EP0132526B1 - Procédé d'hydrogénation de charbon à étape de raffinage intégrée - Google Patents

Procédé d'hydrogénation de charbon à étape de raffinage intégrée Download PDF

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Publication number
EP0132526B1
EP0132526B1 EP84105406A EP84105406A EP0132526B1 EP 0132526 B1 EP0132526 B1 EP 0132526B1 EP 84105406 A EP84105406 A EP 84105406A EP 84105406 A EP84105406 A EP 84105406A EP 0132526 B1 EP0132526 B1 EP 0132526B1
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EP
European Patent Office
Prior art keywords
gas phase
boiling
oil
hydrogenation
intermediate 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
EP84105406A
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German (de)
English (en)
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EP0132526A3 (en
EP0132526A2 (fr
Inventor
Josef Dr. Rer.Nat. Langhoff
Eckard Dr.-Ing. Wolowski
Frank Dr.-Ing. Mirtsch
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RAG AG
Original Assignee
Ruhrkohle AG
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Publication of EP0132526A2 publication Critical patent/EP0132526A2/fr
Publication of EP0132526A3 publication Critical patent/EP0132526A3/de
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Publication of EP0132526B1 publication Critical patent/EP0132526B1/fr
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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 the production of liquid hydrocarbons from coal hydrogenation, with refined products having a comparatively low boiling point being produced in one operation.
  • the overall process can be optimized with regard to the desired product quality.
  • the service life and the optimal reaction conditions of the fixed bed catalyst in the refining stage play a decisive role for economical process control with high availability of the entire system.
  • the quality of the solvent required for mashing the coal is also important.
  • the other part is condensed directly after leaving the bottom phase hydrogenation (hot separator) and supplies the necessary remaining amounts of medium oil and heavy oil for the mashing of the coal.
  • a solvent is formed which is composed of a mixture of hydrogenated (in the gas phase reactor) and non-hydrogenated medium and heavy oils.
  • this process differs from other hydrogenation processes (e.g. EXXON process) in which all of the solvent is hydrated.
  • a disadvantage of the Raichle and Krönig process is that the directly condensed coal oil vapors also contain a considerable part of the light oil produced, which is obtained as a finished product as a finished product.
  • the invention has for its object to reduce the proportion of high-boiling oils in the raw hydrocarbon vapors.
  • the invention is based on the idea of using an intermediate separator.
  • Such intermediate separators are known per se for pure bottom phase hydrogenation (without downstream gas phase hydrogenation).
  • There the intermediate separator is located behind the hot separator.
  • the latter is operated with respect to temperature and pressure in such a way that the amount of solvent, consisting of medium and heavy oil, is obtained in the intermediate separator sump, which is necessary in order to be admixed as a partial flow to another partial flow originating from a vacuum column to ensure the solvent self-sufficiency of the bottom phase hydrogenation.
  • the intermediate separator sump which is necessary in order to be admixed as a partial flow to another partial flow originating from a vacuum column to ensure the solvent self-sufficiency of the bottom phase hydrogenation.
  • the crude coal oils from the bottom phase hydrogenation after leaving the hot separator head are divided by partial condensation in the intermediate separator into a high-boiling liquid fraction and a lower-boiling vapor fraction.
  • the low-boiling coal oil vapors, which drove over the gas phase reactor are made of light, medium and possibly a comparatively small amount of light heavy oil. This division can be varied by changing the intermediate separator temperature. This ensures that the bottom phase hydrogenation is combined with the gas phase hydrogenation in one working process, so that only the crude coal oils which tend to boil less (with only a little heavy oil content) are passed over the gas phase reactors.
  • the crude coal oils, which tend to be high boiling, are largely drawn off upstream of the gas phase reactor and serve as part of the solvent for the pulp mashing.
  • the result of this is that the gas phase reactor, on the one hand, has a better service life and optimal reaction conditions for the production of (partially) refined and low-boiling products and, on the other hand, is relieved of the amount of coal oil (especially heavy oil), which is required as a solvent for mashing the coal.
  • the reaction conditions for a catalyst can be adjusted more optimally if the starting products are not too broad.
  • the use for the gas phase reactor occurs in the desired boiling point.
  • the overall process is optimized in such a way that, on the one hand, there is an optimal refining and conversation of raw coal oils (light, medium and some heavy oil) to refined hydrocarbons with a lighter boiling point and, on the other hand, the service life and the reaction conditions for the catalyst in the gas phase reactor are optimized. Furthermore, in the event that more coal oils are passed through the gas phase reactor than the product quantity, only light and medium oils and the lighter boiling cuts of the heavy oil in the gas phase reactor are hydrated. As a result, the catalyst is gently stressed on the one hand and a hydrated solvent part consisting of medium oil and light heavy oil cuts is produced on the other hand. It is shown that the solvent quality that is decisive for the hydrogenation process is essentially determined by the type and amount of the middle oil and possibly the light heavy oil (e.g. donor effect of the relatively easily hydrolyzed middle oils and light heavy oil fractions N).
  • the part of the solvent thus produced and returned to the pulp mashing then consists on the one hand of the rehydrated solvent oil (free of heavy oil) with a higher boiling point) and on the other hand of the heavy-boiling intermediate separator sump product, which is not hydrated. This results in an improved solvent quality for the hydrogenation process.
  • the liquid portion in the intermediate separator still contains small amounts of light oil
  • stripping with hydrogen-containing gases and partial evaporation of the liquid portion and / or by flash evaporation of the above can be used.
  • Light oil components are largely separated and added to the use for gas phase hydrogenation.
  • FIG. 1 to 5 show flow diagrams of various operating modes according to the invention.
  • the products from the bottom phase hydrogenation 1 in the hot separator 2 are separated into a liquid / solid phase (bottom) and a gas / vapor phase (top) at about 450 ° C.
  • This gas / vapor phase which contains the actual coal oils, is partially cooled in heat exchangers 3, as a result of which the tendentially heavier boiling cuts of the coal oils largely condense out.
  • the liquid phase on the one hand and the gas / vapor phase on the other hand are separated in the intermediate separator 4 at 350 to 420 ° C.
  • the temperature of the intermediate separator 4 which determines the thermodynamic equilibrium and thus the separation of the coal oil into a lower-boiling vapor phase and a heavier-boiling liquid phase, can be varied by alternative connection of the insert product heat exchangers 3, which recover a large part of the waste heat of the products.
  • the hot separator head products are cooled in the heat exchanger 3 to the reaction temperature of the gas phase reactor 6.
  • a considerable part of the heavy oil (e.g. 70%) from the hot separator top product is produced in vapor form.
  • Almost all of the light oil and most of the middle oil are also produced in vapor form.
  • the bottom phase of the intermediate separator contains 15.8 kg of oils (1.5% light oil, 24% medium oil, 74.5% heavy oil), which are recirculated as a solvent component.
  • the top phase of the intermediate separator 4 (use for the gas phase reactor 6) consists of the hydrogenation gas from the bottom phase hydrogenation and 126 kg of oil vapors (14.5% light oil, 55.5% medium oil, 30% heavy oil).
  • the crude coal oils are refined at 390 ° C.
  • the product distribution is approx. 30% light oil, 43.5% medium oil and 26.5% heavy oil.
  • the coal oils from the gas phase reactor 6 are condensed out by cooling 7 and separated from the residual gases in the separator 8.
  • the pre-refined coal oils are divided into gasoline, medium oil and heavy oil. All gasoline and 22% of the middle oil are released as gas. All of the heavy oil and the remaining medium oil (78%) are recirculated as a solvent component, which is hydrated, for mashing the coal. Based on the total solvent, the hydrogenated solvent is 50%.
  • the hot separator head products are cooled in the heat exchanger 3 to an intermediate separator temperature which is below the reaction temperature of the gas phase reactor 6. This ensures that a large part of the heavy oil condenses out.
  • the gas / vapor phase contains comparatively little heavy oil and thus enables optimal reaction conditions for the lumpy catalyst in the gas phase reactor 6.
  • the gas / vapor phase is heated to the reaction temperature of the gas phase reactor 6 by heating 5.
  • the products and gases are fed to a separator 8 via a cooler 7.
  • the bottom phase of the intermediate separator contains 70.5 kg of oils (2.5% light oil, 40.5% medium oil, 57% Heavy oil), which are recirculated as a solvent component.
  • the top phase of the intermediate separator 4 (use for the gas phase reactor 6) consists of the hydrogenation gas from the bottom phase hydrogenation and 71 kg of oil vapors (23% light oil, 63.5% medium oil, 13.5% heavy oil), which in the heating 5 to the reaction temperature 390 ° C of the gas phase reactor 6 are heated.
  • the crude coal oils are refined at 390 ° C and 280 bar by means of refining on a fixed bed catalyst and partially converted to lighter boiling points.
  • the product distribution is approx. 34% light oil, 53.5% medium oil and 12.5% heavy oil.
  • the coal oils from the gas phase reactor 6 are condensed out by cooling 7 and separated from the residual gases in the separator 8.
  • the pre-refined coal oils are divided into gasoline, medium oil and heavy oil. All petrol and 63.5% of the middle oil are released as products. All of the heavy oil and the remaining medium oil (36.5%) are recirculated as a solvent component, which is hydrated, for mashing the coal. Based on the total solvent, the hydrogenated solvent is 15%.
  • the other modes of operation (c-f) represent modifications of the mode of operation b).
  • a small amount of light oil is also produced in the intermediate separator sump. In order to prevent this - albeit small - light oil portion from being returned to the bottom phase hydrogenation as a solvent, this light oil is largely separated from the intermediate separator bottom product and added to the use for the gas phase hydrogenation.
  • the light oil is separated from the intermediate separator sump product by partial evaporation and / or stripping with hydrogenation gas, cycle gas or fresh hydrogen (approx. 97% H 2 ).
  • the evaporation temperature which lies between the intermediate separator and gas phase reactor temperature, and the amount and quality of the stripping gas (e.g. hydrogenation cycle gas, fresh hydrogen, approx. 97% H 2 ) determine the amount of the low-boiling fractions to be evaporated.
  • the intermediate separator sump product can be heated, for example, by means of heat exchanger 5 (e.g. heat recovery of the waste heat from the hot separator head product) or in a heating furnace (e.g. parallel to the heating of the intermediate separator head products).
  • the gas / oil vapors are separated from the bottom product in a further separator 9 and fed to the feed for the gas phase hydrogenation.
  • the bottom product of the intermediate separator 4 at 330 ° - 340 ° C consists of 70.5 kg of oils, which still contain approx. 1.7 kg of light oil.
  • an oil vapor quantity of approx. 18 kg (1.3 kg light oil) results in the separator 9, which can be used in the gas phase reactor 6 can be slammed.
  • the separation of the light oil from the bottom product of the intermediate separator 4 is carried out according to FIG. 3 by relaxing the bottom product with subsequent distillative removal of the low-boiling fractions.
  • the distillation column 10 either only light oil or a mixture of light oil and medium oil can be drawn off, which is compressed again to process pressure by means of a high-pressure pump 11, heated and added to the insert for the gas phase hydrogenation.
  • the procedural justification for this procedure is that the light oil is completely separated from the bottom product of the intermediate separator; depending on the temperature of the intermediate separator 4, a two-phase flow can also be generated in the gas phase reactor 6 by adding light oil and medium oil if optimal reaction conditions in the gas phase reactor require this.
  • the boiling cut in the distillation 10 can be set so that not only the amount of product, but also a solvent fraction (medium oil and possibly heavy oil with a low boiling point) is passed through the gas phase reactor in order to achieve a desired solvent quality (increased hydrated fraction).
  • a solvent fraction medium oil and possibly heavy oil with a low boiling point
  • the bottom product of the intermediate separator 4 at 330-340 ° C consists of 70.5 kg of oils, which still contain approx. 1.7 kg of light oil.
  • the oils By relaxing the oils to atmospheric pressure, some of the oils evaporate, which, however, are converted back into the liquid phase by condensation. The gases released during the expansion of the oils are removed.
  • the light oil (1.7 kg) is completely separated from the residual oil (solvent content 68.8 kg) and fed to the use of gas phase hydrogenation via the pump 11 and the heating device 5. This creates a solvent that is practically free of light oil.
  • This method variant according to FIG. 4 is based on the driving style d.
  • the lighter fractions are separated from the sump by flash evaporation 12. After condensation of these lighter fractions and separation of the gases, they are compressed to high pressure in the liquid phase, heated and then used in the gas phase reactor.
  • the separation of the lighter fractions in the flash evaporation 12 can optionally be enhanced by stripping.
  • the bottom product of the intermediate separator 4 at 330-340 ° C consists of 70.5 kg of oils, which still contain approx. 1.7 kg of light oil.
  • the flash evaporator 12 By relaxing these oils to about atmospheric pressure in the flash evaporator 12, there is a division into 15.5 kg of oil vapors (1.5 kg of light oil) and 55 kg of oils (0.2 kg of light oil).
  • the 15.5 kg of oil vapors with 1.5 kg of light oil are condensed out, separated from the expansion and stripping gases and fed into gas phase hydrogenation via a high-pressure pump 11 and heating device 5.
  • This method variant according to FIG. 5 represents an expansion of the mode of operation e).
  • the lighter fractions are separated from the sump by the flash evaporation 12, possibly by supporting stripping gas. After condensation of these lighter fractions and separation of the gases, they are divided in a subsequent distillation 13 into a low-boiling fraction, which contains virtually all of the light oil, and a heavy-boiling fraction (solvent fraction).
  • the lower-boiling fraction is compressed to high pressure by means of compression 11, heated and added to the use of the gas phase reactor.
  • the 15.5 kg of oil vapors from the flash evaporator 12 consist of 1.5 kg of light oil and 14 kg medium / heavy oil.
  • the 1.5 kg of light oil are separated off, compressed, heated and added to the use of gas phase hydrogenation.
  • the remaining 14 kg medium / heavy oil are added to the solvent.

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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)
  • Organic Low-Molecular-Weight Compounds And Preparation Thereof (AREA)

Claims (8)

1. Procédé d'hydrogénation de charbon comprenant une hydrogénation en phase liquide et une hydrogénation en phase gazeuse prévue en aval, caractérisé en ce que les huiles de charbon provenant de l'hydrogénation en phase liquide sont, par une condensation partielle dans un séparateur intercalaire (4), partagées sous la pression du processus en une fraction liquide à ébullition difficile et en une fraction sous forme de vapeur, à ébullition plus aisée, en ce que la fraction sous forme de vapeur est conduite au-dessus de réacteurs à phase gazeuse et est convertie en produits (partiellement) raffinés à intervalle d'ébullition plus bas, et en ce que la fraction liquide, qui contient la plus grande part des fractions à ébullition difficile, est soutirée en amont du réacteur à phase gazeuse et est amenée au mouillage du charbon en tant que partie du solvant.
2. Procédé suivant la revendication 1, caractérisé en ce qu'on produit une fraction de solvant hydrogénée qui est constituée soit uniquement d'huile moyenne, soit d'un mélange d'huile moyenne et d'huile lourde à intervalle d'ébullition bas, et en ce que, quantitativement, il est conduit au-dessus du réacteur à phase gazeuse davantage d'huiles de charbon que cela ne correspond aux quantités de produit.
3. Procédé suivant la revendication 1, caractérisé en ce que la température du séparateur intercalaire correspond approximativement à la température réactionnelle des réacteurs à phase gazeuse (6) de 350 à 420°C.
4. Procédé suivant la revendication 1, caractérisé en ce que la température du séparateur intercalaire est située en dessous de la température réactionnelle des réacteurs à phase gazeuse et en ce que le produit de tête du séparateur intercalaire est, sous forme de vapeur, chauffé à la température réactionnelle des réacteurs à phase gazeuse, conjointement aux gaz, et est amené au réacteur à phase gazeuse.
5. Procède suivant la revendication 1, caractérisé en ce que la température du séparateur intercalaire est située en dessous de la température réactionnelle des réacteurs à phase gazeuse, en ce que l'huile légère qui est obtenue dans le produit de fond de cuve du séparateur intercalaire est, sous la pression du processus, séparée pour la plus grande part de la phase liquide par une évaporation partielle et/ou un strippage (avec du gaz de recyclage d'hydrogénation ou de l'hydrogène frais) et en ce qu'elle est ajoutée à la charge introduite dans l'hydrogénation en phase gazeuse et est chauffée à la température réactionnelle avant l'entrée dans les réacteurs à phase gazeuse.
6. Procédé suivant la revendication 1, caractérisé en ce que la température du séparateur intercalaire est située en dessous de la température réactionnelle des réacteurs à phase gazeuse, en ce que la phase liquide est détendue et est séparée des éléments à bas point d'ébullition dans une colonne de distillation fractionnée ou dans une simple colonne de distillation primaire, en ce que les éléments à bas point d'ébullition (produit de condensation) sont comprimés à la pression du processus, sont chauffés à la température réactionnelle et sont ajoutés à la charge introduite dans les réacteurs à phase gazeuse.
7. Procédé suivant la revendication 1, caractérisé en ce que les températures du séparateur intercalaire sont situées en dessous de la température réactionnelle des réacteurs à phase gazeuse, en ce que la phase liquide est séparée dans une évaporation de détente, avec éventuellement un strippage, en une phase vapeur à bas point d'ébullition et en une phase liquide à ébullition difficile (solvant), en ce que la phase vapeur à bas point d'ébullition est condensée, en ce qu'elle est separée, dans une distillation ulteriéure, en une fraction à bas point d'ébullition (autre fraction de solvant) et en ce que la fraction à bas point d'ébullition est comprimée, chauffée et amenée aux réacteurs à phase gazeuse.
8. Procédé suivant l'une des revendications 6 et 7, caractérisé en ce que, pour produire un écoulement diphasique dans la colonne de fractionnement ou la colonne de distillation primaire, des fractions à ébullition moyenne sont encore soutirées, en plus de l'huile légère, et elles sont ajoutées à la charge introduite dans les réacteurs à phase gazeuse.
EP84105406A 1983-06-24 1984-05-12 Procédé d'hydrogénation de charbon à étape de raffinage intégrée Expired EP0132526B1 (fr)

Applications Claiming Priority (2)

Application Number Priority Date Filing Date Title
DE3322730A DE3322730A1 (de) 1983-06-24 1983-06-24 Verfahren zur kohlehydrierung mit integrierter raffinationsstufe
DE3322730 1983-06-24

Publications (3)

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EP0132526A2 EP0132526A2 (fr) 1985-02-13
EP0132526A3 EP0132526A3 (en) 1986-06-04
EP0132526B1 true EP0132526B1 (fr) 1988-07-20

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EP84105406A Expired EP0132526B1 (fr) 1983-06-24 1984-05-12 Procédé d'hydrogénation de charbon à étape de raffinage intégrée

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US (1) US4602992A (fr)
EP (1) EP0132526B1 (fr)
JP (1) JPS6013885A (fr)
AU (1) AU557956B2 (fr)
BR (1) BR8403055A (fr)
CA (1) CA1231658A (fr)
DE (2) DE3322730A1 (fr)
PL (1) PL248358A1 (fr)
SU (1) SU1240364A3 (fr)
ZA (1) ZA844753B (fr)

Families Citing this family (3)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US4569749A (en) * 1984-08-20 1986-02-11 Gulf Research & Development Company Coal liquefaction process
DE3519830A1 (de) * 1985-06-03 1986-12-18 Ruhrkohle Ag, 4300 Essen Verfahren zur kohlehydrierung mit integrierten raffinationsstufen
WO2003093815A1 (fr) * 2002-05-01 2003-11-13 Exxonmobil Upstream Research Company Modele de rendements de structures et de compositions chimiques permettant de predire la composition d'hydrocarbures obtenus par thermolyse

Family Cites Families (9)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
JPS5351204A (en) * 1976-10-22 1978-05-10 Kobe Steel Ltd Conversion of coals
DE2654635B2 (de) * 1976-12-02 1979-07-12 Ludwig Dr. 6703 Limburgerhof Raichle Verfahren zur kontinuierlichen Herstellung von Kohlenwasserstoffölen aus Kohle durch spaltende Druckhydrierung
US4222844A (en) * 1978-05-08 1980-09-16 Exxon Research & Engineering Co. Use of once-through treat gas to remove the heat of reaction in solvent hydrogenation processes
US4283267A (en) * 1978-05-11 1981-08-11 Exxon Research & Engineering Co. Staged temperature hydrogen-donor coal liquefaction process
US4266083A (en) * 1979-06-08 1981-05-05 The Rust Engineering Company Biomass liquefaction process
DE3022158C2 (de) * 1980-06-13 1989-11-02 Bergwerksverband Gmbh, 4300 Essen Verfahren zur hydrierenden Kohleverflüssigung
US4400263A (en) * 1981-02-09 1983-08-23 Hri, Inc. H-Coal process and plant design
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
DE3209143A1 (de) * 1982-03-13 1983-09-22 Veba Oel Entwicklungsgesellschaft mbH, 4660 Gelsenkirchen-Buer Verfahren zur mehrstufigen hydrierung von kohle

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BR8403055A (pt) 1985-05-28
AU2970284A (en) 1985-01-31
EP0132526A3 (en) 1986-06-04
CA1231658A (fr) 1988-01-19
ZA844753B (en) 1985-05-29
DE3472800D1 (en) 1988-08-25
PL248358A1 (en) 1985-04-24
US4602992A (en) 1986-07-29
SU1240364A3 (ru) 1986-06-23
DE3322730A1 (de) 1985-01-10
EP0132526A2 (fr) 1985-02-13
AU557956B2 (en) 1987-01-15
JPS6013885A (ja) 1985-01-24

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