EP0209665B1 - Verfahren zur Kohlehydrierung mittels Sumpfphase- und katalysator-Festbetthydrierung - Google Patents

Verfahren zur Kohlehydrierung mittels Sumpfphase- und katalysator-Festbetthydrierung Download PDF

Info

Publication number
EP0209665B1
EP0209665B1 EP86106751A EP86106751A EP0209665B1 EP 0209665 B1 EP0209665 B1 EP 0209665B1 EP 86106751 A EP86106751 A EP 86106751A EP 86106751 A EP86106751 A EP 86106751A EP 0209665 B1 EP0209665 B1 EP 0209665B1
Authority
EP
European Patent Office
Prior art keywords
oil
separator
coal
hydrogenation
hot separator
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.)
Expired
Application number
EP86106751A
Other languages
German (de)
English (en)
French (fr)
Other versions
EP0209665A1 (de
Inventor
Bernd Strobel
Frank Dr. Dipl.-Chem. Friedrich
Eckhard Wolowski
Rainer Löring
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.)
RAG AG
Original Assignee
Ruhrkohle AG
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 Ruhrkohle AG filed Critical Ruhrkohle AG
Publication of EP0209665A1 publication Critical patent/EP0209665A1/de
Application granted granted Critical
Publication of EP0209665B1 publication Critical patent/EP0209665B1/de
Expired legal-status Critical Current

Links

Images

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
    • 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, if appropriate, the catalyst, also the asphalt-containing, non-distillable products of the coal hydrogenation and small to medium amounts of distillable oil.
  • This oil is obtained by distillation or by other measures, such as. B. smoldering or solvent extraction, recovered from the sludge 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-A-2 654 635 also have the disadvantage that oils from all boiling positions 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 system, 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 before 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 sensible 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 a second intermediate separation is carried out at a further reduced temperature after the return oil has been extracted from the intermediate separator.
  • 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 which is practically completely free of heteroatoms (boiling up to 185 ° C.) is obtained from the cold separator.
  • 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 often, the solubility losses in pressurized hydrogen less. The high 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; the same or different catalysts can be used in the reactors 12 and 16, respectively, in order to improve the degree of refinement, saturation, cleavage and hydrogen consumption. B. to achieve the most favorable results for any 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 an amount of oil is continuously condensed as is consumed in the production of the 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 content of the product oil is very high must be largely refined, but the middle oil content 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 from 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 13, which mainly consists of light oil (boiling end 185 ° C.) and has the 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 hydrogenated in the fixed bed reactor 25 at 350 to 420 ° C. under approximately the same pressure as in the other reactors.
  • the addition of hydrogen to flash oil is in the range from 0.5 to 5 m 3 / 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 plant according to Example 1, 126 kg of water-free gas flame 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 m 3 / h of hydrogenation gas (gas quantities below Normal conditions) consisting of 150 m 3 / h of fresh hydrogen and 500 m 3 / h of cycle gas (contains 60% by vote hydrogen) passed through a bottom phase reactor 5.
  • the pressure in 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 Al3O3-Si0 2 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.
  • 105 kg of dry gas coal (corresponding to 100 kg / h, free of water and ash) are mixed with 4 kg / h of dried red mass and 154 kg / h of return oil to form a coal paste and together with 625 m 3 / h hydrogenation gas, consisting of 125 m 3 / h fresh hydrogen and 500 m 3 / h cycle gas (contains 80 vol .-% hydrogen), passed through the bottom phase reactor 5 of 200 l content.
  • the pressure in the reactor 5 is 300 bar, the temperature 470 ° C.
  • the temperature in the hot separator 6 is kept at 440 ° C.
  • 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 are drawn off at the top of the intermediate separator 14 are passed over the fixed bed reactor 16, which also contains 80 kg of a commercially available refining catalyst based on Ni, Mo and A1 2 0 3 .
  • 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 I / 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-Al 2 0 3 refining catalyst. The average temperature of the catalyst bed is kept at 375 ° C.
  • the reactor outlet products are cooled to 20 ° C; 22 kg of light oil per hour with a boiling end of 185 ° C are obtained from the cold separator.
  • the medium oil contains 0.06% basic nitrogen and ⁇ 0.1% oxygen. After 1 month of storage in the air and in the dark, 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 overhead product stream from the intermediate separator 14 is heated to 390 ° C. and passed at 400 ° C. over 80 l of a Co-Mo-Al 2 O 3 catalyst in the reactor 16.
  • the products of the Reactor 16 is obtained from the cold separator 18 54 kg / h product oil with a basic nitrogen content of 10 mg / kg and phenolic oxygen of 15 mg / kg.
  • 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.
  • recycle gas containing 85% by volume of hydrogen is supplied in an amount of 150 m 3 / h.
  • the entire flash oil (25 kg / h) is treated with 125 m 3 / h of 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 the reactor 12 are cooled to 330 ° C.; can be removed from the subsequent intermediate separator 14, 87 kg / h of a medium oil / heavy oil mixture, 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-A1 2 0 3 catalyst, the pressure still being 150 bar and the temperature at 375 ° C is set.
  • 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 fish hydrogen is 100 m 3 / 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 m 3 / 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 m 3 / 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.
EP86106751A 1985-06-03 1986-05-16 Verfahren zur Kohlehydrierung mittels Sumpfphase- und katalysator-Festbetthydrierung Expired EP0209665B1 (de)

Applications Claiming Priority (2)

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

Publications (2)

Publication Number Publication Date
EP0209665A1 EP0209665A1 (de) 1987-01-28
EP0209665B1 true EP0209665B1 (de) 1988-10-12

Family

ID=6272296

Family Applications (1)

Application Number Title Priority Date Filing Date
EP86106751A Expired EP0209665B1 (de) 1985-06-03 1986-05-16 Verfahren zur Kohlehydrierung mittels Sumpfphase- und katalysator-Festbetthydrierung

Country Status (7)

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

Families Citing this family (5)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
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 (en) * 2006-09-18 2017-07-05 Jeffrey P. Newton Production of lower molecular weight hydrocarbons
CN103384713B (zh) * 2011-01-05 2015-08-19 莱斯拉有限公司 有机质的加工
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

Family Cites Families (15)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US3726784A (en) * 1971-02-18 1973-04-10 Exxon Research Engineering Co Integrated coal liquefaction and hydrotreating process
DE2654635B2 (de) * 1976-12-02 1979-07-12 Ludwig Dr. 6703 Limburgerhof Raichle Verfahren zur kontinuierlichen Herstellung von Kohlenwasserstoffölen aus Kohle durch spaltende Druckhydrierung
US4330391A (en) * 1976-12-27 1982-05-18 Chevron Research Company Coal liquefaction process
US4338182A (en) * 1978-10-13 1982-07-06 Exxon Research & Engineering Co. Multiple-stage hydrogen-donor coal liquefaction
DE3105030A1 (de) * 1981-02-12 1982-09-02 Basf Ag, 6700 Ludwigshafen Verfahren zur kontinuierlichen herstellung von kohlenwasserstoffoelen aus kohle durch druckhydrierung in zwei stufen
ZA822056B (en) * 1981-08-05 1983-02-23 Lummus Co Coal liquefaction
US4400261A (en) * 1981-10-05 1983-08-23 International Coal Refining Company Process for coal liquefaction by separation of entrained gases from slurry exiting staged dissolvers
DE3209143A1 (de) * 1982-03-13 1983-09-22 Veba Oel Entwicklungsgesellschaft mbH, 4660 Gelsenkirchen-Buer Verfahren zur mehrstufigen hydrierung von kohle
DE3311552A1 (de) * 1983-03-30 1984-10-04 Veba Oel Entwicklungsgesellschaft mbH, 4660 Gelsenkirchen-Buer Verfahren zur hydrierung von kohle
DE3311356A1 (de) * 1983-03-29 1984-10-11 GfK Gesellschaft für Kohleverflüssigung mbH, 6600 Saarbrücken Verfahren zum hydrieren von kohle
DE3322730A1 (de) * 1983-06-24 1985-01-10 Ruhrkohle Ag, 4300 Essen Verfahren zur kohlehydrierung mit integrierter raffinationsstufe
EP0151399B1 (de) * 1984-01-19 1989-08-02 Ruhrkohle Aktiengesellschaft Hydriergasführung in Kohleverflüssigungsanlagen
DE3402264A1 (de) * 1984-01-24 1985-08-01 Basf Ag, 6700 Ludwigshafen Verfahren zur kontinuierlichen herstellung von kohlenwasserstoffoelen durch spaltende druckhydrierung
US4596650A (en) * 1984-03-16 1986-06-24 Lummus Crest, Inc. Liquefaction of sub-bituminous coal
DE3420197A1 (de) * 1984-05-30 1985-12-12 Ruhrkohle Ag, 4300 Essen Verfahren zur herstellung eines dieselkraftstoffes aus kohlemitteloel

Also Published As

Publication number Publication date
JPS62285983A (ja) 1987-12-11
ZA864365B (en) 1986-12-09
DE3519830C2 (ja) 1993-07-22
EP0209665A1 (de) 1987-01-28
AU8658089A (ja) 1987-12-03
AU581990B2 (en) 1989-03-09
JPH0784597B2 (ja) 1995-09-13
DE3660919D1 (en) 1988-11-17
SU1468427A3 (ru) 1989-03-23
US4741822A (en) 1988-05-03
DE3519830A1 (de) 1986-12-18

Similar Documents

Publication Publication Date Title
EP0073355B1 (de) Verfahren zur Herstellung flüssiger Kohlenwasserstoffe
DE2654635B2 (de) Verfahren zur kontinuierlichen Herstellung von Kohlenwasserstoffölen aus Kohle durch spaltende Druckhydrierung
DE1768566A1 (de) Kohlehydrierverfahren
DE2909103A1 (de) Verfahren zur kohleverfluessigung
DE1119438B (de) Verfahren zum Raffinieren von schwefelhaltigen schweren OElen
DD147851A5 (de) Integriertes kohleverfluessigungs-vergasungsverfahren
DE2733186A1 (de) Kohleverfluessigungsverfahren
DE2622426C2 (ja)
DD251781A5 (de) Verfahren zur behandlung von rueckstaenden von schweroel auf erdoelbasis
DE2635388A1 (de) Kohleverfluessigungsverfahren
CH542276A (de) Verfahren zur Trennung eines heissen gemischtphasigen Konversionsproduktes
DE3033259A1 (de) Verfahren zur loesungsmittelraffination von kohle
DE1770575A1 (de) Verfahren zum Hydrieren von Kohlenwasserstoffen
DE3022581A1 (de) Verfahren zur kohleverfluessigung
EP0209665B1 (de) Verfahren zur Kohlehydrierung mittels Sumpfphase- und katalysator-Festbetthydrierung
DE2824062A1 (de) Kohleverfluessigungsverfahren
DE3529795A1 (de) Verfahren zur verfluessigung von kohle
DE2834712C2 (de) Verfahren zur Rückführung des bei der Herstellung von Synthesegas in wäßriger Suspension anfallenden Rußes in den Gasgenerator
DD147679A5 (de) Kohleverfluessigungsverfahren mit verbessertem schlammrueckfuehrungssystem
DD147676A5 (de) Kohleverfluessigungsverfahren mit zusatz fremder mineralien
DE3236504C2 (ja)
DD158794A5 (de) Verfahren zur herstellung einer normalerweise festen geloesten kohle
EP0027962B1 (de) Verfahren zur Herstellung flüssiger Kohlenwasserstoffe aus Kohle
DE3022158C2 (de) Verfahren zur hydrierenden Kohleverflüssigung
EP0166858B1 (de) Verfahren zur Herstellung eines Dieselkraftstoffes aus Kohlemittelöl

Legal Events

Date Code Title Description
PUAI Public reference made under article 153(3) epc to a published international application that has entered the european phase

Free format text: ORIGINAL CODE: 0009012

17P Request for examination filed

Effective date: 19861129

AK Designated contracting states

Kind code of ref document: A1

Designated state(s): BE DE FR GB NL

17Q First examination report despatched

Effective date: 19880308

GRAA (expected) grant

Free format text: ORIGINAL CODE: 0009210

AK Designated contracting states

Kind code of ref document: B1

Designated state(s): BE DE FR GB NL

GBT Gb: translation of ep patent filed (gb section 77(6)(a)/1977)
REF Corresponds to:

Ref document number: 3660919

Country of ref document: DE

Date of ref document: 19881117

ET Fr: translation filed
PLBE No opposition filed within time limit

Free format text: ORIGINAL CODE: 0009261

STAA Information on the status of an ep patent application or granted ep patent

Free format text: STATUS: NO OPPOSITION FILED WITHIN TIME LIMIT

26N No opposition filed
PGFP Annual fee paid to national office [announced via postgrant information from national office to epo]

Ref country code: GB

Payment date: 19930413

Year of fee payment: 8

PG25 Lapsed in a contracting state [announced via postgrant information from national office to epo]

Ref country code: GB

Effective date: 19940516

GBPC Gb: european patent ceased through non-payment of renewal fee

Effective date: 19940516

PGFP Annual fee paid to national office [announced via postgrant information from national office to epo]

Ref country code: BE

Payment date: 19950530

Year of fee payment: 10

PGFP Annual fee paid to national office [announced via postgrant information from national office to epo]

Ref country code: NL

Payment date: 19950531

Year of fee payment: 10

PGFP Annual fee paid to national office [announced via postgrant information from national office to epo]

Ref country code: FR

Payment date: 19960515

Year of fee payment: 11

PG25 Lapsed in a contracting state [announced via postgrant information from national office to epo]

Ref country code: BE

Effective date: 19960531

BERE Be: lapsed

Owner name: RUHRKOHLE A.G.

Effective date: 19960531

PG25 Lapsed in a contracting state [announced via postgrant information from national office to epo]

Ref country code: NL

Effective date: 19961201

NLV4 Nl: lapsed or anulled due to non-payment of the annual fee

Effective date: 19961201

PG25 Lapsed in a contracting state [announced via postgrant information from national office to epo]

Ref country code: FR

Free format text: LAPSE BECAUSE OF NON-PAYMENT OF DUE FEES

Effective date: 19980130

REG Reference to a national code

Ref country code: FR

Ref legal event code: ST

PGFP Annual fee paid to national office [announced via postgrant information from national office to epo]

Ref country code: DE

Payment date: 20020521

Year of fee payment: 17

PG25 Lapsed in a contracting state [announced via postgrant information from national office to epo]

Ref country code: DE

Free format text: LAPSE BECAUSE OF NON-PAYMENT OF DUE FEES

Effective date: 20031202