WO2014065420A1 - オレフィン及び単環芳香族炭化水素の製造方法、並びにエチレン製造装置 - Google Patents
オレフィン及び単環芳香族炭化水素の製造方法、並びにエチレン製造装置 Download PDFInfo
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- WO2014065420A1 WO2014065420A1 PCT/JP2013/079042 JP2013079042W WO2014065420A1 WO 2014065420 A1 WO2014065420 A1 WO 2014065420A1 JP 2013079042 W JP2013079042 W JP 2013079042W WO 2014065420 A1 WO2014065420 A1 WO 2014065420A1
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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
- C10G11/00—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils
- C10G11/02—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils characterised by the catalyst used
- C10G11/04—Oxides
- C10G11/05—Crystalline alumino-silicates, e.g. molecular sieves
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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
- C10G35/00—Reforming naphtha
- C10G35/04—Catalytic reforming
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- B—PERFORMING OPERATIONS; TRANSPORTING
- B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
- B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
- B01J19/00—Chemical, physical or physico-chemical processes in general; Their relevant apparatus
- B01J19/24—Stationary reactors without moving elements inside
- B01J19/245—Stationary reactors without moving elements inside placed in series
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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
- C10G35/00—Reforming naphtha
- C10G35/04—Catalytic reforming
- C10G35/06—Catalytic reforming characterised by the catalyst used
- C10G35/095—Catalytic reforming characterised by the catalyst used containing crystalline alumino-silicates, e.g. molecular sieves
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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
- C10G63/00—Treatment of naphtha by at least one reforming process and at least one other conversion process
- C10G63/02—Treatment of naphtha by at least one reforming process and at least one other conversion process plural serial stages only
- C10G63/04—Treatment of naphtha by at least one reforming process and at least one other conversion process plural serial stages only including at least one cracking step
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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
- C10G69/00—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one other conversion process
- C10G69/02—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one other conversion process plural serial stages only
- C10G69/04—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one other conversion process plural serial stages only including at least one step of catalytic cracking in the absence of hydrogen
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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
- C10G69/00—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one other conversion process
- C10G69/02—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one other conversion process plural serial stages only
- C10G69/06—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one other conversion process plural serial stages only including at least one step of thermal cracking in the absence of hydrogen
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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
- C10G69/00—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one other conversion process
- C10G69/02—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one other conversion process plural serial stages only
- C10G69/08—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one other conversion process plural serial stages only including at least one step of reforming naphtha
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- B—PERFORMING OPERATIONS; TRANSPORTING
- B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
- B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
- B01J2219/00—Chemical, physical or physico-chemical processes in general; Their relevant apparatus
- B01J2219/24—Stationary reactors without moving elements inside
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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
- C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
- C10G2300/10—Feedstock materials
- C10G2300/1037—Hydrocarbon fractions
- C10G2300/1048—Middle distillates
- C10G2300/1051—Kerosene having a boiling range of about 180 - 230 °C
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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
- C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
- C10G2300/10—Feedstock materials
- C10G2300/1037—Hydrocarbon fractions
- C10G2300/1048—Middle distillates
- C10G2300/1055—Diesel having a boiling range of about 230 - 330 °C
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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
- C10G2400/00—Products obtained by processes covered by groups C10G9/00 - C10G69/14
- C10G2400/20—C2-C4 olefins
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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
- C10G2400/00—Products obtained by processes covered by groups C10G9/00 - C10G69/14
- C10G2400/22—Higher olefins
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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
- C10G2400/00—Products obtained by processes covered by groups C10G9/00 - C10G69/14
- C10G2400/30—Aromatics
Definitions
- the present invention relates to a method for producing olefins and monocyclic aromatic hydrocarbons, and an ethylene production apparatus, and more particularly, a method for producing olefins having 2 to 4 carbon atoms and monocyclic aromatic hydrocarbons having 6 to 8 carbon atoms, and ethylene production. Relates to the device.
- This application claims priority to Japanese Patent Application No. 2012-236132 filed in Japan on October 25, 2012, the contents of which are incorporated herein by reference.
- Oils containing polycyclic aromatic components such as light cycle oil (hereinafter referred to as “LCO”), which is a cracked light oil produced by fluid catalytic cracking (hereinafter referred to as “FCC”) equipment, have so far been mainly used. It was used as a fuel base for light oil and heavy oil.
- LCO light cycle oil
- FCC fluid catalytic cracking
- high-added monocyclic aromatic hydrocarbons having 6 to 8 carbon atoms for example, benzene, toluene, crude oil, etc.
- BTX a technique for efficiently producing xylene
- Patent Document 1 BTX can be efficiently produced from the pyrolytic heavy oil obtained from the ethylene production apparatus, but light olefin is not produced, so the light olefin which is the original purpose of the ethylene production apparatus As a result, it becomes impossible to meet the demand to increase the production efficiency.
- it is demanded to increase the production efficiency of light olefins in an ethylene production apparatus it is naturally required to increase production efficiency while suppressing an increase in cost.
- the present invention has been made in view of the above circumstances, and the object of the present invention is to enable production of light olefins from an ethylene production apparatus with higher production efficiency and while suppressing an increase in cost, and also with respect to BTX efficiently.
- An object of the present invention is to provide an olefin and monocyclic aromatic hydrocarbon production method and an ethylene production apparatus which can be produced.
- the method for producing olefins and monocyclic aromatic hydrocarbons of the present invention comprises a cracking furnace and hydrogen, ethylene, propylene, C4 fraction, monocyclic aromatic having 6 to 8 carbon atoms from the cracked product produced in the cracking furnace.
- the feedstock oil is brought into contact with and reacted with an olefin containing a crystalline aluminosilicate and a catalyst for producing a monocyclic aromatic hydrocarbon to produce a olefin having 2 to 4 carbon atoms and a monocyclic aromatic hydrocarbon having 6 to 8 carbon atoms.
- a cracking and reforming reaction step to obtain a product comprising: By treating a part or all of the product obtained in the cracking and reforming reaction step with a product recovery unit of the ethylene production apparatus, a part of or all of the product has 2 to 4 carbon atoms. And a product recovery step for recovering each of the olefin and the monocyclic aromatic hydrocarbon having 6 to 8 carbon atoms.
- the hydrogenation reaction process of partially hydrogenating a part or all of the said raw material oil before the said cracking reforming reaction process.
- a hydrogen partial pressure is 1 to 9 MPa
- a hydrogenation temperature is 150 to 400 ° C.
- a hydrogenation catalyst is used as hydrogenation conditions for hydrogenating the raw material oil.
- the inorganic carrier containing aluminum oxide is selected from 10 to 30% by mass of at least one metal selected from Group 6 metals of the periodic table based on the total catalyst mass, and from Group 8 to 10 metals of the periodic table It is preferable to use a catalyst obtained by supporting 1 to 7% by mass of at least one metal.
- the product recovery step a part of the product obtained in the cracking and reforming reaction step is processed by a product recovery device of the ethylene production apparatus, and the cracking and reforming is performed. It is preferable to have a recycling step for returning the heavy fraction having 9 or more carbon atoms in the product obtained in the reaction step to the cracking and reforming reaction step.
- the raw material oil is reacted in the presence of a saturated hydrocarbon having 1 to 3 carbon atoms.
- the cracking and reforming reaction step in the cracking and reforming reaction step, two or more fixed bed reactors are used, and the cracking and reforming reaction and the catalyst for producing the olefin and monocyclic aromatic hydrocarbon are periodically switched. It is preferable that the reproduction of the above is repeated alternately or sequentially.
- the crystalline aluminosilicate contained in the olefin and monocyclic aromatic hydrocarbon production catalyst used in the cracking and reforming reaction step is mainly medium pore zeolite and / or large pore zeolite. The component is preferably used.
- the olefin and monocyclic aromatic hydrocarbon production catalyst used in the cracking and reforming reaction step preferably contains phosphorus.
- the ethylene production apparatus of the present invention comprises a cracking furnace, A product recovery device for separating and recovering hydrogen, ethylene, propylene, C4 fraction and fractions containing monocyclic aromatic hydrocarbons having 6 to 8 carbon atoms from the decomposition products produced in the cracking furnace; A pyrolysis heavy oil obtained from the cracking furnace and having a distillation property of 90% by volume distillation temperature of 390 ° C. or less is used as a raw material oil, and the olefin and monocyclic aroma containing crystalline aluminosilicate with respect to this raw material oil.
- a cracking and reforming reaction apparatus for obtaining a product containing an olefin having 2 to 4 carbon atoms and a monocyclic aromatic hydrocarbon having 6 to 8 carbon atoms by contacting with and reacting with a catalyst for producing an aromatic hydrocarbon;
- Product supply means for supplying a part or all of the product obtained in the cracking and reforming reaction apparatus to the product recovery apparatus.
- the production apparatus preferably includes a hydrogenation reaction apparatus that partially hydrogenates a part or all of the raw material oil before the cracking and reforming reaction apparatus.
- the product supply means is configured to supply a part of the product obtained in the cracking reforming reaction apparatus to the product recovery apparatus, and the cracking reforming reaction apparatus It is preferable to have a recycling means for returning the heavy fraction having 9 or more carbon atoms in the product obtained in step 1 to the cracking reforming reaction apparatus.
- the cracking and reforming reaction apparatus includes two or more fixed bed reactors, and these are periodically switched while the cracking and reforming reaction and the olefin and monocyclic aromatic hydrocarbons are produced. It is preferable that the regeneration of the catalyst is alternately or sequentially repeated.
- light olefins can be produced with higher production efficiency and while suppressing an increase in cost, and BTX can also be efficiently produced. Can be manufactured.
- FIG. 1 is a view for explaining an embodiment of an ethylene production apparatus according to the present invention
- FIG. 2 shows a cracking and reforming process of the ethylene production apparatus shown in FIG. It is a figure for demonstrating.
- the part other than the cracking and reforming process shown in FIG. 2 may be a known ethylene production apparatus having a decomposition process and a separation and purification process.
- An ethylene production apparatus described in Patent Document 1 can be given. Therefore, the embodiment of the ethylene production apparatus according to the present invention includes an ethylene production apparatus in which the cracking and reforming process of the present invention is added to the existing ethylene production apparatus.
- the ethylene production apparatus is called a steam cracker or a steam cracking apparatus.
- a steam cracker As shown in FIG. 1, hydrogen, ethylene, propylene, C 4 And a product recovery device 2 for separating and recovering each fraction and a fraction containing a monocyclic aromatic hydrocarbon having 6 to 8 carbon atoms (BTX fraction: cracked gasoline).
- BTX fraction cracked gasoline
- the cracking furnace 1 thermally decomposes raw materials such as a naphtha fraction, a kerosene fraction, and a light oil fraction to produce hydrogen, ethylene, propylene, C4 fraction, and BTX fraction, and a heavier residue than the BTX fraction.
- Pyrolytic heavy oil is produced as oil (bottom oil). This pyrolytic heavy oil is sometimes called Heavy Aromatic Residue oil (HAR oil).
- the operating conditions of the cracking furnace 1 are not particularly limited and can be operated under general conditions. For example, a method of operating the raw material together with diluted water vapor at a thermal decomposition reaction temperature of 770 to 850 ° C. and a residence time (reaction time) of 0.1 to 0.5 seconds can be mentioned.
- the lower limit of the thermal decomposition reaction temperature is more preferably 775 ° C. or higher, and further preferably 780 ° C. or higher.
- the upper limit of the thermal decomposition reaction temperature is more preferably 845 ° C. or less, and further 840 ° C. or less. preferable.
- the steam / raw material (mass ratio) is preferably 0.2 to 0.9, more preferably 0.25 to 0.8, and still more preferably 0.3 to 0.7.
- the residence time (reaction time) of the raw material is more preferably 0.15 to 0.45 seconds, and further preferably 0.2 to 0.4 seconds.
- the product recovery device 2 includes a pyrolysis heavy oil separation step 3 and further supplies hydrogen, ethylene, propylene, C4 fraction, monocyclic aromatic hydrocarbon having 6 to 8 carbon atoms (BTX fraction: cracked gasoline). Each recovery unit for separating and recovering the contained fraction is provided.
- the pyrolysis heavy oil separation step 3 is a distillation column that separates the decomposition product obtained in the cracking furnace 1 into a component having a lower boiling point and a higher component before being subjected to the main distillation.
- the low boiling point component separated in the pyrolysis heavy oil separation step 3 is taken out as a gas and pressurized by the cracked gas compressor 4.
- the predetermined boiling point is such that the low-boiling component mainly includes products intended by the ethylene production apparatus, that is, hydrogen, ethylene, propylene, C4 fraction, and cracked gasoline (BTX fraction). Is set.
- the high boiling point component (bottom fraction) separated in the pyrolysis heavy oil separation step 3 becomes pyrolysis heavy oil, which may be further separated as necessary.
- pyrolysis heavy oil For example, gasoline fraction, light pyrolysis heavy oil, heavy pyrolysis heavy oil and the like can be separated and recovered by a distillation tower or the like.
- the gas (cracked gas) separated in the pyrolysis heavy oil separation process 3 and pressurized by the cracking gas compressor 4 is subjected to cleaning, etc., and then a component having a higher boiling point than hydrogen and hydrogen in the cryogenic separation process 5. Separated. Next, the component having a boiling point higher than that of hydrogen is supplied to the demethanizer 6 and methane is separated and recovered. Under such a configuration, a hydrogen recovery unit 7 and a methane recovery unit 8 are formed on the downstream side of the cryogenic separation step 5. The recovered hydrogen and methane are both used in the cracking and reforming process 21 described later.
- the high boiling point components separated in the demethanizer 6 are supplied to the deethanizer 9.
- the deethanizer 9 separates ethylene and ethane into components having higher boiling points.
- the ethylene and ethane separated by the deethanizer 9 are separated into ethylene and ethane by the ethylene rectifying tower 10 and recovered.
- an ethane recovery unit 11 and an ethylene recovery unit 12 are formed on the downstream side of the ethylene rectification column 10.
- the recovered ethylene becomes a main product manufactured by an ethylene manufacturing apparatus.
- the recovered ethane can be supplied to the cracking furnace 1 together with raw materials such as a naphtha fraction, a kerosene fraction, and a light oil fraction, and can be recycled.
- the high boiling point component separated in the deethanizer 9 is supplied to the depropanizer 13. Then, the depropanizer 13 separates propylene and propane into components having higher boiling points. Propylene and propane separated in the depropanizer 13 are separated and recovered by a propylene fractionator 14 by rectification. Under such a configuration, a propane recovery unit 15 and a propylene recovery unit 16 are formed on the downstream side of the propylene rectification column 14. The recovered propylene is also a main product produced with ethylene production equipment together with ethylene.
- the high boiling point component separated in the depropanizer 13 is supplied to the depentanizer 17.
- the depentanizer 17 separates the component having 5 or less carbon atoms and the component having a higher boiling point, that is, the component having 6 or more carbon atoms.
- the component having 5 or less carbon atoms separated by the depentane tower 17 is separated by the debutane tower 18 into a C4 fraction mainly composed of 4 carbon components and a fraction mainly composed of 5 carbon components. Each will be collected.
- the component having 4 carbon atoms separated by the debutane tower 18 can be further supplied to an extractive distillation apparatus or the like, and separated and recovered into butadiene, butane, isobutane and butylene, respectively. Under such a configuration, a butylene recovery unit (not shown) is formed on the downstream side of the debutane tower 18.
- the cracked gasoline (BTX fraction) collected by the cracked gasoline recovery unit 19 is supplied to a BTX purification device 20 that separates and recovers the cracked gasoline into benzene, toluene, and xylene. Here, they can be separated and recovered, and it is desirable to install them from the viewpoint of chemical production.
- the component having 9 or more carbon atoms (C9 +) contained in the cracked gasoline is separated from the BTX fraction by the BTX purification device 20 and recovered.
- An apparatus for separation can be provided in the BTX purification apparatus 20.
- the component having 9 or more carbon atoms can be used as a raw material oil for the production of olefin and BTX, which will be described later, in the same manner as the pyrolysis heavy oil separated in the pyrolysis heavy oil separation step 3.
- FIG. 1 and FIG. 2 an embodiment of an ethylene production apparatus according to the present invention and a hydrocarbon production method using the ethylene production apparatus, that is, a carbon number of 2 to 4 according to the present invention.
- a method for producing olefins and monocyclic aromatic hydrocarbons having 6 to 8 carbon atoms will be described.
- the ethylene production apparatus is separated in the pyrolysis heavy oil separation step 3 as shown in FIG. 1 and is heavier than the recovered pyrolysis heavy oil (HAR oil), that is, the BTX fraction.
- HAR oil recovered pyrolysis heavy oil
- olefins and BTX fractions are mainly produced using hydrocarbons having 9 or more carbon atoms (aromatic hydrocarbons) as feedstock.
- the remaining heavy oil recovered from the cracked gasoline recovery unit 19 from the BTX fraction can also be used as a raw material.
- the apparatus configuration shown in FIG. 2 is provided.
- the apparatus configuration shown in FIG. 2 is for producing olefins having 2 to 4 carbon atoms and monocyclic aromatic hydrocarbons (BTX fraction) having 6 to 8 carbon atoms, and is obtained from the above-mentioned ethylene production apparatus.
- the olefin and BTX fraction are produced using pyrolysis heavy oil as a raw oil.
- the properties obtained by the distillation test vary greatly depending on the decomposition temperature and the decomposition raw material, but the 10 vol% distillation temperature (T10) is preferably 145 ° C. or higher and 230 ° C. or lower.
- the 90 vol% distillation temperature (T90) and the end point are not limited because they vary greatly depending on the fraction used, but if the fraction is obtained directly from the pyrolysis heavy oil separation step 3, for example 90 vol%
- the distillation temperature (T90) is preferably 400 ° C or higher and 600 ° C or lower
- the end point (EP) is preferably 450 ° C or higher and 800 ° C or lower.
- the density at 15 ° C. is 1.03 g / cm 3 or more and 1.08 g / cm 3 or less
- the kinematic viscosity at 50 ° C. is 20 mm 2 / s or more and 45 mm 2 / s or less
- the sulfur content (sulfur content) is 200 mass ppm. It is preferable that the content is 700 mass ppm or less
- the nitrogen content (nitrogen content) is 20 mass ppm or less
- the aromatic content is 80 volume% or more.
- the distillation test is a value measured in accordance with “Petroleum product-distillation test method” defined in JIS K 2254, and the density at 15 ° C. is the “crude oil and petroleum product defined in JIS K 2249” -Kinematic viscosity at 50 ° C measured according to "Density test method and density / mass / capacity conversion table (extract)" is JIS K 2283 "Crude oil and petroleum products” -Values obtained according to the "Kinematic Viscosity Test Method and Viscosity Index Calculation Method” and the sulfur content is defined as “Radiation Excitation” in "Crude Oil and Petroleum Products-Sulfur Content Test Method” defined in JIS K 2541-1992.
- the sulfur content measured in accordance with the “Method” and the nitrogen content are the nitrogen content measured in accordance with JIS K 2609 “Crude Oil and Petroleum Products—Nitrogen Content Testing Method” It refers Petroleum Institute method JPI-5S-49-97 the content of total aromatic content measured in "Petroleum products - - hydrocarbon type test method high performance liquid chromatography” means respectively.
- the pyrolyzed heavy oil is not directly used as a raw material oil, but the pyrolyzed heavy oil is preliminarily cut at a predetermined cut temperature (90% by volume distillation temperature is 90% by distillation). 390 ° C.) and separated into a light fraction (light pyrolysis heavy oil) and a heavy fraction (heavy pyrolysis heavy oil). And let the light fraction as shown below be feedstock.
- the heavy fraction is stored separately and used, for example, as fuel.
- the feedstock oil according to the present invention is a pyrolytic heavy oil obtained from the above-described ethylene production apparatus, and has a distillation property of 90 vol% distillation temperature of 390 ° C or lower. That is, a light pyrolysis heavy oil that has been distilled in the front distillation column 30 and has a distillation property of 90 vol% distillation temperature adjusted to 390 ° C. or lower is used as the raw material oil. Thus, by setting the 90% by volume distillation temperature to 390 ° C.
- the feedstock mainly consists of aromatic hydrocarbons having 9 to 12 carbon atoms, and a catalyst for producing olefins and monocyclic aromatic hydrocarbons, which will be described later,
- the yield of olefin and BTX fraction can be increased.
- the 10 vol% distillation temperature (T10) is preferably 140 ° C or higher and 220 ° C or lower
- the 90 vol% distillation temperature (T90) is 220 ° C or higher and 380 ° C. Or less, more preferably T10 is 160 ° C. or higher and 200 ° C.
- T90 is 240 ° C. or higher and 350 ° C. or lower.
- T90 90 volume% distillation temperature
- the distillation property is measured in accordance with “Petroleum product-distillation test method” defined in JIS K 2254.
- the raw material oil which concerns on this invention contains the pyrolysis heavy oil obtained from an ethylene manufacturing apparatus, it may contain another base material.
- the number of carbons separated and recovered by the cracked gasoline recovery unit 19 as described above, in addition to the light pyrolysis heavy oil obtained by distillation treatment in the front distillation column 30 Nine or more components (aromatic hydrocarbons) are also used.
- the fraction whose distillation property 90 volume% distillation temperature (T90) is adjusted to 390 ° C. or less does not necessarily need to be subjected to distillation treatment in the front distillation column 30. Therefore, as will be described later, separately from the pyrolysis heavy oil shown in FIG. 2, a hydrogenation reaction device 31 or a cracking reforming reaction device which is a device constituting the cracking reforming process 21 on the rear stage side of the front distillation column 30. It is also possible to supply to 33 directly.
- Part or all of the raw material oil obtained in this way is partially hydrogenated by the hydrogenation reactor 31. That is, part or all of the feedstock is subjected to the hydrogenation reaction step.
- only the light pyrolysis heavy oil that is, only a part of the raw material oil is subjected to partial hydrogenation treatment.
- Hydrocarbons with 9 carbon atoms of some fractions when pyrolyzed heavy oil is separated into a plurality of fractions or residual oils when other chemicals or fuels are produced from these separated fractions.
- Hydrogenation treatment can be omitted for main components and components having 9 or more carbon atoms separated and recovered by the cracked gasoline recovery unit 19. However, it goes without saying that these components may also be partially hydrogenated by the hydrogenation reactor 31.
- Pyrolytic heavy oil obtained from an ethylene production apparatus usually has a very high content of aromatic hydrocarbons. Therefore, in the present embodiment, a necessary fraction in the pyrolyzed heavy oil separated earlier, that is, light HAR is used as a raw material oil, and this raw material oil is hydrogenated in the hydrogenation reaction device 31 (hydrogenation reaction step). To process. However, a large amount of hydrogen is required for hydrotreating the feedstock until it is hydrocracked, and when a completely hydrogenated feedstock is used, a catalyst for producing olefins and monocyclic aromatic hydrocarbons described later is used. Thus, the production efficiency of the olefin and BTX fractions in the cracking and reforming reaction step by contacting and reacting with each other becomes extremely low.
- the raw oil is not completely hydrogenated but only partially hydrogenated. That is, mainly the bicyclic aromatic hydrocarbons in the feedstock oil are selectively hydrogenated and converted to monocyclic aromatic hydrocarbons (such as naphthenobenzenes) in which only one aromatic ring is hydrogenated.
- monocyclic aromatic hydrocarbons such as naphthenobenzenes
- examples of the monocyclic aromatic hydrocarbon include indane, tetralin, alkylbenzene, and the like.
- the hydrogenation treatment is partially performed in this manner, the amount of heat generated during the treatment can be reduced while simultaneously reducing the amount of hydrogen consumed in the hydrogenation reaction step.
- naphthalene which is a typical example of a bicyclic aromatic hydrocarbon
- the hydrogen consumption per mole of naphthalene is 5 moles, but when hydrogenating to tetralin, the hydrogen consumption is Can be realized at 2 moles.
- feedstock oil pyrolytic heavy oil
- the hydrogen consumption required to hydrogenate this fraction to indanes is that hydrogen is converted from naphthalene to decalin. Even less than the amount needed to convert. Therefore, it becomes possible to more efficiently convert the bicyclic aromatic hydrocarbons in the feedstock oil to naphthenobenzenes.
- the hydrogen recovered by the hydrogen recovery unit 7 can be used as the hydrogen used in this hydrogenation reaction step. That is, the hydrogen recovered by the hydrogen recovery unit 7 is supplied to the hydrogenation reaction device 31 to perform a hydrogenation process. Therefore, by using hydrogen generated by the same ethylene production apparatus, the space and cost required for hydrogen storage and movement can be minimized.
- the hydrogenation reaction apparatus 31 for performing such a hydrogenation treatment a known hydrogenation reactor can be used.
- the hydrogen partial pressure at the reactor inlet is preferably 1 to 9 MPa.
- the lower limit is more preferably 1.2 MPa or more, and further preferably 1.5 MPa or more.
- 7 MPa or less is more preferable, and 5 MPa or less is further more preferable.
- the LHSV (Liquid Hourly Space Velocity) of the hydrogenation reaction step by the hydrogenation reactor 31 is preferably 0.05 to 10 h ⁇ 1 . More preferably at least 0.1 h -1 as the lower limit, 0.2 h -1 or more is more preferable. Further, more preferably 5h -1 or less as the upper limit, 3h -1 or less is more preferable.
- LHSV is less than 0.05 h ⁇ 1 , there is a concern that the construction cost of the reactor becomes excessive and the economic efficiency is impaired.
- the LHSV exceeds 10 h ⁇ 1 the hydrotreatment of the feedstock does not proceed sufficiently, and the target hydride may not be obtained.
- the reaction temperature (hydrogenation temperature) in the hydrogenation reaction step by the hydrogenation reactor 31 is preferably 150 ° C. to 400 ° C. As a minimum, 170 degreeC or more is more preferable, and 190 degreeC or more is further more preferable. Moreover, as an upper limit, 380 degrees C or less is more preferable, and 370 degrees C or less is further more preferable.
- the reaction temperature is lower than 150 ° C., the hydrogenation treatment of the raw material oil tends not to be sufficiently achieved.
- the reaction temperature exceeds 400 ° C., the generation of gas as a by-product increases, so the yield of the hydrotreated oil decreases, which is not desirable.
- the hydrogen / oil ratio in the hydrogenation reaction step by the hydrogenation reactor 31 is preferably 100 to 2000 NL / L.
- 110 NL / L or more is more preferable, and 120 NL / L or more is further more preferable.
- 1800 NL / L or less is more preferable, and 1500 NL / L or less is further more preferable.
- the hydrogen / oil ratio is less than 100 NL / L, coke formation on the catalyst proceeds at the reactor outlet, and the catalyst life tends to be shortened.
- the hydrogen / oil ratio exceeds 2000 NL / L, there is a concern that the construction cost of the recycle compressor becomes excessive and the economic efficiency is impaired.
- the reaction mode in the hydrogenation treatment of the hydrogenation reaction apparatus 31 is not particularly limited, it can usually be selected from various processes such as a fixed bed and a moving bed, and among them, a fixed bed is preferable. Moreover, it is preferable that the hydrogenation reaction apparatus 31 is columnar.
- the hydrotreating catalyst accommodated in the hydrogenation reaction apparatus 31 and used for the hydrotreating of the feedstock selectively hydrogenates the bicyclic aromatic hydrocarbons in the feedstock and only has one aromatic ring.
- the catalyst is not limited as long as it is a catalyst that can be converted into a hydrogenated monocyclic aromatic hydrocarbon (such as naphthenobenzenes).
- a preferred hydrotreating catalyst contains at least one metal selected from Group 6 metals of the periodic table and at least one metal selected from Group 8 to 10 metals of the periodic table.
- the Group 6 metal of the periodic table molybdenum, tungsten, and chromium are preferable, and molybdenum and tungsten are particularly preferable.
- the Group 8-10 metal of the periodic table iron, cobalt, and nickel are preferable, and cobalt and nickel are more preferable. These metals may be used alone or in combination of two or more. As specific examples of metal combinations, molybdenum-cobalt, molybdenum-nickel, tungsten-nickel, molybdenum-cobalt-nickel, tungsten-cobalt-nickel, and the like are preferably used.
- the periodic table is a long-period type periodic table defined by the International Union of Pure and Applied Chemistry (IUPAC).
- the hydrotreating catalyst is preferably one in which the metal is supported on an inorganic carrier containing aluminum oxide.
- the inorganic carrier containing the aluminum oxide include alumina, alumina-silica, alumina-boria, alumina-titania, alumina-zirconia, alumina-magnesia, alumina-silica-zirconia, alumina-silica-titania, and various types. Examples include a carrier in which a porous inorganic compound such as various clay minerals such as zeolite, ceviolite, and montmorillonite is added to alumina, among which alumina is particularly preferable.
- the inorganic carrier composed of a plurality of metal oxides such as alumina-silica may be a simple mixture of these oxides or a complex oxide.
- the catalyst for hydrotreating is an inorganic carrier containing aluminum oxide and at least one selected from Group 6 metals of the periodic table on the basis of the total catalyst mass, which is the total mass of the inorganic carrier and the metal.
- a catalyst obtained by supporting 10 to 30% by mass of metal and 1 to 7% by mass of at least one metal selected from Group 8 to 10 metals of the periodic table is preferable.
- the precursor of the metal species used when the metal is supported on the inorganic carrier is not particularly limited, but an inorganic salt of the metal, an organometallic compound, or the like is used, and a water-soluble inorganic salt is preferably used.
- the In the loading step loading is performed using a solution of these metal precursors, preferably an aqueous solution.
- a known method such as an immersion method, an impregnation method, a coprecipitation method, or the like is preferably employed.
- the carrier on which the metal precursor is supported is preferably dried and then calcined in the presence of oxygen, and the metal species is once converted to an oxide. Furthermore, it is preferable to convert the metal species into a sulfide by a sulfidation treatment called pre-sulfidation before performing the hydrogenation treatment of the raw material oil.
- the conditions for the preliminary sulfidation are not particularly limited, but a sulfur compound is added to a petroleum fraction or pyrolysis heavy oil (hereinafter referred to as a preliminary sulfidation feedstock oil), and the temperature is 200 to 380 ° C., and the LHSV is 1 to It is preferable that the catalyst is continuously brought into contact with the hydrotreating catalyst under the conditions of 2h ⁇ 1 , the pressure being the same as that in the hydrotreating operation, and the treating time being 48 hours or longer.
- the sulfur compound added to the pre-sulfided raw material oil is not limited, but dimethyl disulfide (DMDS), sulfazole, hydrogen sulfide and the like are preferable. It is preferable to add about mass%.
- the hydrotreated oil of the raw material oil obtained from the hydrogenation reaction apparatus 31 (hydrogenation reaction step) described above has the following properties.
- the distillation properties are such that 10% by volume distillation temperature (T10) is 140 ° C. or higher and 200 ° C. or lower, 90% by volume distillation temperature (T90) is 200 ° C. or higher and 390 ° C. or lower, more preferably T10 is 160 ° C. or higher and 190 ° C. or lower, T90 is 210 ° C. or higher and 370 ° C. or lower.
- T10 is less than 140 ° C.
- the raw material oil formed by including this hydrotreated oil may contain xylene, which is one of the target products.
- T90 exceeds 390 ° C. (becomes heavy)
- catalyst performance deteriorates due to metal poisoning, coke deposition, etc. on the hydrotreating catalyst, and for the production of olefins and monocyclic aromatic hydrocarbons described later. This is not preferable from the viewpoints that coke deposition on the catalyst increases and the predetermined performance cannot be obtained, and that hydrogen consumption increases and is not economical.
- the hydrotreated oil of this raw material oil is supplied to the cracking and reforming reaction apparatus 33 after the hydrogen is removed in the subsequent dehydrogenation tower 32 and is supplied to the cracking and reforming reaction step.
- the cracking reforming reaction apparatus 33 directly receives a fraction mainly composed of hydrocarbons having about 9 to 10 carbon atoms, which does not contain many polycyclic aromatics together with the hydrotreated oil and has a low necessity for hydrogenation. It can also be supplied.
- a heating furnace (not shown) is provided between the dehydrogenation tower 32 and the cracking reforming reaction apparatus 33, and the hydrotreated oil of the raw material oil and the C9 fraction are subjected to a predetermined temperature as a pretreatment. To be heated. That is, when the cracking and reforming reaction apparatus 33 is brought into contact with the catalyst, it is preferable that these raw material oils and the like are in a gas phase state. Keep it close to this. Moreover, the hydrogen removed and recovered by the dehydrogenation tower 32 is returned to the hydrogenation reaction device 31 again and can be used for the hydrogenation treatment, or the hydrogen can be recovered again by the ethylene production device.
- the cracking and reforming reaction apparatus 33 contains an olefin and a monocyclic aromatic hydrocarbon production catalyst.
- the raw material oil (including hydrotreated oil) supplied to the catalyst is brought into contact with and reacted with carbon.
- a product containing an olefin having 2 to 4 carbon atoms and a monocyclic aromatic hydrocarbon having 6 to 8 carbon atoms is obtained.
- the catalyst for producing olefins and monocyclic aromatic hydrocarbons contains crystalline aluminosilicate.
- the content of the crystalline aluminosilicate in the catalyst may be determined according to the required reactivity and selectivity of the cracking reforming reaction or the shape and strength of the catalyst, and is not particularly limited, but is 10 to 100% by mass. preferable. Since it is used in a fixed bed reactor, it may be a catalyst consisting only of crystalline aluminosilicate. If a binder is added to increase the strength, the content of crystalline aluminosilicate is preferably 20 to 95% by mass, more preferably 25 to 90% by mass. However, if the content of crystalline aluminosilicate is less than 10%, the amount of catalyst for obtaining sufficient catalytic activity becomes excessive, which is not preferable.
- the crystalline aluminosilicate is preferably mainly composed of medium pore zeolite and / or large pore zeolite because the yield of monocyclic aromatic hydrocarbons can be further increased.
- the medium pore zeolite is a zeolite having a 10-membered ring skeleton structure. Examples of the medium pore zeolite include AEL type, EUO type, FER type, HEU type, MEL type, MFI type, NES type, and TON type. And zeolite having a WEI type crystal structure. Among these, the MFI type is preferable because the yield of monocyclic aromatic hydrocarbons can be further increased.
- the large pore zeolite is a zeolite having a 12-membered ring skeleton structure.
- Examples of the large pore zeolite include AFI type, ATO type, BEA type, CON type, FAU type, GME type, LTL type, and MOR type. , Zeolites of MTW type and OFF type crystal structures.
- the BEA type, FAU type, and MOR type are preferable in terms of industrial use, and the BEA type is preferable because the yield of monocyclic aromatic hydrocarbons can be further increased.
- the crystalline aluminosilicate may contain, in addition to the medium pore zeolite and the large pore zeolite, a small pore zeolite having a skeleton structure having a 10-membered ring or less, and a very large pore zeolite having a skeleton structure having a 14-membered ring or more.
- examples of the small pore zeolite include zeolites having crystal structures of ANA type, CHA type, ERI type, GIS type, KFI type, LTA type, NAT type, PAU type, and YUG type.
- Examples of the ultra-large pore zeolite include zeolites having CLO type and VPI type crystal structures.
- the crystalline aluminosilicate has a molar ratio of silicon to aluminum (Si / Al ratio) of 100 or less, preferably 50 or less.
- Si / Al ratio of the crystalline aluminosilicate exceeds 100, the yield of monocyclic aromatic hydrocarbons becomes low.
- the Si / Al ratio of the crystalline aluminosilicate is preferably 10 or more in order to obtain a sufficient yield of monocyclic aromatic hydrocarbons.
- the olefin and monocyclic aromatic hydrocarbon production catalyst according to the present invention may further contain gallium and / or zinc. By including gallium and / or zinc, more efficient BTX production can be expected.
- gallium and / or zinc As crystalline aluminosilicate containing gallium and / or zinc, gallium is incorporated in the lattice skeleton of crystalline aluminosilicate (crystalline aluminogallosilicate), or zinc is incorporated in the lattice skeleton of crystalline aluminosilicate.
- the Ga-supported crystalline aluminosilicate and / or the Zn-supported crystalline aluminosilicate is a material in which gallium and / or zinc is supported on the crystalline aluminosilicate by a known method such as an ion exchange method or an impregnation method.
- the gallium source and zinc source used at this time are not particularly limited, and examples thereof include gallium salts such as gallium nitrate and gallium chloride, zinc salts such as gallium oxide, zinc nitrate and zinc chloride, and zinc oxide.
- the upper limit of the content of gallium and / or zinc in the catalyst is preferably 5% by mass or less, more preferably 3% by mass or less, and more preferably 2% by mass or less when the total amount of the catalyst is 100% by mass. More preferably, it is more preferably 1% by mass or less. If the content of gallium and / or zinc exceeds 5% by mass, the yield of monocyclic aromatic hydrocarbons is lowered, which is not preferable. Further, the lower limit of the content of gallium and / or zinc is preferably 0.01% by mass or more, and more preferably 0.1% by mass or more, when the total amount of the catalyst is 100% by mass. If the gallium and / or zinc content is less than 0.01% by mass, the yield of monocyclic aromatic hydrocarbons may be low, which is not preferable.
- Crystalline aluminogallosilicate and / or crystalline aluminodine silicate is a structure in which the SiO 4 , AlO 4 and GaO 4 / ZnO 4 structures are tetrahedrally coordinated in the skeleton, and gel crystallization by hydrothermal synthesis, It can be obtained by inserting gallium and / or zinc into the lattice skeleton of the crystalline aluminosilicate, or inserting aluminum into the lattice skeleton of the crystalline gallosilicate and / or crystalline zinc silicate.
- the catalyst for producing olefins and monocyclic aromatic hydrocarbons preferably contains phosphorus.
- the phosphorus content in the catalyst is preferably 0.1 to 10.0% by mass when the total amount of the catalyst is 100% by mass.
- the lower limit of the phosphorus content is preferably 0.1% by mass or more, and more preferably 0.2% by mass or more because it can prevent a decrease in yield of monocyclic aromatic hydrocarbons over time.
- the upper limit of the phosphorus content is preferably 10.0% by mass or less, more preferably 6.0% by mass or less, and more preferably 3.0% by mass or less because the yield of monocyclic aromatic hydrocarbons can be increased. Further preferred.
- the method for incorporating phosphorus into the catalyst for producing olefins and monocyclic aromatic hydrocarbons is not particularly limited.
- Examples include a method of supporting phosphorus on a silicate, a method of containing a phosphorus compound during zeolite synthesis and replacing part of the crystalline aluminosilicate skeleton with phosphorus, a method of using a crystal accelerator containing phosphorus during zeolite synthesis, and the like. It is done.
- the aqueous solution containing phosphate ions used at that time is not particularly limited, but phosphoric acid, diammonium hydrogen phosphate, ammonium dihydrogen phosphate, and other water-soluble phosphates can be dissolved in water at an arbitrary concentration. What was prepared can be used preferably.
- Such a catalyst for producing olefins and monocyclic aromatic hydrocarbons includes a crystalline aluminogallosilicate / crystalline aluminodine silicate carrying phosphorus as described above, or a crystalline aluminosilicate carrying gallium / zinc and phosphorus.
- the catalyst for producing olefins and monocyclic aromatic hydrocarbons is formed into a powder form, a granular form, a pellet form or the like according to the reaction mode of the cracking reforming reaction apparatus 33 (cracking reforming reaction step).
- a cracking reforming reaction step For example, in the case of a fixed bed, it is formed in a granular or pellet form, and in the case of a fluidized bed, it is formed in a powder form.
- an inert oxide may be blended into the catalyst as a binder and then molded using various molding machines.
- the binder is preferably an inorganic substance such as silica or alumina.
- the olefin and monocyclic aromatic hydrocarbon production catalyst contains a binder or the like
- a material containing phosphorus as a binder may be used as long as the preferable range of the phosphorus content described above is satisfied.
- the binder and the crystalline aluminosilicate supported on gallium and / or zinc are mixed, or the binder and the crystalline aluminogallosilicate and / or crystal are mixed. After mixing with the basic aluminodine silicate, phosphorus may be added to produce the catalyst.
- reaction format As the reaction type of the cracking and reforming reaction apparatus 33, that is, the reaction form when the raw material oil is brought into contact with the olefin and the catalyst for producing monocyclic aromatic hydrocarbons by the cracking and reforming reaction apparatus 33 and subjected to the cracking and reforming reaction, Examples include a bed, a moving bed, and a fluidized bed.
- the fixed bed is preferable because the cost of the apparatus is much lower than that of the fluidized bed or moving bed. Therefore, it is possible to repeat the reaction and regeneration in one fixed bed reactor, but it is preferable to install two or more reactors in order to carry out the reaction continuously.
- a fixed bed cracking reforming reaction apparatus 33 (fixed bed reactor) is used, and two fixed bed reactors 33 are used. In FIG. 2, two fixed bed reactors 33 are shown. However, the number is not limited to this, and any number of fixed bed reactors 33 can be installed as long as there are two or more.
- the activity of the catalyst is reduced due to the adhesion of coke. Therefore, the catalyst is regenerated after being operated for a predetermined time. That is, using two or more cracking reforming reaction apparatuses 33 (fixed bed reactors), the cracking reforming reaction and regeneration of the catalyst for producing olefins and monocyclic aromatic hydrocarbons are repeated while periodically switching them.
- the operating time for continuous operation with one decomposition reforming reaction apparatus 33 is several hours to 10 days, although it varies depending on the size of the apparatus and various operating conditions (reaction conditions). If the number of reactors of the cracking reforming reactor 33 (fixed bed reactor) is increased, the continuous operation time per reactor can be shortened, and the decrease in the activity of the catalyst can be suppressed, so that regeneration is required. Time can be shortened.
- reaction temperature The reaction temperature at the time of contacting and reacting the feedstock with the catalyst is not particularly limited, but is preferably 350 to 700 ° C, more preferably 400 to 650 ° C. When the reaction temperature is less than 350 ° C., the reaction activity is not sufficient. When the reaction temperature exceeds 700 ° C., it is disadvantageous in terms of energy, and at the same time, the production of coke is remarkably increased and the production efficiency of the target product is lowered.
- reaction pressure The reaction pressure when contacting and reacting the raw material oil with the catalyst is 0.1 MPaG to 2.0 MPaG. That is, the contact between the raw material oil and the olefin and monocyclic aromatic hydrocarbon production catalyst is performed under a pressure of 0.1 MPaG to 2.0 MPaG. Since the present invention has a completely different reaction concept from the conventional method by hydrocracking, it does not require any high-pressure conditions that are advantageous in hydrocracking. Rather, an unnecessarily high pressure is not preferable because it promotes decomposition and by-produces a light gas that is not intended. In addition, the fact that the high pressure condition is not required is advantageous in designing the reactor. Therefore, when the reaction pressure is 0.1 MPaG to 2.0 MPaG, the cracking and reforming reaction can be performed efficiently.
- the contact time between the feedstock and the catalyst is not particularly limited as long as the desired reaction proceeds substantially.
- the gas passage time on the catalyst is preferably 2 to 150 seconds, more preferably 3 to 100 seconds. More preferably, it is ⁇ 80 seconds. If the contact time is less than 5 seconds, substantial reaction is difficult. If the contact time exceeds 300 seconds, the accumulation of carbonaceous matter in the catalyst due to coking or the like will increase, or the amount of light gas generated due to decomposition will increase.
- the regeneration treatment is performed by removing coke from the catalyst surface. Specifically, air is passed through the cracking and reforming reaction device 33 to burn the coke adhering to the catalyst surface. Since the cracking and reforming reaction apparatus 33 is maintained at a sufficiently high temperature, the coke adhering to the catalyst surface is easily burned by simply circulating air. However, if normal air is supplied to the cracking reforming reaction apparatus 33 and distributed, rapid combustion may occur. Therefore, it is preferable to supply the air whose oxygen concentration has been lowered by mixing nitrogen in advance to the cracking reforming reaction apparatus 33 and to distribute it. That is, as the air used for the regeneration treatment, it is preferable to use, for example, an oxygen concentration reduced to about several to 10%. Further, the reaction temperature and the regeneration temperature are not necessarily the same, and a preferable temperature can be appropriately set.
- methane acts as a diluent that lowers the concentration of the hydrocarbons on the catalyst surface to prevent the heavy hydrocarbons derived from the feedstock from adhering to the catalyst surface and the catalytic reaction from proceeding ( to disturb. Therefore, methane comes to suppress that the heavy hydrocarbon derived from raw material oil adheres to the catalyst surface, and becomes coke.
- the methane recovered by the methane recovery unit 8 is used as the methane to be supplied to the cracking / reforming reaction apparatus 33. That is, the methane recovered by the methane recovery unit 8 is supplied to the cracking / reforming reactor 33 as a diluent.
- the methane supplied to the cracking and reforming reaction apparatus 33 is heat-treated at a predetermined temperature in a heating furnace (not shown) provided upstream of the cracking and reforming reaction apparatus 33 together with the raw material oil.
- ethane and propane can also be used instead of methane, it is more preferable to use methane having the lowest reactivity and recovering a sufficient amount in the same ethylene production apparatus.
- the methane / oil ratio in the cracking / reforming reaction step by the cracking / reforming reactor 33 is preferably 20 to 2000 NL / L. As a minimum, 30 NL / L or more is more preferable, and 50 NL / L or more is further more preferable. Moreover, as an upper limit, 1800 NL / L or less is more preferable, and 1500 NL / L or less is further more preferable.
- the methane / oil ratio is less than 20 NL / L, the dilution effect is not sufficient, and the adhesion of coke to the catalyst surface cannot be sufficiently suppressed.
- a fluidized bed that can continuously remove coke adhering to the catalyst and can carry out the reaction stably. It can also be used. In that case, it is more preferable to use a continuously regenerating fluidized bed in which the catalyst circulates between the reactor and the regenerator and the reaction-regeneration is continuously repeated.
- a fluidized bed reactor has a higher apparatus cost than the fixed bed reactor, it is preferable to use the above fixed bed reactor in order to suppress the cost increase of the entire ethylene production apparatus.
- the cracking reforming reaction product derived from the cracking reforming reaction apparatus 33 includes a gas containing an olefin having 2 to 4 carbon atoms, a BTX fraction, and an aromatic hydrocarbon having C9 or more. Therefore, the cracking / reforming reaction product is separated into each component by the purification / recovery device 34 provided at the subsequent stage of the cracking / reforming reaction device 33, and purified and recovered.
- the purification and recovery device 34 includes a BTX fraction recovery tower 35 and a gas separation tower 36.
- the BTX fraction tower 35 distills the cracking and reforming reaction product and separates it into a light fraction having 8 or less carbon atoms and a heavy fraction having 9 or more carbon atoms.
- the gas separation tower 36 distills the light fraction having 8 or less carbon atoms separated by the BTX fraction collection tower 35, and a BTX fraction containing benzene, toluene and crude xylene, and a gas fraction having a lower boiling point than these. To separate. In these BTX fraction collection tower 35 and gas separation tower 36, since the fraction obtained in each is reprocessed as will be described later, it is not necessary to increase the distillation accuracy, and the distillation operation is carried out relatively roughly. Can do.
- the gas fraction separated in the gas separation tower 36 mainly includes C4 such as hydrogen, ethylene, propylene, butylene and the like. Distillate, BTX is included. Therefore, these gas fractions, that is, gas fractions that become a part of the product obtained in the cracking reforming reaction step, are processed again by the product recovery apparatus 2 shown in FIG. That is, these gas fractions are subjected to the pyrolysis heavy oil separation step 3 together with the cracked product obtained in the cracking furnace 1.
- hydrogen and methane are separated and recovered mainly by processing with the cracked gas compressor 4 and the demethanizer tower 6 and the like, and further ethylene is recovered by processing with the deethanizer tower 9 and the ethylene fractionator 10.
- propylene is recovered by treatment in the depropanizer 13 and the propylene fractionator 14, and treated in the depentane tower 17, the debutane tower 18 and the like, butylene, butadiene, etc., and cracked gasoline (BTX distillate). Min).
- the benzene, toluene, and xylene separated by the gas separation tower 36 shown in FIG. 2 are supplied to the BTX purification apparatus 20 shown in FIG. 1, and purified and rectified into benzene, toluene, and xylene, respectively, and separated and recovered as products. To do. Further, in the present embodiment, BTX is collected together, but may be collected separately depending on the apparatus configuration at the subsequent stage. For example, xylene may be supplied directly to a paraxylene production apparatus, not a BTX purification apparatus.
- the heavy fraction (bottom fraction) having 9 or more carbon atoms separated by the BTX fraction collection tower 35 is returned to the hydrogenation reactor 31 by a recycling path 37 (recycling process) as a recycling means. Together with the light pyrolysis heavy oil derived from the distillation column 30, it is again subjected to the hydrogenation reaction step. That is, this heavy fraction (bottom fraction) is returned to the cracking and reforming reaction device 33 via the hydrogenation reaction device 31 and used for the cracking and reforming reaction step.
- a heavy component having a distillation property of 90% by volume distillation temperature (T90) exceeding 390 ° C. is supplied to the hydrogenation reactor 31 (hydrogenation reaction step).
- the purification / recovery of the cracking / reforming reaction product derived from the cracking / reforming reaction apparatus 33 and the recycling to the cracking / reforming reaction step have been described. It is also possible to return to the recovery device 2 and perform recovery processing, in which case the installation of the purification recovery device 34 is not necessary.
- a heavy fraction having 9 or more carbon atoms (bottom fraction) obtained from the bottom of the BTX fraction collection tower 35 is recycled to the hydrogenation reactor 31, and a fraction having 8 or less carbon atoms obtained from the tower top. May be returned to the product recovery apparatus 2 of the ethylene production apparatus and processed in a lump.
- the heat obtained from the ethylene production apparatus Since the raw material oil composed of cracked heavy oil is subjected to the cracking and reforming reaction by the cracking and reforming reaction device 33, a part of the obtained product is recovered by the product recovery device 2 of the ethylene production apparatus. Light olefins by-produced in the cracking and reforming reaction apparatus 33 can be easily recovered by the existing product recovery apparatus 2 without constructing a new apparatus. Therefore, light olefins can be produced with higher production efficiency while suppressing an increase in cost. Further, the BTX fraction can also be efficiently produced by the cracking and reforming reaction apparatus 33.
- a hydrogenation reaction device 31 that partially hydrogenates part of the raw material oil (light pyrolysis heavy oil) on the front side (front) of the cracking reforming reaction device 33 (cracking reforming reaction step). Therefore, it is possible to suppress the amount of hydrogen consumed in the hydrogenation reaction step and the amount of heat generated during the treatment, and further, the decomposition reforming reaction device 33 (decomposition reforming reaction). BTX can be more efficiently produced in the step).
- the heavy fraction having 9 or more carbon atoms in the product obtained in the cracking and reforming reaction device 33 is again passed through the hydrogenation reaction device 31 (hydrogenation reaction step). Since there is a recycling path (recycling means 37, recycling process) for returning to the cracking and reforming reaction device 33 (cracking reforming reaction step), the production of the BTX fraction is further performed in the device configuration for performing the cracking and reforming process 21.
- the efficiency can be increased, and the production efficiency of light olefins by the ethylene production apparatus can be increased.
- the methane acts as a diluent, so that the coke can be prevented from adhering to the catalyst surface. Therefore, it is possible to suppress a decrease in the activity of the catalyst, increase the production efficiency of the olefin and the BTX fraction, and reduce the cost required for the regeneration treatment of the catalyst.
- the cracking reforming reaction apparatus 33 two or more fixed bed reactors are used as the cracking reforming reaction apparatus 33, and the cracking reforming reaction and regeneration of the catalyst for producing olefins and monocyclic aromatic hydrocarbons are repeated while periodically switching them. Therefore, the BTX fraction can be produced with high production efficiency.
- a fixed bed reactor having a much lower apparatus cost than a fluidized bed reactor is used, the cost of the apparatus configuration used for the cracking and reforming process 21 can be sufficiently reduced.
- the light olefin produced together with the BTX fraction can be easily recovered by the existing product recovery device 2 of the ethylene production apparatus, the light olefin is produced with high production efficiency together with the BTX fraction. be able to.
- the cracking and reforming reaction apparatus 33 performs the cracking and reforming reaction, and a part of the obtained product is recovered by the product recovery apparatus 2 of the ethylene production apparatus. All the products obtained by the reaction may be recovered by the product recovery apparatus 2 of the ethylene production apparatus.
- the hydrogenation reaction device 31 hydrogenation reaction step
- only a part of the raw material oil (light pyrolysis heavy oil) is partially hydrogenated. You may make it partially hydrogenate in the apparatus 31 (hydrogenation reaction process).
- the raw material oil used for the cracking and reforming process is only the one derived from the pyrolytic heavy oil of the connected ethylene production apparatus, but the properties as the raw material oil described in the present application are used. As long as the above conditions are satisfied, oils derived from pyrolytic heavy oils from other ethylene production equipment may be combined and used as raw material oils.
- the obtained kneaded material was extruded into a shape of a cylinder having a diameter of 1.5 mm by an extrusion molding machine, dried at 110 ° C. for 1 hour, and then fired at 550 ° C. to obtain a molded carrier.
- An impregnation solution prepared by taking 300 g of the obtained molded carrier, adding molybdenum trioxide, cobalt nitrate (II) hexahydrate, phosphoric acid (concentration 85%) to 150 ml of distilled water and adding malic acid until dissolved. Impregnation while spraying.
- Catalyst A has a SiO 2 content of 1.9% by mass, a TiO 2 content of 2.0% by mass on a carrier basis, a MoO 3 loading of 22.9% by mass on a catalyst basis, and a CoO carrier.
- the amount was 2.5% by mass, and the amount of P 2 O 5 supported was 4.0% by mass.
- pyrolysis heavy oil A The physical property value, distillation property, aromatic content, etc. of the pyrolysis heavy oil obtained from an ethylene production apparatus were measured. The results are shown in Table 1. Pyrolyzed heavy oil A was prepared by fractionating only its light components by distillation operation. Further, pyrolytic heavy oil C was prepared by recovering unreacted oil produced as a by-product when a petroleum resin was produced from a heavy oil fraction lighter than pyrolytic heavy oil A. Furthermore, pyrolysis heavy oil D was prepared by separating and recovering only the light fraction from the mixed fraction of pyrolysis heavy oil A and pyrolysis heavy oil C by distillation. About pyrolysis heavy oil B and C, D, the physical-property value, distillation property, aromatic content rate, etc. were measured. The results are shown in Table 2.
- the catalyst A was charged into a fixed bed continuous flow reactor, and the catalyst was first presulfided. That is, a density of 0.8516 g / ml at 15 ° C., an initial boiling point of 231 ° C. in a distillation test, a final boiling point of 376 ° C., a sulfur content of 1.18% by mass as a sulfur atom based on the mass of a pre-sulfurized raw material oil, hue 1% by weight of DMDS based on the mass of the fraction is added to the fraction corresponding to straight-distilled gas oil of L1.5 (preliminary sulfurized feedstock), and this is continuously added to the catalyst A for 48 hours. Supplied.
- the obtained hydrogenated pyrolysis heavy oils are designated as B-1 and D-1, respectively, and their properties are shown in Table 3.
- the solution (B) was gradually added to the solution (A) while stirring the solution (A) at room temperature.
- the resulting mixture was vigorously stirred with a mixer for 15 minutes to break up the gel into a milky homogeneous fine state.
- this mixture was put into a stainless steel autoclave, and a crystallization operation was performed under self-pressure under the conditions of a temperature of 165 ° C., a time of 72 hours, and a stirring speed of 100 rpm.
- the product was filtered to recover the solid product, and washing and filtration were repeated 5 times using about 5 liters of deionized water.
- the solid substance obtained by filtration was dried at 120 ° C., and further calcined at 550 ° C. for 3 hours under air flow.
- the obtained fired product was confirmed to have an MFI structure. Further, the SiO 2 / Al 2 O 3 ratio (molar ratio) was 65 by X-ray fluorescence analysis (model name: Rigaku ZSX101e). In addition, the aluminum element contained in the lattice skeleton calculated from this result was 1.3% by mass.
- a first solution was prepared by dissolving 202 g of tetraethylammonium hydroxide aqueous solution (40% by mass) in 59.1 g of silicic acid (SiO 2: 89% by mass). This first solution was added to a second solution prepared by dissolving 0.74 g Al-pellets and 2.69 g sodium hydroxide in 17.7 g water. In this way, the first solution and the second solution are mixed, and the composition (molar ratio of oxide) is 2.4Na 2 O-20.0 (TEA) 2 -Al 2 O. 3 was obtained -64.0SiO 2 -612H 2 O reaction mixture.
- the reaction mixture was placed in a 0.3 L autoclave and heated at 150 ° C. for 6 days.
- the resulting product was then separated from the mother liquor and washed with distilled water.
- the obtained product was confirmed to be BEA type zeolite from the XRD pattern.
- the BEA type zeolite was calcined at 550 ° C. for 3 hours to obtain a proton type BEA zeolite.
- hydrothermal treatment was performed in an environment of a treatment temperature of 650 ° C., a treatment time of 6 hours, and water vapor of 100% by mass. Thereafter, 99.2 parts (400 kgf) of hydrothermal degradation catalyst obtained by mixing 9 parts of phosphorus-containing proton-type MFI zeolite, which was also hydrothermally treated, with 1 part of hydrothermally-treated phosphorus-supported proton type BEA zeolite.
- the tablet C was molded by applying pressure, coarsely pulverized, and adjusted to a size of 20 to 28 mesh to obtain a granular catalyst C.
- Example 8 when making each raw material oil contact-react with a catalyst, nitrogen was introduce
- Example 8 the same experiment was performed by changing the diluent to methane.
- Examples 1 to 8 in which pyrolytic heavy oil having a predetermined property was used as the feedstock were Comparative Examples 1 in which pyrolysis heavy oil having a boiling point exceeding 400 ° C. was used as the feedstock.
- olefins having 2 to 4 carbon atoms and monocyclic aromatic hydrocarbons (benzene, toluene, xylene) having 6 to 8 carbon atoms can be produced with good yield.
- Comparative Example 2 coke was excessively formed on the catalyst, and the reaction tube was blocked in the middle, and the evaluation could not be continued until 24 hours later. Therefore, in Examples 1 to 8 of the present invention, it was confirmed that olefins and BTX can be efficiently produced from pyrolytic heavy oil obtained from an ethylene production apparatus.
- Example 5 when Example 5 and Example 6 are compared, by partially hydrogenating the raw material, the olefin having 2 to 4 carbon atoms and the monocyclic aromatic hydrocarbon having 6 to 8 carbon atoms (benzene, It was confirmed that toluene, xylene) can be produced.
- the BTX yield was lower in Example 2 than in Example 5, it was confirmed that it was more efficient to use two or more reactors repeatedly while repeating reaction regeneration. .
- Example 8 is a result substantially the same as Example 2, the use of methane as a diluent can stably produce olefins and aromatic hydrocarbons without increasing coke on the catalyst. (The yield in Table 4 excludes methane gas used as a diluent).
- Example 9 The liquid product obtained in Example 4 was distilled to recover only the heavier than BTX.
- the recovered liquid was mixed with pyrolysis heavy oil B at a ratio of 2: 1 and again hydrogenated under the same conditions as those for obtaining hydrogenated pyrolysis heavy oil B-1, then the same conditions as in Example 4
- the catalytic activity was evaluated.
- the results are shown in Table 5. From the results shown in Table 5, it was confirmed that olefins and BTX can be more efficiently produced from pyrolyzed heavy oil obtained from an ethylene production apparatus by repeatedly using heavy components as raw materials.
- the present invention relates to a method for producing an olefin having 2 to 4 carbon atoms and a monocyclic aromatic hydrocarbon having 6 to 8 carbon atoms, and an ethylene production apparatus.
- BTX can be produced with higher production efficiency and while suppressing an increase in cost, and light olefins can also be produced efficiently.
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Abstract
Description
本願は2012年10月25日に日本に出願された特願2012-236132号に対して優先権を主張し、その内容をここに援用する。
この特許文献1の芳香族炭化水素の製造方法は、従来では前記熱分解重質油(分解重質油)がコンビナート内でボイラー等の燃料等に使われることがほとんどであったのに対し、前記熱分解重質油を水素化処理した後、単環芳香族炭化水素製造用触媒に接触させ反応させることで、BTXを製造するようにしている。
しかしながら、前記特許文献1の芳香族炭化水素の製造方法では、熱分解重質油から製造されるのは基本的にBTX留分のみであり、軽質オレフィンが製造可能なことは開示されていない。したがって、前記特許文献1の技術を採用すると、エチレン製造装置から得られる熱分解重質油からBTXを効率よく製造できる反面、軽質オレフィンが製造されないため、エチレン製造装置の本来の目的である軽質オレフィンの製造に関しては、結果的にその生産効率を高めたいとの要望に応えられなくなってしまう。
また、エチレン製造装置では、軽質オレフィンの生産効率を高めることが要望されているものの、当然ながらコストが上昇するのを抑えつつ、生産効率を高めることが求められている。
前記分解改質反応工程で得られた生成物の一部または全てを、前記エチレン製造装置の生成物回収装置で処理することにより、該生成物の一部または全てから、炭素数2~4のオレフィン並びに炭素数6~8の単環芳香族炭化水素をそれぞれ回収する生成物回収工程と、を有する。
また、前記製造方法において、前記水素化反応工程では、前記原料油を水素化する水素化条件として、水素分圧を1~9MPa、水素化温度を150~400℃とするとともに、水素化触媒として、アルミニウム酸化物を含む無機担体に全触媒質量を基準として周期表第6族金属から選択される少なくとも1種の金属を10~30質量%と、周期表第8~10族金属から選択される少なくとも1種の金属を1~7質量%とを担持させて得られる触媒を用いる、ことが好ましい。
また、前記製造方法において、前記生成物回収工程では、前記分解改質反応工程で得られた生成物の一部を、前記エチレン製造装置の生成物回収装置で処理するようにし、前記分解改質反応工程で得られた生成物のうちの炭素数9以上の重質留分を、前記分解改質反応工程に戻すリサイクル工程を有する、ことが好ましい。
また、前記製造方法において、前記分解改質反応工程では、炭素数1~3の飽和炭化水素を共存させた状態で、前記原料油を反応させることが好ましい。
また、前記製造方法において、前記分解改質反応工程では、2基以上の固定床反応器を用い、これらを定期的に切り替えながら分解改質反応と前記オレフィン及び単環芳香族炭化水素製造用触媒の再生とを交互にもしくは順次繰り返す、ことが好ましい。
また、前記製造方法においては、前記分解改質反応工程で用いるオレフィン及び単環芳香族炭化水素製造用触媒に含有される結晶性アルミノシリケートが、中細孔ゼオライト及び/又は大細孔ゼオライトを主成分としたものであることが好ましい。
また、前記製造方法においては、前記分解改質反応工程で用いるオレフィン及び単環芳香族炭化水素製造用触媒が、リンを含むことが好ましい。
前記分解炉で生成した分解生成物から水素、エチレン、プロピレン、C4留分、炭素数6~8の単環芳香族炭化水素を含む留分をそれぞれ分離回収する生成物回収装置と、
前記分解炉から得られる熱分解重質油でかつ蒸留性状の90容量%留出温度が390℃以下のものを原料油とし、この原料油に対して結晶性アルミノシリケートを含むオレフィン及び単環芳香族炭化水素製造用触媒に接触させ、反応させて、炭素数2~4のオレフィン並びに炭素数6~8の単環芳香族炭化水素を含む生成物を得る分解改質反応装置と、
前記分解改質反応装置で得られた生成物の一部または全てを、前記生成物回収装置に供給する生成物供給手段とを備える。
また、前記製造装置において、前記生成物供給手段は、前記分解改質反応装置で得られた生成物の一部を、前記生成物回収装置に供給するように構成され、前記分解改質反応装置で得られた生成物のうちの炭素数9以上の重質留分を、前記分解改質反応装置に戻すリサイクル手段を有する、ことが好ましい。
また、前記製造装置において、前記分解改質反応装置は、2基以上の固定床反応器を備え、これらが定期的に切り替えられながら分解改質反応と前記オレフィン及び単環芳香族炭化水素製造用触媒の再生とを交互にもしくは順次繰り返すよう構成されていることが好ましい。
まず、本発明に係るエチレン製造装置の一実施形態の概略構成と、本発明の製造方法に係るプロセスについて、図1を参照して説明する。
なお、本発明に係るエチレン製造装置の実施形態のうち、図2に示す分解改質プロセス以外の部分は、分解工程と分離精製工程を備えた公知のエチレン製造装置であってよく、一例として非特許文献1に記載されたエチレン製造装置をあげることができる。従って、本発明に係るエチレン製造装置の実施形態には、既存のエチレン製造装置に本発明の分解改質プロセスを追加したものも含まれる。
なお、回収されたエチレンは、エチレン製造装置で製造する主製品となる。また、回収されたエタンは、ナフサ留分や灯油留分、軽油留分等の原料とともに分解炉1に供給し、リサイクルすることもできる。
本発明における熱分解重質油の性状としては、特に規定されないものの、以下の性状を有することが好ましい。
蒸留試験により得られる性状は、分解温度や分解原料により大きく変動するが、10容量%留出温度(T10)は、145℃以上230℃以下のものが好ましく使用される。90容量%留出温度(T90)並びに終点に関しては、用いる留分によりさらに大きく変化するため制限はないが、熱分解重質油分離工程3から直接得られる留分であれば、例えば90容量%留出温度(T90)は400℃以上600℃以下、終点(EP)は450℃以上800℃以下の範囲のものが好ましく使用される。
本発明に係る原料油は、前記したエチレン製造装置から得られる熱分解重質油で、かつ、蒸留性状の90容量%留出温度が390℃以下のものである。すなわち、前留塔30にて蒸留処理され、蒸留性状の90容量%留出温度が390℃以下に調整された軽質熱分解重質油が、原料油として用いられる。このように90容量%留出温度を390℃以下にすることで、原料油は炭素数が9~12の芳香族炭化水素が主となり、後述するオレフィン及び単環芳香族炭化水素製造用触媒との接触及び反応による分解改質反応工程において、オレフィンおよびBTX留分の収率を高めることができる。また、オレフィンおよびBTX留分の収率をより高めるためには、好ましくは10容量%留出温度(T10)が140℃以上220℃以下、90容量%留出温度(T90)が220℃以上380℃以下、より好ましくはT10が160℃以上200℃以下、T90が240℃以上350℃以下である。
なお、分解改質プロセス21に供される際に原料油蒸留性状の90容量%留出温度(T90)が390℃以下である場合は、必ずしも前留塔30にて蒸留処理する必要はない。
なお、本発明に係る原料油は、エチレン製造装置から得られる熱分解重質油を含むものであれば、他の基材を含むものであってもよい。
また、蒸留性状の90容量%留出温度(T90)が390℃以下に調整されている留分は、必ずしも前留塔30にて蒸留処理をする必要がない。そのため、後述するように図2に示す熱分解重質油とは別に、前留塔30の後段側にて分解改質プロセス21を構成する装置である水素化反応装置31あるいは分解改質反応装置33に直接供給することも可能である。
本実施形態では、前記の軽質熱分解重質油のみ、すなわち原料油の一部のみを部分水素化処理する。熱分解重質油を複数の留分に分離した際の一部の留分あるいはこれらの分離した留分から他の化学品または燃料を製造した際の残油等のうち炭素数9の炭化水素を主とする成分や分解ガソリン回収部19にて分離回収された炭素数9以上の成分については、水素化処理を省略できる。ただし、これらの成分についても、水素化反応装置31によって部分水素化処理してもよいのはもちろんである。
エチレン製造装置から得られる熱分解重質油は、通常、芳香族炭化水素の含有量が非常に多い。そこで、本実施形態では、先に分離した熱分解重質油中の必要な留分、すなわち軽質HARを原料油とし、この原料油を水素化反応装置31(水素化反応工程)にて水素化処理する。ただし、原料油を水素化分解するまで水素化処理するには多量の水素が必要になるとともに、完全に水素化された原料油を用いると後述するオレフィン及び単環芳香族炭化水素製造用触媒との接触及び反応による分解改質反応工程におけるオレフィンおよびBTX留分の製造効率が極めて低くなってしまう。
一方、水素/油比が2000NL/Lを超える場合には、リサイクルコンプレッサーの建設費が過大になり、経済性が損なわれる懸念がある。
予備硫化の条件としては、特に限定されないものの、石油留分または熱分解重質油(以下、予備硫化原料油という。)に硫黄化合物を添加し、これを温度200~380℃、LHSVが1~2h-1、圧力は水素化処理運転時と同一、処理時間48時間以上の条件にて、前記水素化処理用触媒に連続的に接触せしめることが好ましい。前記予備硫化原料油に添加する硫黄化合物としては、限定されないものの、ジメチルジスルフィド(DMDS)、サルファゾール、硫化水素等が好ましく、これらを予備硫化原料油に対して予備硫化原料油の質量基準で1質量%程度添加することが好ましい。
以上に説明した水素化反応装置31(水素化反応工程)から得られる、原料油の水素化処理油は、以下の性状を有することが好ましい。
蒸留性状は、10容量%留出温度(T10)が140℃以上200℃以下、90容量%留出温度(T90)が200℃以上390℃以下、より好ましくはT10が160℃以上190℃以下、T90が210℃以上370℃以下である。T10が140℃未満では、この水素化処理油を含んで形成される原料油に、目的物の一つであるキシレンを含有する可能性があるため、好ましくない。一方、T90が390℃を超える(重質になる)と、水素化処理触媒への金属被毒、コーク析出等により触媒性能が低下すること、及び後述するオレフィン及び単環芳香族炭化水素製造用触媒へのコーク析出が多くなり所定の性能が出なくなること、水素消費量が多くなり経済的でなくなること、といった点から好ましくない。
(オレフィン及び単環芳香族炭化水素製造用触媒)
オレフィン及び単環芳香族炭化水素製造用触媒は、結晶性アルミノシリケートを含むものである。触媒の結晶性アルミノシリケートの含有量は、必要とされる分解改質反応の反応性や選択性もしくは触媒の形状や強度に応じて決定すればよく、特に限定されないものの、10~100質量%が好ましい。固定床反応器に用いるので結晶性アルミノシリケートのみからなる触媒であってよい。強度を高めるためにバインダーを添加するのであれば、結晶性アルミノシリケートの含有量は20~95質量%が好ましく、25~90質量%がより好ましい。しかし、結晶性アルミノシリケートの含有量が10%を下回ると、十分な触媒活性を得るための触媒量が過大となるため好ましくない。
結晶性アルミノシリケートとしては、単環芳香族炭化水素の収率をより高くできることから、中細孔ゼオライト及び/又は大細孔ゼオライトを主成分としたものであることが好ましい。
中細孔ゼオライトは、10員環の骨格構造を有するゼオライトであり、中細孔ゼオライトとしては、例えば、AEL型、EUO型、FER型、HEU型、MEL型、MFI型、NES型、TON型、WEI型の結晶構造のゼオライトが挙げられる。これらの中でも、単環芳香族炭化水素の収率をより高くできることから、MFI型が好ましい。
大細孔ゼオライトは、12員環の骨格構造を有するゼオライトであり、大細孔ゼオライトとしては、例えば、AFI型、ATO型、BEA型、CON型、FAU型、GME型、LTL型、MOR型、MTW型、OFF型の結晶構造のゼオライトが挙げられる。これらの中でも、工業的に使用できる点では、BEA型、FAU型、MOR型が好ましく、単環芳香族炭化水素の収率をより高くできることから、BEA型が好ましい。
ここで、小細孔ゼオライトとしては、例えば、ANA型、CHA型、ERI型、GIS型、KFI型、LTA型、NAT型、PAU型、YUG型の結晶構造のゼオライトが挙げられる。
超大細孔ゼオライトとしては、例えば、CLO型、VPI型の結晶構造のゼオライトが挙げられる。
また、結晶性アルミノシリケートのSi/Al比は、単環芳香族炭化水素の十分な収率を得るためには、10以上であることが好ましい。
ガリウムおよび/または亜鉛を含む結晶性アルミノシリケートとしては、結晶性アルミノシリケートの格子骨格内にガリウムが組み込まれたもの(結晶性アルミノガロシリケート)、結晶性アルミノシリケートの格子骨格内に亜鉛が組み込まれたもの(結晶性アルミノジンコシリケート)、結晶性アルミノシリケートにガリウムを担持したもの(Ga担持結晶性アルミノシリケート)、結晶性アルミノシリケートに亜鉛を担持したもの(Zn担持結晶性アルミノシリケート)、それらを少なくとも1種以上含んだものが挙げられる。
ガリウム及び/又は亜鉛の含有量が5質量%を超えると、単環芳香族炭化水素の収率が低くなるため好ましくない。
また、ガリウム及び/又は亜鉛の含有量の下限は、触媒全量を100質量%とした場合、0.01質量%以上であることが好ましく、0.1質量%以上であることがより好ましい。ガリウム及び/又は亜鉛の含有量が0.01質量%未満であると、単環芳香族炭化水素の収率が低くなることがあり好ましくない。
具体的には、固定床で用いる場合、バインダーとしてはシリカ、アルミナなどの無機物質が好ましく用いられる。
また、オレフィン及び単環芳香族炭化水素製造用触媒がバインダーを含有する場合、バインダーとガリウム及び/又は亜鉛担持結晶性アルミノシリケートとを混合した後、またはバインダーと結晶性アルミノガロシリケート及び/又は結晶性アルミノジンコシリケートとを混合した後に、リンを添加して触媒を製造してもよい。
分解改質反応装置33の反応形式、すなわち分解改質反応装置33によって前記原料油をオレフィン及び単環芳香族炭化水素製造用触媒と接触させ、分解改質反応させる際の反応形式としては、固定床、移動床、流動床等が挙げられる。
特に、固定床は流動床や移動床に比べて装置コストが格段に安価であり、好ましい。したがって、固定床反応器1基で反応と再生を繰り返す事も可能であるが、反応を連続して行うために2基以上の反応器を設置するのがよい。本実施形態では、図2に示すように固定床の分解改質反応装置33(固定床反応器)を用いるとともに、この固定床反応器33を2基用いている。なお、図2では固定床反応器33を2基記載しているが、これに限定されることなく、2基以上であれば任意の数、設置することができる。
原料油を触媒と接触、反応させる際の反応温度は、特に制限されないものの、350~700℃が好ましく、400~650℃がより好ましい。反応温度が350℃未満では、反応活性が十分でない。反応温度が700℃を超えると、エネルギー的に不利になると同時に、コーク生成が著しく増大し目的物の製造効率が低下する。
原料油を触媒と接触、反応させる際の反応圧力は、0.1MPaG~2.0MPaGである。すなわち、原料油とオレフィン及び単環芳香族炭化水素製造用触媒との接触を、0.1MPaG~2.0MPaGの圧力下で行う。
本発明は、水素化分解による従来の方法とは反応思想が全く異なるため、水素化分解では優位とされる高圧条件を全く必要としない。むしろ、必要以上の高圧は、分解を促進し、目的としない軽質ガスを副生するため好ましくない。また、高圧条件を必要としないことは、反応装置設計上においても優位である。そのため、反応圧力が0.1MPaG~2.0MPaGであれば、分解改質反応を効率的に行うことが可能である。
原料油と触媒との接触時間は、実質的に所望する反応が進行すれば特に制限されないものの、例えば、触媒上のガス通過時間で2~150秒が好ましく、3~100秒がより好ましく、5~80秒がさらに好ましい。接触時間が5秒未満では、実質的な反応が困難である。接触時間が300秒を超えると、コーキング等による触媒への炭素質の蓄積が多くなる、または分解による軽質ガスの発生量が多くなり、さらには装置も巨大となり好ましくない。
分解改質反応装置33によって分解改質反応処理(分解改質反応工程)を所定時間行ったら、分解改質反応処理の運転は別の分解改質反応装置33に切り替え、分解改質反応処理の運転を停止した分解改質反応装置33については、活性が低下したオレフィン及び単環芳香族炭化水素製造用触媒の再生を行う。
なお、分解改質反応装置33での分解改質反応処理においては、触媒表面へのコークの付着を抑制するため、炭素数1~3の飽和炭化水素、例えばメタンを、図2に示すように分解改質反応装置33に供して該メタンを共存させた状態で、原料油を処理するのが好ましい。メタンは、ほとんど反応性がなく、したがって分解改質反応装置33内にて前記触媒と接触しても、反応を起こすことがない。よって、原料油に由来する重質の炭化水素が触媒表面に付着し触媒反応が進むのを、メタンは触媒表面での前記炭化水素の濃度を下げる希釈剤として作用することにより、これを抑制(妨害)する。したがって、メタンは、原料油に由来する重質の炭化水素が触媒表面に付着してコークとなるのを、抑制するようになる。
分解改質反応装置33から導出された分解改質反応生成物には、炭素数2~4のオレフィンを含有するガス、BTX留分、C9以上の芳香族炭化水素が含まれる。そこで、分解改質反応装置33の後段に設けられた精製回収装置34により、この分解改質反応生成物を各成分に分離し、精製回収する。
BTX留分塔35は、前記の分解改質反応生成物を蒸留し、炭素数8以下の軽質留分と炭素数9以上の重質留分とに分離する。ガス分離塔36は、BTX留分回収塔35で分離された炭素数8以下の軽質留分を蒸留し、ベンゼン、トルエン、粗キシレンを含むBTX留分と、これらより低沸点のガス留分とに分離する。なお、これらBTX留分回収塔35、ガス分離塔36では、後述するようにそれぞれで得られる留分を再処理するため、その蒸留精度を高める必要はなく、蒸留操作を比較的大まかに行うことができる。
前記したようにガス分離塔36では、その蒸留操作を比較的大まかに行っているため、ガス分離塔36で分離されたガス留分には、主に、水素、エチレン、プロピレン、ブチレン等のC4留分、BTXが含まれる。そこで、これらガス留分、すなわち前記分解改質反応工程で得られた生成物の一部となるガス留分を、図1に示した生成物回収装置2で再度処理する。すなわち、これらガス留分を、分解炉1で得られた分解生成物とともに、熱分解重質油分離工程3に供する。そして、主に分解ガスコンプレッサー4、脱メタン塔6等にて処理することで水素やメタンを分離回収し、さらに脱エタン塔9、エチレン精留塔10にて処理することでエチレンを回収する。また、脱プロパン塔13、プロピレン精留塔14にて処理することでプロピレンを回収し、脱ペンタン塔17、脱ブタン塔18等にて処理することでブチレンやブタジエンなどと、分解ガソリン(BTX留分)を回収する。
また、BTX留分回収塔35で分離された炭素数9以上の重質留分(ボトム留分)については、リサイクル手段としてのリサイクル路37(リサイクル工程)によって水素化反応装置31に戻し、前留塔30から導出される軽質熱分解重質油とともに再度水素化反応工程に供する。すなわち、この重質留分(ボトム留分)は、水素化反応装置31を経て分解改質反応装置33に戻され、分解改質反応工程に供されるようになる。なお、リサイクル工程(リサイクル路37)では、例えば蒸留性状の90容量%留出温度(T90)が390℃を超えるような重質分については、水素化反応装置31(水素化反応工程)に供する前にカットバックし、重質熱分解重質油とともに貯留するのが好ましい。90容量%留出温度(T90)が390℃を超える留分がほとんど含まれない場合でも、反応性の低い留分が蓄積される場合などは、一定量を系外に排出することが好ましい。
例えば、前記実施形態では、分解改質反応装置33によって分解改質反応させ、得られた生成物の一部をエチレン製造装置の生成物回収装置2で回収処理するようにしたが、分解改質反応によって得られた生成物の全てを、エチレン製造装置の生成物回収装置2で回収処理するようにしてもよい。
また、水素化反応装置31(水素化反応工程)では、原料油の一部(軽質熱分解重質油)のみを部分的に水素化するようにしたが、原料油の全てを、水素化反応装置31(水素化反応工程)にて部分的に水素化するようにしてもよい。
また、水素化反応装置31(水素化反応工程)で使用する水素としては、水素回収部7で回収された水素だけでなく、公知の水素製造方法で得た水素を利用してもよい。
また、前記実施形態では、分解改質プロセスに供される原料油は連結されたエチレン製造装置の熱分解重質油に由来するもののみを用いたが、本願に記載された原料油としての性状を満たす限り、他のエチレン製造装置からの熱分解重質油に由来する油を合わせて原料油としてもよい。
濃度5質量%のアルミン酸ナトリウム水溶液1kgに水ガラス3号を加え、70℃に保温した容器に入れた。また、濃度2.5質量%の硫酸アルミニウム水溶液1kgに硫酸チタン(IV)水溶液(TiO2含有量として24質量%)を加えた溶液を、70℃に保温した別の容器において調製し、この溶液を、上述のアルミン酸ナトリウムを含む水溶液に15分間で滴下した。上記水ガラスおよび硫酸チタン水溶液の量は、所定のシリカ、チタニアの含有量となるように調整した。
使用する三酸化モリブデン、硝酸コバルト(II)6水和物およびリン酸の量は、所定の担持量となるよう調整した。含浸溶液に含浸した試料を110℃で1時間乾燥した後、550℃で焼成し、触媒Aを得た。触媒Aは、担体基準で、SiO2の含有量が1.9質量%、TiO2の含有量が2.0質量%、触媒基準でMoO3の担持量が22.9質量%、CoOの担持量が2.5質量%、P2O5担持量が4.0質量%であった。
エチレン製造装置から得た熱分解重質油(熱分解重質油Aとする。)の、物性値、蒸留性状、芳香族含有率等を測定した。結果を表1に示す。熱分解重質油Aを蒸留操作によりその軽質分のみを分取して熱分解重質油Bを調製した。また、熱分解重質油Aよりも軽質な重質油留分から石油樹脂を製造した際に副生する未反応油を回収することにより熱分解重質油Cを調製した。さらに、熱分解重質油Aと熱分解重質油Cとの混合留分から軽質分のみを蒸留により分離回収することで熱分解重質油Dを調製した。熱分解重質油BおよびC、Dについて、物性値、蒸留性状、芳香族含有率等を測定した。結果を表2に示す。
固定床連続流通式反応装置に上記触媒Aを充填し、まず触媒の予備硫化を行った。すなわち、15℃における密度0.8516g/ml、蒸留試験における初留点231℃、終留点376℃、予備硫化原料油の質量を基準とした硫黄原子としての硫黄分1.18質量%、色相L1.5である直留系軽油相当の留分(予備硫化原料油)に、該留分の質量基準で1質量%のDMDSを添加し、これを48時間前記触媒Aに対して連続的に供給した。
その後、表2に示す熱分解重質油Bおよび熱分解重質油Dを原料油として用い、反応温度300℃、LHSV=1.0h-1、水素油比500NL/L、圧力3MPaにて水素化処理を行った。得られた水素化熱分解重質油をそれぞれB-1、D-1とし、その性状を表3に示す。
また、表1、2、3の各組成は、シリカゲルクロマト分別により得た飽和分および芳香族分について、EIイオン化法による質量分析(装置:日本電子(株)製、JMS-700)を行い、ASTM D2425“Standard Test Method for Hydrocarbon Types in Middle Distillates by Mass Spectrometry”に準拠して炭化水素のタイプ分析により算出した。
硅酸ナトリウム(Jケイ酸ソーダ3号、SiO2:28~30質量%、Na:9~10質量%、残部水、日本化学工業(株)製)の1706.1gおよび水の2227.5gからなる溶液(A)と、Al2(SO4)3・14~18H2O(試薬特級、和光純薬工業(株)製)の64.2g、テトラプロピルアンモニウムブロマイドの369.2g、H2SO4(97質量%)の152.1g、NaClの326.6gおよび水の2975.7gからなる溶液(B)をそれぞれ調製した。
得られた混合物をミキサーで15分間激しく撹拌し、ゲルを解砕して乳状の均質微細な状態にした。
次いで、この混合物をステンレス製のオートクレーブに入れ、温度を165℃、時間を72時間、撹拌速度を100rpmとする条件で、自己圧力下に結晶化操作を行った。結晶化操作の終了後、生成物を濾過して固体生成物を回収し、約5リットルの脱イオン水を用いて洗浄と濾過を5回繰り返した。濾別して得られた固形物を120℃で乾燥し、さらに空気流通下、550℃で3時間焼成した。
59.1gのケイ酸(SiO2 :89質量%)に四エチルアンモニウムヒドロオキシド水溶液(40質量%)を202gに溶解することにより、第一の溶液を調製した。この第一の溶液を、0.74gのAl-ペレット及び2.69gの水酸化ナトリウムを17.7gの水に溶解して調製した第二の溶液に加えた。このようにして第一の溶液と第二の溶液の二つの溶液を混合して、組成(酸化物のモル比換算)が、2.4Na2O-20.0(TEA)2-Al2O3-64.0SiO2-612H2Oの反応混合物を得た。
この反応混合物を0.3Lオートクレーブに入れ、150℃で6日間加熱した。そして、得られた生成物を母液から分離し、蒸留水で洗った。
得られた生成物は、X線回析分析(機種名:Rigaku RINT-2500V)の結果、XRDパターンよりBEA型ゼオライトであることが確認された。
その後、硝酸アンモニウム水溶液(30質量%)でイオン交換した後、BEA型ゼオライトを550℃で3時間焼成を行い、プロトン型BEAゼオライトを得た。
次いで、得られたプロトン型BEAゼオライト30gに、2.0質量%のリン(プロトン型BEAゼオライト総質量を100質量%とした値)が担持されるようにリン酸水素二アンモニウム水溶液30gを含浸させ、120℃で乾燥した。その後、空気流通下、780℃で3時間焼成して、プロトン型BEAゼオライトとリンとを含有する触媒を得た。得られた触媒の初期活性における影響を排除するため、処理温度650℃、処理時間6時間、水蒸気100質量%の環境下で水熱処理を実施した。その後、水熱処理したリン担持プロトン型BEAゼオライト1部に対して、同じく水熱処理したリン含有プロトン型MFIゼオライト9部を混合する事により得られた水熱劣化処理触媒に39.2MPa(400kgf)の圧力をかけて打錠成型し、粗粉砕して20~28メッシュのサイズに揃えて、粒状体の触媒Cを得た。
触媒BまたはC(10ml)を反応器に充填した流通式反応装置を用い、反応温度を550℃、反応圧力を0.1MPaG、LHSV=1とする分子状水素非共存下の条件のもとで、表4に示す各原料油を対応する触媒と接触、反応させた。用いた原料油と触媒との組み合わせにより、表4に示すように実施例1~8、および比較例1、2とした。なお、各原料油を触媒と接触反応させる際、希釈剤として、原料油に対して窒素を容積で1:1となるように導入した。実施例8においては希釈材をメタンに変更して同様の実験を行った。
したがって、本発明の実施例1~8では、エチレン製造装置から得られる熱分解重質油から、オレフィン並びにBTXを効率よく製造できることが確認された。
また、実施例5に比べ実施例2ではBTX収率が低下していることから、反応再生を繰り返しながら2基以上の反応器を繰返し利用することが、より効率的であることが確認された。
また、実施例8は実施例2とほぼ同等の結果であることから、希釈材としてメタンを用いることで触媒上のコークを増加させることなく、安定的にオレフィン並びに芳香族炭化水素を製造できることが確認された(表4中の収率は、希釈剤として用いたメタンガスを除いている。)。
実施例4で得られた液生成物を蒸留し、BTXよりも重質分のみを回収した。回収液を熱分解重質油Bと2:1の割合で混合させ、再度、水素化熱分解重質油B-1を得た条件と同条件で水素化した後に、実施例4と同条件で触媒活性を評価した。その結果を表5に示す。表5に示す結果より、重質分を繰返し原料として用いる事で、エチレン製造装置から得られる熱分解重質油から、オレフィン並びにBTXをより効率よく製造できることが確認された。
Claims (12)
- 分解炉と、該分解炉で生成した分解生成物から水素、エチレン、プロピレン、C4留分、炭素数6~8の単環芳香族炭化水素を含む留分をそれぞれ分離回収する生成物回収装置と、を備えたエチレン製造装置より得られる熱分解重質油であって且つ蒸留性状の90容量%留出温度が390℃以下の原料油を、結晶性アルミノシリケートを含むオレフィン及び単環芳香族炭化水素製造用触媒と接触させ、反応させて、炭素数2~4のオレフィン並びに炭素数6~8の単環芳香族炭化水素を含む生成物を得る分解改質反応工程と、
前記分解改質反応工程で得られた生成物の一部または全てを、前記エチレン製造装置の生成物回収装置で処理することにより、該生成物の一部または全てから、炭素数2~4のオレフィン並びに炭素数6~8の単環芳香族炭化水素をそれぞれ回収する生成物回収工程と、を有するオレフィン及び単環芳香族炭化水素の製造方法。 - 前記分解改質反応工程の前に、前記原料油の一部または全てを部分水素化する水素化反応工程を有する、請求項1記載のオレフィン及び単環芳香族炭化水素の製造方法。
- 前記水素化反応工程では、前記原料油を水素化する水素化条件として、水素分圧を1~9MPa、水素化温度を150~400℃とするとともに、水素化触媒として、アルミニウム酸化物を含む無機担体に全触媒質量を基準として周期表第6族金属から選択される少なくとも1種の金属を10~30質量%と、周期表第8~10族金属から選択される少なくとも1種の金属を1~7質量%とを担持させて得られる触媒を用いる、請求項2記載のオレフィン及び単環芳香族炭化水素の製造方法。
- 前記生成物回収工程では、前記分解改質反応工程で得られた生成物の一部を、前記エチレン製造装置の生成物回収装置で処理するようにし、
前記分解改質反応工程で得られた生成物のうちの炭素数9以上の重質留分を、前記分解改質反応工程に戻すリサイクル工程を有する、請求項1~3のいずれか一項に記載のオレフィン及び単環芳香族炭化水素の製造方法。 - 前記分解改質反応工程では、炭素数1~3の飽和炭化水素を共存させた状態で、前記原料油を反応させる請求項1~4のいずれか一項に記載のオレフィン及び単環芳香族炭化水素の製造方法。
- 前記分解改質反応工程では、2基以上の固定床反応器を用い、これらを定期的に切り替えながら分解改質反応と前記オレフィン及び単環芳香族炭化水素製造用触媒の再生とを繰り返す、請求項1~5のいずれか一項に記載のオレフィン及び単環芳香族炭化水素の製造方法。
- 前記分解改質反応工程で用いるオレフィン及び単環芳香族炭化水素製造用触媒に含有される結晶性アルミノシリケートが、中細孔ゼオライト及び/又は大細孔ゼオライトを主成分としたものである請求項1~6のいずれか一項に記載のオレフィン及び単環芳香族炭化水素の製造方法。
- 前記分解改質反応工程で用いるオレフィン及び単環芳香族炭化水素製造用触媒が、リンを含む請求項1~7のいずれか一項に記載のオレフィン及び単環芳香族炭化水素の製造方法。
- 分解炉と、
前記分解炉で生成した分解生成物から水素、エチレン、プロピレン、C4留分、炭素数6~8の単環芳香族炭化水素を含む留分をそれぞれ分離回収する生成物回収装置と、
前記分解炉から得られる熱分解重質油でかつ蒸留性状の90容量%留出温度が390℃以下のものを原料油とし、この原料油に対して結晶性アルミノシリケートを含むオレフィン及び単環芳香族炭化水素製造用触媒に接触させ、反応させて、炭素数2~4のオレフィン並びに炭素数6~8の単環芳香族炭化水素を含む生成物を得る分解改質反応装置と、
前記分解改質反応装置で得られた生成物の一部または全てを、前記生成物回収装置に供給する生成物供給手段と、を備えるエチレン製造装置。 - 前記分解改質反応装置の前に、前記原料油の一部または全てを部分水素化する水素化反応装置を有する、請求項9記載のエチレン製造装置。
- 前記生成物供給手段は、前記分解改質反応装置で得られた生成物の一部を、前記生成物回収装置に供給するように構成され、
前記分解改質反応装置で得られた生成物のうちの炭素数9以上の重質留分を、前記分解改質反応装置に戻すリサイクル手段を有する、請求項9又は10に記載のエチレン製造装置。 - 前記分解改質反応装置は、2基以上の固定床反応器を備え、これらが定期的に切り替えられながら分解改質反応と前記オレフィン及び単環芳香族炭化水素製造用触媒の再生とを繰り返すよう構成されている請求項9~11のいずれか一項に記載のエチレン製造装置。
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