WO2006010330A1 - A process for direct liquefaction of coal - Google Patents

A process for direct liquefaction of coal Download PDF

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Publication number
WO2006010330A1
WO2006010330A1 PCT/CN2005/001132 CN2005001132W WO2006010330A1 WO 2006010330 A1 WO2006010330 A1 WO 2006010330A1 CN 2005001132 W CN2005001132 W CN 2005001132W WO 2006010330 A1 WO2006010330 A1 WO 2006010330A1
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WIPO (PCT)
Prior art keywords
coal
reaction
oil
liquefaction
catalyst
Prior art date
Application number
PCT/CN2005/001132
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English (en)
French (fr)
Inventor
Yuzhuo Zhang
Geping Shu
Jialu Jin
Minli Cui
Xiuzhang Wu
Xiangkun Ren
Yaowu Xu
Shipu Liang
Jianwei Huang
Ming Yuan
Juzhong Gao
Yufei Zhu
Original Assignee
Shenhua Group Corporation Limited
China Shenhua Coal Liquefaction Corporation
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Publication date
Application filed by Shenhua Group Corporation Limited, China Shenhua Coal Liquefaction Corporation filed Critical Shenhua Group Corporation Limited
Priority to EP05771295.2A priority Critical patent/EP1783194B1/en
Priority to ES05771295.2T priority patent/ES2540745T3/es
Priority to JP2007522903A priority patent/JP4866351B2/ja
Priority to US11/572,638 priority patent/US7763167B2/en
Priority to AU2005266712A priority patent/AU2005266712B2/en
Priority to CA2575445A priority patent/CA2575445C/en
Priority to UAA200702177A priority patent/UA83585C2/uk
Priority to PL05771295T priority patent/PL1783194T3/pl
Publication of WO2006010330A1 publication Critical patent/WO2006010330A1/zh

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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
    • 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
    • C10G1/06—Production of liquid hydrocarbon mixtures from oil-shale, oil-sand, or non-melting solid carbonaceous or similar materials, e.g. wood, coal by destructive hydrogenation
    • C10G1/065—Production of liquid hydrocarbon mixtures from oil-shale, oil-sand, or non-melting solid carbonaceous or similar materials, e.g. wood, coal by destructive hydrogenation in the presence of a solvent
    • 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
    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/10—Feedstock materials
    • C10G2300/1074—Vacuum distillates
    • 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
    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/10—Feedstock materials
    • C10G2300/1077—Vacuum residues
    • 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
    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/20—Characteristics of the feedstock or the products
    • C10G2300/30—Physical properties of feedstocks or products
    • C10G2300/301—Boiling range
    • 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
    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/40—Characteristics of the process deviating from typical ways of processing
    • C10G2300/4081—Recycling aspects
    • 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
    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/40—Characteristics of the process deviating from typical ways of processing
    • C10G2300/42—Hydrogen of special source or of special composition
    • 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
    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/40—Characteristics of the process deviating from typical ways of processing
    • C10G2300/44—Solvents

Definitions

  • the invention relates to a method of direct liquefaction of coal. Background technique
  • the liquefaction process used in Germany was a bubbling bed reactor, the solid-liquid separation was carried out by pressure filtration or centrifugal separation, and the iron-containing natural mineral catalyst was used.
  • the circulating solvent due to pressure filtration and centrifugation contained difficult-to-react asphalt, plus The liquefaction catalyst has low activity, so the liquefaction conditions are very severe, the reaction pressure is 70 MPa, and the reaction condition is 480 °C.
  • the H - COAL process was developed in the United States in the early 1980s.
  • the H-Linux process uses a forced circulation suspension bed reactor with a reaction pressure of 20 MPa and a reaction temperature of 455 °C.
  • the liquefaction catalyst is a nickel-molybdenum or cobalt-molybdenum/alumina carrier petroleum-based hydrogenation catalyst, which is separated by hydrocyclone and decompressed.
  • the circulating solvent was separated by distillation.
  • the process is easy to control the reaction temperature and stable in product properties due to the use of a forced circulation suspension bed reactor and a petroleum-based hydrogenation catalyst.
  • the petroleum-based hydrogenation catalyst is deactivated in the coal liquefaction reaction system, the catalyst renewal period is short, and the liquefied oil production cost is high.
  • the IGOR+ process uses a bubbling bed reactor, a vacuum distillation separation of the circulating solvent, and an on-line fixed-bed hydrogenation reactor for circulation
  • the solvent and product are hydrogenated at different depths, and the liquefaction catalyst is red mud. Since the process adopts hydrogenation-recycling solvent after hydrogenation, the coal slurry has stable properties, high coal slurry concentration, easy preheating, and heat exchange with the high-temperature separator gas phase, and high heat utilization rate.
  • reaction pressure 30 MPa reaction pressure 30 MPa
  • reaction temperature 470 ° C reaction temperature 470 ° C
  • the on-line fixed-bed hydrogenation reactor has the risk of short catalyst deactivation inactivation cycle
  • bubbling bed reactor pair High calcium coal has mineral deposits.
  • the NEDOL process employs a bubble-bed reactor, a vacuum distillation separation solvent, and an off-line fixed bed hydrogenation reactor for hydrogenation of the circulating solvent using ultrafinely pulverized natural pyrite (0.7 ⁇ ). Since the process adopts hydrogenation-recycling solvent after hydrogenation, the coal slurry has stable properties, high coal slurry concentration, easy preheating, and heat exchange with the high-temperature separator gas phase, high heat utilization rate, mild reaction conditions, typical The operating conditions were a reaction pressure of 17 MPa and a reaction temperature of 450 °C.
  • the natural pyrite has high hardness, difficult to be ultra-fine pulverized, and high in cost; the gas retention coefficient of the bubbling bed reactor is large, and the utilization rate is low; the bubbling bed reactor has a low liquid velocity and mineral deposit; the fixed bed hydrogenation reactor There is a risk of a short operating cycle.
  • the object of the present invention is to provide a direct coal liquefaction method, which can stably operate for a long period of time, has high reactor utilization rate, prevents mineral deposition, moderates reaction conditions, maximizes liquid yield, and simultaneously provides high quality raw materials for further processing of liquefied products. .
  • the invention comprises the following steps:
  • the process for preparing the coal slurry comprises the following steps: (a) drying and pulverizing the raw coal through a coal pretreatment device to prepare a coal powder of a certain particle size; (b) preparing the coal powder with a catalyst raw material in a catalyst The preparation device is formed into a catalyst of ultrafine particles; (c) the coal slurry is prepared by mixing the coal powder with the catalyst in a coal slurry preparation device with a hydrogen supply solvent.
  • the step of liquefying the reaction comprises the following steps: (a) mixing the coal slurry with hydrogen and then preheating into the first forced circulation suspension bed reactor for reaction; (b) suspending the first forced circulation The outlet material of the bed reactor is replenished with hydrogen and sent to a second forced circulation suspension bed reactor for reaction.
  • the conditions of the liquefaction reaction are:
  • the gas-liquid separation process comprises the following steps: (a) feeding the reaction material into a high temperature separator for gas-liquid separation, wherein the high temperature separator is controlled at a temperature of 420 ° C; (b) the gas phase of the high temperature separator Part of it is sent to a cryogenic separator for further gas-liquid separation, wherein the temperature of the cryogenic separator is room temperature.
  • reaction conditions for the hydrogenation are:
  • the hydrogen-donating circulating solvent is a product obtained by hydrogenating a coal direct liquefied oil. Its distillation range is 220 - 450 °C.
  • the residue of the vacuum distillation column has a solid content of 50 - 55 wt%.
  • the distillation product of the atmospheric distillation column overhead oil and the vacuum distillation column decompression oil has a distillation range of C5 - 530 °C.
  • the forced circulation suspended bed hydrogenation reactor is a reactor having an internal member, a circulation pump at the bottom, and a catalyst replacement.
  • the invention provides a coal direct liquefaction method, which adopts a high active liquefaction catalyst, a hydrogen supply circulating solvent, a forced circulation suspension bed reactor, a vacuum distillation separation asphalt and a solid, and a forced circulation suspension bed hydrogenation reactor, which can be stably stabilized for a long period of time. Operation, high reactor utilization, prevention of mineral shield deposition, mild reaction conditions, maximum liquid yield, and high quality raw materials for further processing of liquefied products.
  • Figure 1 is a flow chart of a direct coal liquefaction process of the present invention. detailed description
  • the figures in the figure represent: 1. Raw coal, 2. Coal pretreatment device, 3. Catalyst raw material, 4. Catalyst preparation device, 5. Coal slurry preparation device, 6. Hydrogen, 7. First forced circulation suspended bed reaction , 8, second forced circulation suspended bed reactor, 9, high temperature separator, 10, cryogenic separator, 11, atmospheric distillation tower, 12, vacuum distillation column, 13, liquefied oil forced circulation suspension bed hydrogenation reaction , 14, gas-liquid separator, 15, product fractionation tower, 16, hydrogen supply solvent.
  • the liquefied raw coal 1 is dried and pulverized by the coal pretreatment device 2 to produce pulverized coal of a certain particle size.
  • the catalyst raw material 3 is passed through the catalyst preparation device 4 to form an ultrafine particle catalyst.
  • the coal powder and the catalyst are mixed in the coal slurry preparation device 5 and the hydrogen supply solvent 16 to form a coal slurry.
  • the coal slurry is mixed with the hydrogen gas 6 and then preheated into the first forced circulation suspended bed reactor 7, and the first forced circulation suspended bed reactor 7 outlet material is supplied with hydrogen to the second forced circulation suspended bed reactor 8.
  • the second forced circulation suspended bed reactor 8 reacts the material into the high temperature separator 9
  • the gas-liquid separation was carried out, and the high temperature separator 9 controlled the temperature at 420 °C.
  • the gas phase portion of the high temperature separator 9 enters the low temperature separator 10 for further gas-liquid separation, and the temperature of the low temperature separator 10 is room temperature.
  • the gas phase portion of the cryogenic separator 10 is mixed with hydrogen for recycling, and the exhaust gas portion is discharged from the system.
  • the liquid phase portion of the high temperature separator 9 and the cryogenic separator 10 enters the atmospheric distillation column 11 to separate the light fraction, and the bottom portion of the atmospheric distillation column 11 enters the vacuum distillation column 12 for the removal of the pitch and the solid, and the vacuum distillation is carried out.
  • the bottom material of the tower 12 is the liquefied residue. In order to ensure that the residue can be smoothly removed at a certain temperature, the solid content in the residue is generally controlled to be 50 - 55 wt%.
  • the distillate oil of the atmospheric distillation column 11 and the vacuum distillation column 12 is mixed with the hydrogen gas 6 into the forced circulation suspension bed hydrogenation reactor 13 to carry out catalytic hydrogenation for the purpose of improving the hydrogen supply performance of the solvent. Due to the high content of condensed aromatic hydrocarbons and heteroatoms in the liquefied oil, the composition is complicated, and the catalyst is easily deactivated by carbon deposition. After the forced circulation suspension bed reactor, the catalyst can be periodically updated, the operation cycle can be extended indefinitely, and the fixed bed reaction is also avoided. The risk of increasing the pressure difference of the catalyst deposit.
  • the forced circulation suspended bed hydrogenation reactor 13 outlet material enters the gas-liquid separator 14 for gas-liquid separation, and the gas phase portion of the gas-liquid separator 14 is mixed with hydrogen for recycling, and the exhaust gas portion is discharged to the system.
  • the liquid-liquid material of the gas-liquid separator 14 enters the product fractionation column 15, and the product and the hydrogen-donating solvent 16 are fractionated.
  • the products are all gasoline and diesel fractions.
  • the pulverized coal is lignite or young bituminous coal having a water content of 0.5 to 4.0% by weight and a particle size of ⁇ 0.15 mm.
  • ⁇ -FeOOH ultrafine ⁇ -hydrated iron oxide
  • the hydrogen-donating circulating solvent used is a product obtained by hydrogenating a coal direct liquefied oil, and has a distillation range of 220 to 450 °C. Due to the pre-hydrogenation of the circulating solvent, the solvent is stable and the slurry is good. It can be prepared into a high concentration coal slurry with a solid concentration of 45 - 55wt ° / ⁇ , and the coal slurry has good fluidity.
  • the viscosity of the coal slurry is less than 400CP ( 60 ° C) Because of the pre-hydrogenation of the circulating solvent, the hydrogen supply performance of the solvent is good, and the high-activity liquefaction catalyst is added, the liquefaction reaction condition is mild, the reaction pressure is 17 - 19 MPa, and the reaction temperature is 440 - 465 °C. Due to the pre-addition of circulating solvent Hydrogen, solvent has hydrogen supply performance, can prevent condensation of free radical fragments during coal thermal decomposition during coal preheating and heat exchange, prevent coking, prolong operation cycle and improve heat utilization.
  • the forced circulation suspended bed reactor used has a low gas retention coefficient, a high liquid phase utilization rate of the reactor, and a high liquid velocity due to a forced circulation pump, and no mineral deposits in the reactor.
  • two forced circulation suspended bed reactors are employed. Due to the full back-mixed flow in the forced circulation suspended bed reactor, the axial temperature distribution is uniform, the reaction temperature is easy to control, the reaction temperature can be controlled by the feed temperature, and the reactor side line quenching hydrogen control is not required, and the product property is stable. Due to the low gas retention coefficient of the forced circulation suspended bed reactor, the liquid phase utilization of the reactor is high; due to the high liquid velocity in the forced circulation suspended bed reactor, there is no mineral deposition in the reactor.
  • the removal of bitumen and solids is carried out by vacuum distillation.
  • Vacuum distillation is a mature and effective separation method for removing bitumen and solids.
  • the distillate distilled under reduced pressure does not contain bitumen. It can provide qualified raw materials for hydrogenation of circulating solvent, and the residue of vacuum distillation. Containing solids 50 - 55wt%; Due to the use of highly active liquefaction catalysts, the amount of addition is small, the oil content in the residue is small, and there are many diesel fractions in the product.
  • the recycle solvent and product are hydrogenated using a forced recycle suspension bed hydrogenation reactor.
  • the forced circulation suspension bed hydrogenation reactor adopts the upflow type, the catalyst can be periodically renewed, the hydrogen supply solvent after hydrogenation has good hydrogen supply performance, the product property is stable, the operation cycle can be extended indefinitely, and the fixed bed reaction is also avoided due to the catalyst. The risk of an increase in the pressure difference of the deposit.
  • the following is a liquefaction result of direct liquefaction of a young bituminous coal using a preferred embodiment of the present invention.
  • test conditions are:
  • Reactor temperature first reactor 455 ° C, second reactor 455 ° C.
  • Reaction pressure First reactor 19.0 MPa, second reactor 19.0 MPa.
  • Coal slurry concentration 45/55 (dry coal / solvent, mass ratio).
  • Catalyst addition amount Liquefaction catalyst: 1.0 wt% (/dry coal).
  • S/Fe 2 (molar ratio).
  • Table 1 shows the liquefaction results of a young bituminous coal in a continuous coal liquefaction process test apparatus of the present invention (the data in the table is based on dry ashless base coal).
  • Table 2 shows the liquefaction results of the same young bituminous coal in a direct liquefaction process continuous test unit (the data in the table is based on dry ash-free coal).
  • Table 1 Liquefaction results of a young bituminous coal on the coal direct liquefaction continuous test device of the present invention

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  • Engineering & Computer Science (AREA)
  • Oil, Petroleum & Natural Gas (AREA)
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Description

一种煤炭直接液化的方法 技术领域
本发明涉及一种煤炭直接液化的方法。 背景技术
1913年， 德国的柏吉乌斯 （Bergius ) 进行了从煤或煤焦油通过高 温高压加氢生产液体燃料研究并获得世界上第一个煤直接液化专利， 从而为煤的直接液化奠定了基础。 1927年，德国 I.GFarbenindustrie (燃 料公司）在 Leuna建立了世界上第一个煤直接液化厂。 二次大战期间， 德国一共有 12 套直接液化装置建成投产， 生产能力达到 423 >< 104t/a (吨 /年），为发动第二次世界大战的德国提供了 2/3的航空燃料和 50% 的汽车和装甲车用油。 当时德国采用的液化工艺为鼓泡床反应器， 固 液分离采用加压过滤或离心分离工艺， 采用含铁天然矿物催化剂， 由 于加压过滤和离心分离的循环溶剂含有难反应的沥青， 加上液化催化 剂活性低， 所以液化条件十分苛刻, 反应压力为 70MPa, 反应条件为 480°C。
二战结束后， 德国煤直接液化工厂全部关闭。 七十年代初石油危 机后， 寻找石油替代品引起了各工业发达国家的高度重视， 研究开发 了许多煤炭直接液化新工艺。
美国于 80年代初开发了 H - COAL工艺。 H - COAL工艺采用强 制循环悬浮床反应器， 反应压力 20MPa， 反应温度 455 °C , 液化催化 剂为镍 -钼或钴 -钼 /氧化铝载体的石油系加氢催化剂，采用旋液分离 和减压蒸馏分离循环溶剂。 该工艺由于采用强制循环悬浮床反应器和 石油系加氢催化剂， 反应温度控制容易， 产品性质稳定。 但由于石油 系加氢催化剂在煤液化反应体系中失活严重， 催化剂更新周期短， 液 化油生产成本高。
德国在 80年代末开发了 IGOR+工艺。 IGOR+工艺采用鼓泡床反 应器， 减压蒸馏分离循环溶剂， 并采用在线固定床加氢反应器对循环 溶剂和产品进行不同深度的加氢， 液化催化剂采用赤泥。 该工艺由于 全部采用加氢后的供氢性循环溶剂， 煤浆性质稳定， 煤浆浓度高， 预 热容易， 而且可以与高温分离器气相进行换热， 热利用率高。 但由于 赤泥催化剂活性低，反应条件苛刻，典型操作条件： 反应压力 30MPa, 反应温度 470°C ; 在线固定床加氢反应器存在催化剂结焦失活操作周 期短的风险； 鼓泡床反应器对高钙煤有矿物质沉积现象。
日本在 90年代末开发完成了 NEDOL工艺。 NEDOL工艺采用鼓 泡床反应器， 减压蒸馏分离循环溶剂， 并采用离线的固定床加氢反应 器对循环溶剂进行加氢， 液化催化剂采用超细粉碎的天然黄铁矿（0.7 μ )。 该工艺由于全部采用加氢后的供氢性循环溶剂， 煤浆性质稳定， 煤浆浓度高， 预热容易， 而且可以与高温分离器气相进行换热， 热利 用率高， 反应条件緩和， 典型的操作条件为反应压力 17MPa, 反应温 度 450°C。 但天然黄铁矿硬度大， 超细粉碎困难， 成本高； 鼓泡床反 应器气体滞留系数大， 利用率低； 鼓泡床反应器液速低， 有矿物质沉 积； 固定床加氢反应器存在操作周期短的风险。 发明内容
本发明的目的是提供一种煤炭直接液化方法， 能长期稳定运转， 反应器利用率高， 防止矿物质沉积， 反应条件緩和， 最大限度提高液 体收率， 并同时为液化产品进一步加工提供优质原料。
本发明的包括以下步骤：
( 1 )将原料煤制备成一种煤浆；
( 2 ) 将所述煤浆经过预处理后送入一个反应系统中进行液化反 应；
( 3 )将反应产物在分离器中进行气液分离， 其中的液相部分通过 一个蒸馏塔进行分离， 形成轻质油分和塔底物料；
( 4 ) 将所述塔底物料送入另一个蒸馏塔分离为馏出油和残渣； ( 5 )将所述的轻质油分和馏出油进行混合， 将混合产物送入一个 强制循环悬浮床加氢反应器进行催化加氢； ( 6 )将加氢产物通过一个分馏塔分离出产品油和其他供氢性循环 溶剂。
优选地， 制备煤浆的过程包括以下步骤： （a) 将原料煤经过一个 煤前处理装置干燥粉碎后， 制成一定粒度的煤粉； （b) 将煤粉与一种 催化剂原料在一个催化剂制备装置制成超细颗粒的催化剂； （c) 将煤 粉与所述催化剂在一个煤浆制备装置中与一种供氢性溶剂混合制成所 述煤浆。
根据本发明， 所述液化反应的步骤包括以下步骤： （a) 所述煤浆 与氢气混合后经过预热进入第一强制循环悬浮床反应器中进行反应； ( b )将第一强制循环悬浮床反应器的出口物料补充氢气后送入第二强 制循环悬浮床反应器中进行反应。 其中所述液化反应的条件为：
反应温度： 430-465°C;
反应压力： 15- 19MPa;
气液比： 600- 1000NL/kg;
煤浆空速： 0.7- 1.0t/m3 · h;
催化剂添加量： Fe/干煤 = 0.5-1.0 wt%。
所述气液分离的过程包括以下步骤： （a) 将反应物料送入一个高 温分离器中进行气液分离， 其中高温分离器的控制温度在 420°C; (b) 将高温分离器的气相部分送入一个低温分离器中进行进一步气液分 离， 其中低温分离器的温度为室温。
根据本发明的一个优选实施例， 所述的液化催化剂为直径 20-30 纳米、 长 100- 180纳米的 γ -水合氧化铁 ( γ -FeOOH), 同时含有 硫， 硫的添加比例为 S/Fe=2 (摩尔比）。
根据本发明， 所述加氢的反应条件为：
反应温度： 330- 390°C;
反应压力： 10- 15MPa;
气液比： 600- 1000NL/kg;
空速： 0.8- 2.5h_1。
所述的供氢性循环溶剂是煤炭直接液化油经过加氢得到的产物， 其馏程为 220 - 450°C。
所述的减压蒸镏塔的残渣含固体量为 50 - 55wt %。
所述的常压蒸馏塔塔顶油和减压蒸馏塔减压油混合后的产物的馏 程为 C5 - 530°C。
另外， 所述的强制循环悬浮床加氢反应器是带有内构件、 底部带 有循环泵、 可以进行催化剂置换的反应器。
本发明提供的一种煤炭直接液化方法， 采用高活性液化催化剂、 供氢性循环溶剂、强制循环悬浮床反应器、减压蒸馏分离沥青和固体， 强制循环悬浮床加氢反应器， 能长期稳定运转， 反应器利用率高， 防 止矿物盾沉积， 反应条件緩和， 最大限度地提高液体收率， 并同时为 液化产品进一步加工提供的优质原料。 附图说明
下面参照附图， 可以更容易理解本发明的技术方案。 附图中： 图 1是本发明的煤炭直接液化方法的一个流程图。 具体实施方式
图 中各个数字分别代表： 1、 原料煤， 2、 煤前处理装置， 3、 催化剂原料， 4、 催化剂制备装置， 5、 煤浆制备装置， 6、 氢气， 7、 第一强制循环悬浮床反应器， 8、 第二强制循环悬浮床反应器， 9、 高 温分离器， 10、 低温分离器， 11、 常压蒸馏塔， 12、 减压蒸馏塔， 13、 液化油强制循环悬浮床加氢反应器， 14、 气液分离器， 15、 产品分馏 塔， 16、 供氢性溶剂。
液化原料煤 1经过煤前处理装置 2干燥粉碎后制成一定粒度的煤 粉。 催化剂原料 3经过催化剂制备装置 4制成超细颗粒的催化剂。 煤 粉和催化剂在煤浆制备装置 5与供氢性溶剂 16混合制成煤浆。 煤浆 与氢气 6混合后经过预热进入第一强制循环悬浮床反应器 7, 第一强 制循环悬浮床反应器 7出口物料补充氢气后进入第二强制循环悬浮床 反应器 8。 第二强制循环悬浮床反应器 8反应物料进入高温分离器 9 进行气液分离， 高温分离器 9控制温度在 420°C。 高温分离器 9气相 部分进入低温分离器 10进一步气液分离，低温分离器 10温度为室温。 低温分离器 10气相部分与氢气混合循环使用， 废气部分被排出系统。 高温分离器 9和低温分离器 10的液相部分进入常压蒸馏塔 11分离出 轻质馏分， 常压蒸馏塔 11塔底物料进入减压蒸馏塔 12进行沥青和固 体的脱除， 减压蒸馏塔 12 塔底物料即为液化残渣， 为了保证残渣能 在一定温度下顺利排除， 一般控制残渣中的固体含量为 50 - 55wt %。 常压蒸馏塔 11和减压蒸馏塔 12的馏出油与氢气 6混合全部进入强制 循环悬浮床加氢反应器 13 进行以提高溶剂供氢性能为目的的催化加 氢。 由于液化油稠环芳烃和杂原子含量高， 組成复杂， 催化剂容易积 炭失活， 采用强制循环悬浮床反应器后， 催化剂可以定期更新， 操作 周期可以无限延长， 而且也避免了固定床反应由于催化剂积炭压差增 大的风险。 强制循环悬浮床加氢反应器 13 出口物料进入气液分离器 14进行气液分离， 气液分离器 14气相部分与氢气混合循环使用， 废 气部分被排出系统。 气液分离器 14液相物料进入产品分馏塔 15, 分 馏出产品和供氢性溶剂 16。 产品全部为汽油、 柴油馏分。
其中， 所述的煤粉为含水量 0.5-4.0wt%、 粒度 < 0.15mm的褐煤或 年轻烟煤。
其中， 采用的液化催化剂为超细 γ -水合氧化铁（ γ - FeOOH ), 其直径 20 - 30纳米、 长为 100 - 180纳米， 同时添加 ；琉的添加比 例为 S/Fe=2 (摩尔比）。 由于该催化剂活性高、 添加量少， Fe/干煤为 0.5-1.0wt%, 煤液化转化率高， 残渣中由于催化剂带出的液化油少， 增 加了蒸馏油产率。
其中， 采用的供氢性循环溶剂是煤炭直接液化油经过加氢得到的 产物， 其馏程为 220 - 450°C。 由于循环溶剂采用预加氢， 溶剂性质稳 定， 成浆性好， 可以制备成含固体浓度 45 - 55wt °/ 々高浓度煤浆， 而 且煤浆流动性好， 煤浆粘度小于 400CP ( 60°C ); 由于循环溶剂采用预 加氢， 溶剂供氢性能好， 加上高活性液化催化剂， 液化反应条件温和， 反应压力 17 - 19MPa, 反应温度 440 - 465 °C。 由于循环溶剂采用预加 氢， 溶剂具有供氢性能， 在煤浆预热和换热过程中， 能阻止煤热分解 过程中自由基碎片的缩合， 防止结焦， 延长操作周期， 提高热利用率。
其中， 采用的强制循环悬浮床反应器由于气体滞留系数低， 反应 器液相利用率高， 而且由于有强制循环泵， 液速高， 反应器内没有矿 物质沉积。 在本发明的优选实施例中， 采用两个强制循环的悬浮床反 应器。 由于强制循环悬浮床反应器内为全返混流，轴向温度分布均匀， 反应温度控制容易， 通过进料温度即可控制反应温度， 不需要采用反 应器侧线急冷氢控制， 产品性质稳定。 由于强制循环悬浮床反应器气 体滞留系数低， 反应器液相利用率高； 由于强制循环悬浮床反应器内 液速高， 反应器内没有矿物质沉积。
在本发明的优选实施例中， 采用减压蒸馏的方法进行沥青和固体 物的脱除。 减压蒸馏是一种成熟和有效的脱除沥青和固体的分离方 法， 减压蒸馏的馏出物不含沥青， 可为循环溶剂的加氢增加供氢性提 供合格原料， 减压蒸馏的残渣含固体 50 - 55wt % ; 由于使用高活性的 液化催化剂， 添加量少， 残渣中含油量少， 产品中柴油馏分多。
在本发明的优选实施例中， 循环溶剂和产品采用强制循环悬浮床 加氢反应器进行加氢。 由于强制循环悬浮床加氢反应器采用上流式， 催化剂可以定期更新， 加氢后的供氢性溶剂供氢性能好， 产品性质稳 定， 操作周期可以无限延长， 而且也避免了固定床反应由于催化剂积 炭压差增大的风险。
下面是使用本发明的优选实施例对一种年轻烟煤进行直接液化 的液化结果。
试验条件为：
反应器温度： 第一反应器 455 °C , 第二反应器 455 °C。
反应压力： 第一反应器 19.0MPa, 第二反应器 19.0MPa。
煤浆浓度： 45/55 (干煤 /溶剂， 质量比）。
催化剂添加量： 液化催化剂： 1.0wt%( /干煤)。
硫添加量： S/Fe=2 (摩尔比）。
气液比： 1000NL/kg 煤浆 循环氢浓度： H2 85vol%
表 1是一种年轻烟煤在本发明的一种煤直接液化工艺连续试验装 置上的液化结果 （表中数据以干燥无灰基煤为基准）。 表 2 为同一种 年轻烟煤在某直接液化工艺连续试验装置上的液化结果（表中数据以 千燥无灰基煤为基准）。 表 1 一种年轻烟煤在本发明的煤直接液化连续试验装置上的液化结 果
Figure imgf000009_0001
表 2 —种年轻烟煤在现有技术直接液化工艺连续试验装置上的液化 结果
Figure imgf000009_0002
对比表 1、 表 2可以得出， 本发明设备的转化率、 油收率高于现有 技术设备， 有机残渣减少， 液化效果好。

Claims

权 利 要 求
1. 一种煤炭直接液化的方法， 其特征在于, 包括以下步骤： ( 1 )将原料煤制备成一种煤浆；
( 2 ) 将所述煤浆经过预处理后送入一个反应系统中进行液化反 应；
(3) 将反应产物在分离器 （9， 10) 中进行气液分离， 其中的液 相部分通过一常压蒸馏塔（ 11 )进行分离， 形成轻质油分和塔底物料；
(4)将所述塔底物料送入一个减压蒸馏塔（ 12)分离为馏出油和 残渣；
( 5 )将所述的轻盾油分和馏出油进行混合， 将混合产物送入一个 强制循环悬浮床加氢反应器 （ 13) 进行催化加氢；
(6)将加氢产物通过一个分馏塔（15)分离出产品油和其他供氢 性循环溶剂。
2. 根据权利要求 1所述的方法， 其特征在于， 步骤 (1)包括以下步 骤：
(a) 将原料煤经过一个煤前处理装置 （2) 干燥粉碎后， 制成一 定粒度的煤粉；
(b)将煤粉与一种催化剂原料（3) 在一个催化剂制备装置 （4) 制成超细颗粒的催化剂；
(c) 将煤粉与所述催化剂在一个煤浆制备装置 （5) 中与一种供 氢性溶剂 （ 16) 混合制成所述煤浆。
3. 根据权利要求 1所述的方法， 其特征在于， 所述液化反应步骤 包括以下步骤：
(a) 所述煤浆与氢气 （6) 混合后经过预热进入第一强制循环悬 浮床反应器 （7) 中进行反应；
(b) 将第一强制循环悬浮床反应器 （7) 的出口物料补充氢气后 送入第二强制循环悬浮床反应器 （8) 中进行反应；
其中所述液化反应的条件为： 反应温度： 430- 465 °C;
反应压力： 15- 19MPa;
气液比： 600- 1000NL/kg;
煤浆空速： 0.7- 1.0t/m3 ' h;
催化剂添加量： Fe/干煤 = 0.5-1.0 wt%。
4. 根据权利要求 1 所述的方法， 其特征在于， 所述步驟 （3) 中 气液分离的过程包括以下步骤：
(a) 将反应物料送入一个高温分离器 （9) 中进行气液分离， 其 中高温分离器 （9) 的控制温度在 420 °C;
(b) 将高温分离器 （9) 的气相部分送入一个低温分离器 （ 10) 中进行进一步气液分离， 其中低温分离器 （ 10) 的温度为室温。
5. 根据权利要求 2所述的方法， 其特征在于, 所述的催化剂为直 径 20- 30纳米、 长 100- 180纳米的 γ -水合氧化铁 ( γ - FeOOH), 同时含有硫， 硫的添加比例为 S/Fe=2 (摩尔比）。
6. 根据权利要求 1所述的方法， 其特征在于， 所述加氢的反应条 件为：
反应温度： 330- 390°C;
反应压力： 10_ 15MPa;
气液比： 600 - 1000NL/kg;
空速： 0.8- 2,5h_1。
7. 根据权利要求 1所述的方法， 其特征在于， 所述的供氢性循环 溶剂是煤炭直接液化油经过加氢得到的产物， 其馏程为 220-450°C。
8.根据权利要求 1所述的方法，其特征在于所述的减压蒸馏塔（ 12 ) 的残渣含固体量为 50 - 55wt%。
9. 根据权利要求 1所述的方法， 其特征在于， 所述的常压蒸馏塔 塔顶油和减压蒸馏塔减压油混合后的产物的馏程为 C5 - 530°C。
10.根据权利要求 1所述的方法， 其特征在于， 所述的强制循环悬 浮床加氢反应器 （ 13) 是带有内构件、 底部带有循环泵、 可以进行催 化剂置换的反应器。
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AU2005266712B2 (en) 2009-08-13
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CN1587351A (zh) 2005-03-02
UA83585C2 (uk) 2008-07-25
US20090152171A1 (en) 2009-06-18
US20090283450A2 (en) 2009-11-19
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