken from the following intermediate separator 14, which are used entirely as return oil.
  • the sludge from the hot separator 6 supplies 25 kg / h of flash oil when treated in a flash evaporator, which is sulfurized by saturation with hydrogen sulfide gas and is then hydrogenated and used as described.
  • the vapor and gaseous top product of the intermediate separator is heated from 330 ° C to 370 ° C and passed through a reactor with 80 kg of solid Ni-Mo-Al2O3 catalyst, the pressure still being 150 bar and the temperature set at 375 ° C becomes.
  • 56.5 kg / h of product oil are obtained from the cold separator, which consists of 40% light oil and 60% middle oil; the basic nitrogen content is 8 mg / kg, the phenolic oxygen content is approximately 15 mg / kg. After 1 month in the absence of air and light, the oil remains water-white.
  • Example 2 A test as in Example 2 is carried out, but the system does not contain a fixed bed reactor 12 after the hot separator 6 or a fixed bed reactor 16 after the intermediate separator 14.
  • the amount of fresh hydrogen is 100 m3 / h.
  • the hourly amount of distillate oil from the flash evaporation of the sludge is 30 kg.
  • 124 kg / h of oil are obtained per hour from the intermediate separator 14, combined with the blowdown distillate and continuously returned to the production of pulp.
  • the return oil contains 45% medium oil, the rest is heavy oil.
  • An hourly product oil quantity of 49.5 kg is obtained at the cold separator 18, which consists of 23% light oil and 77% medium oil.
  • the basic nitrogen content of the oil is 0.76% and the phenolic oxygen content is 2.7%. After a month, the initially yellowish oil is colored black. Without refining return oil and product oil, a lower yield is obtained with a much poorer oil quality.
  • Example 6 A test is carried out as in Example 6, but a fixed bed reactor 16 with 160 kg of catalyst charge is operated after the intermediate separator 14. The amount of fresh hydrogen is 125 m3 / h.
  • the product oil in the cold separator is 48.5 kg of water-bright and storage-stable oil per hour. It consists of 42% light oil and 58% medium oil. The basic nitrogen content is 12 mg / kg. The oil quality is therefore similar, but the yield is significantly lower than in the process according to the invention.
  • the first fixed bed reactor 12 contains 160 kg of catalyst; the second fixed bed reactor 16 is not operated.
  • the amount of fresh hydrogen is 125 m3 / h; the hourly amount of sludge distillate is 20 kg.
  • the oil gain from the cold separator 18 is 55 kg per hour.
  • the oil consists of 36% light oil and 64% medium oil.
  • the nitrogen content is 100 mg / kg. After a month, the originally colorless oil turned yellow. The yield is good in this mode of operation according to the prior art; however, the oil quality is insufficient.

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  • Chemical & Material Sciences (AREA)
  • Engineering & Computer Science (AREA)
  • Oil, Petroleum & Natural Gas (AREA)
  • Life Sciences & Earth Sciences (AREA)
  • Wood Science & Technology (AREA)
  • Chemical Kinetics & Catalysis (AREA)
  • General Chemical & Material Sciences (AREA)
  • Organic Chemistry (AREA)
  • Production Of Liquid Hydrocarbon Mixture For Refining Petroleum (AREA)
  • Catalysts (AREA)
EP86106751A 1985-06-03 1986-05-16 Procédé d'hydrogénation de charbon en phase liquide et en lit fixe avec catalyseur Expired EP0209665B1 (fr)

Applications Claiming Priority (2)

Application Number Priority Date Filing Date Title
DE3519830 1985-06-03
DE19853519830 DE3519830A1 (de) 1985-06-03 1985-06-03 Verfahren zur kohlehydrierung mit integrierten raffinationsstufen

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EP0209665A1 true EP0209665A1 (fr) 1987-01-28
EP0209665B1 EP0209665B1 (fr) 1988-10-12

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EP86106751A Expired EP0209665B1 (fr) 1985-06-03 1986-05-16 Procédé d'hydrogénation de charbon en phase liquide et en lit fixe avec catalyseur

Country Status (7)

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US (1) US4741822A (fr)
EP (1) EP0209665B1 (fr)
JP (1) JPH0784597B2 (fr)
AU (2) AU581990B2 (fr)
DE (2) DE3519830A1 (fr)
SU (1) SU1468427A3 (fr)
ZA (1) ZA864365B (fr)

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JPH11106757A (ja) * 1997-10-08 1999-04-20 Nippon Brown Coal Liquefaction Corp 石炭液化油の安定化方法
US7720677B2 (en) * 2005-11-03 2010-05-18 Coding Technologies Ab Time warped modified transform coding of audio signals
EP2125995B1 (fr) * 2006-09-18 2017-07-05 Jeffrey P. Newton Production d'hydrocarbures à bas poids moléculaire
CA2822875C (fr) * 2011-01-05 2018-02-20 Licella Pty Ltd. Traitement de matiere organique
CN106957681A (zh) * 2017-03-31 2017-07-18 北京中科诚毅科技发展有限公司 一种提高加氢反应体系氢分压的方法及其设计方法和用途

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EP0123161A1 (fr) * 1983-03-30 1984-10-31 VEBA OEL Entwicklungs-Gesellschaft mbH Procédé d'hydrogénation de charbon
DE3322730A1 (de) * 1983-06-24 1985-01-10 Ruhrkohle Ag, 4300 Essen Verfahren zur kohlehydrierung mit integrierter raffinationsstufe

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EP0151399B1 (fr) * 1984-01-19 1989-08-02 Ruhrkohle Aktiengesellschaft Distribution du gaz d'hydrogène dans des unités de liquéfaction de charbons
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DE3311356A1 (de) * 1983-03-29 1984-10-11 GfK Gesellschaft für Kohleverflüssigung mbH, 6600 Saarbrücken Verfahren zum hydrieren von kohle
EP0123161A1 (fr) * 1983-03-30 1984-10-31 VEBA OEL Entwicklungs-Gesellschaft mbH Procédé d'hydrogénation de charbon
DE3322730A1 (de) * 1983-06-24 1985-01-10 Ruhrkohle Ag, 4300 Essen Verfahren zur kohlehydrierung mit integrierter raffinationsstufe

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SU1468427A3 (ru) 1989-03-23
AU8658089A (fr) 1987-12-03
EP0209665B1 (fr) 1988-10-12
AU581990B2 (en) 1989-03-09
ZA864365B (en) 1986-12-09
JPS62285983A (ja) 1987-12-11
DE3519830A1 (de) 1986-12-18
DE3519830C2 (fr) 1993-07-22
DE3660919D1 (en) 1988-11-17
JPH0784597B2 (ja) 1995-09-13
US4741822A (en) 1988-05-03

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