CN114456839B - Coking naphtha processing method - Google Patents
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- 238000004939 coking Methods 0.000 title claims abstract description 18
- 238000003672 processing method Methods 0.000 title claims abstract description 9
- 238000005984 hydrogenation reaction Methods 0.000 claims abstract description 175
- 239000003054 catalyst Substances 0.000 claims abstract description 83
- 238000006243 chemical reaction Methods 0.000 claims abstract description 49
- 229910052739 hydrogen Inorganic materials 0.000 claims abstract description 47
- 239000002994 raw material Substances 0.000 claims abstract description 47
- 239000001257 hydrogen Substances 0.000 claims abstract description 46
- UFHFLCQGNIYNRP-UHFFFAOYSA-N Hydrogen Chemical compound [H][H] UFHFLCQGNIYNRP-UHFFFAOYSA-N 0.000 claims abstract description 44
- 239000007791 liquid phase Substances 0.000 claims abstract description 33
- 238000001704 evaporation Methods 0.000 claims abstract description 26
- 230000008020 evaporation Effects 0.000 claims abstract description 26
- 239000007795 chemical reaction product Substances 0.000 claims abstract description 21
- 239000000463 material Substances 0.000 claims abstract description 19
- 239000012071 phase Substances 0.000 claims abstract description 17
- 239000007789 gas Substances 0.000 claims abstract description 16
- 239000000047 product Substances 0.000 claims abstract description 13
- 150000001993 dienes Chemical class 0.000 claims abstract description 10
- 238000000034 method Methods 0.000 claims description 45
- 238000004073 vulcanization Methods 0.000 claims description 28
- 239000002184 metal Substances 0.000 claims description 26
- 229910052751 metal Inorganic materials 0.000 claims description 26
- XUIMIQQOPSSXEZ-UHFFFAOYSA-N Silicon Chemical compound [Si] XUIMIQQOPSSXEZ-UHFFFAOYSA-N 0.000 claims description 23
- 229910052710 silicon Inorganic materials 0.000 claims description 23
- 239000010703 silicon Substances 0.000 claims description 23
- 239000003795 chemical substances by application Substances 0.000 claims description 22
- 239000003921 oil Substances 0.000 claims description 22
- 238000001035 drying Methods 0.000 claims description 19
- 150000001336 alkenes Chemical class 0.000 claims description 15
- PNEYBMLMFCGWSK-UHFFFAOYSA-N aluminium oxide Inorganic materials [O-2].[O-2].[O-2].[Al+3].[Al+3] PNEYBMLMFCGWSK-UHFFFAOYSA-N 0.000 claims description 14
- JRZJOMJEPLMPRA-UHFFFAOYSA-N olefin Natural products CCCCCCCC=C JRZJOMJEPLMPRA-UHFFFAOYSA-N 0.000 claims description 13
- 238000007670 refining Methods 0.000 claims description 8
- RWSOTUBLDIXVET-UHFFFAOYSA-N Dihydrogen sulfide Chemical group S RWSOTUBLDIXVET-UHFFFAOYSA-N 0.000 claims description 7
- 238000006477 desulfuration reaction Methods 0.000 claims description 7
- 229910000037 hydrogen sulfide Inorganic materials 0.000 claims description 7
- 230000023556 desulfurization Effects 0.000 claims description 6
- WQOXQRCZOLPYPM-UHFFFAOYSA-N dimethyl disulfide Chemical compound CSSC WQOXQRCZOLPYPM-UHFFFAOYSA-N 0.000 claims description 6
- 238000000926 separation method Methods 0.000 claims description 5
- 229910052799 carbon Inorganic materials 0.000 claims description 4
- 229910044991 metal oxide Inorganic materials 0.000 claims description 4
- 150000004706 metal oxides Chemical class 0.000 claims description 4
- 229910052976 metal sulfide Inorganic materials 0.000 claims description 4
- 229910052717 sulfur Inorganic materials 0.000 claims description 4
- 238000004519 manufacturing process Methods 0.000 claims description 3
- 238000002156 mixing Methods 0.000 claims description 3
- 238000004090 dissolution Methods 0.000 claims description 2
- 238000004821 distillation Methods 0.000 claims description 2
- 229910052750 molybdenum Inorganic materials 0.000 claims description 2
- 229910052759 nickel Inorganic materials 0.000 claims description 2
- 238000001228 spectrum Methods 0.000 claims description 2
- DFHAXXVZCFXGOQ-UHFFFAOYSA-K trisodium phosphonoformate Chemical compound [Na+].[Na+].[Na+].[O-]C(=O)P([O-])([O-])=O DFHAXXVZCFXGOQ-UHFFFAOYSA-K 0.000 claims description 2
- 229910052721 tungsten Inorganic materials 0.000 claims description 2
- QGJOPFRUJISHPQ-UHFFFAOYSA-N Carbon disulfide Chemical group S=C=S QGJOPFRUJISHPQ-UHFFFAOYSA-N 0.000 claims 3
- 238000009833 condensation Methods 0.000 claims 1
- 230000005494 condensation Effects 0.000 claims 1
- VGGSQFUCUMXWEO-UHFFFAOYSA-N Ethene Chemical compound C=C VGGSQFUCUMXWEO-UHFFFAOYSA-N 0.000 abstract description 7
- 239000005977 Ethylene Substances 0.000 abstract description 7
- 238000005336 cracking Methods 0.000 abstract description 2
- 239000000243 solution Substances 0.000 description 17
- 229910018072 Al 2 O 3 Inorganic materials 0.000 description 10
- 238000011049 filling Methods 0.000 description 10
- 238000005470 impregnation Methods 0.000 description 10
- 238000011068 loading method Methods 0.000 description 9
- 239000007788 liquid Substances 0.000 description 6
- 230000000694 effects Effects 0.000 description 5
- 239000000126 substance Substances 0.000 description 5
- QGZKDVFQNNGYKY-UHFFFAOYSA-N Ammonia Chemical compound N QGZKDVFQNNGYKY-UHFFFAOYSA-N 0.000 description 4
- 239000008367 deionised water Substances 0.000 description 4
- 229910021641 deionized water Inorganic materials 0.000 description 4
- 230000008021 deposition Effects 0.000 description 4
- 150000005673 monoalkenes Chemical class 0.000 description 4
- XLYOFNOQVPJJNP-UHFFFAOYSA-N water Chemical compound O XLYOFNOQVPJJNP-UHFFFAOYSA-N 0.000 description 4
- -1 VIB metals Chemical class 0.000 description 3
- QGAVSDVURUSLQK-UHFFFAOYSA-N ammonium heptamolybdate Chemical compound N.N.N.N.N.N.O.O.O.O.O.O.O.O.O.O.O.O.O.O.O.O.O.O.O.O.O.O.O.O.[Mo].[Mo].[Mo].[Mo].[Mo].[Mo].[Mo] QGAVSDVURUSLQK-UHFFFAOYSA-N 0.000 description 3
- 238000001354 calcination Methods 0.000 description 3
- UFMZWBIQTDUYBN-UHFFFAOYSA-N cobalt dinitrate Chemical compound [Co+2].[O-][N+]([O-])=O.[O-][N+]([O-])=O UFMZWBIQTDUYBN-UHFFFAOYSA-N 0.000 description 3
- 229910001981 cobalt nitrate Inorganic materials 0.000 description 3
- 238000001816 cooling Methods 0.000 description 3
- 230000003111 delayed effect Effects 0.000 description 3
- 238000005516 engineering process Methods 0.000 description 3
- 229930195733 hydrocarbon Natural products 0.000 description 3
- 150000002430 hydrocarbons Chemical class 0.000 description 3
- 238000002360 preparation method Methods 0.000 description 3
- 230000001105 regulatory effect Effects 0.000 description 3
- VHUUQVKOLVNVRT-UHFFFAOYSA-N Ammonium hydroxide Chemical compound [NH4+].[OH-] VHUUQVKOLVNVRT-UHFFFAOYSA-N 0.000 description 2
- IJGRMHOSHXDMSA-UHFFFAOYSA-N Atomic nitrogen Chemical compound N#N IJGRMHOSHXDMSA-UHFFFAOYSA-N 0.000 description 2
- OKTJSMMVPCPJKN-UHFFFAOYSA-N Carbon Chemical compound [C] OKTJSMMVPCPJKN-UHFFFAOYSA-N 0.000 description 2
- 229910021529 ammonia Inorganic materials 0.000 description 2
- 235000011114 ammonium hydroxide Nutrition 0.000 description 2
- 230000009286 beneficial effect Effects 0.000 description 2
- 230000000052 comparative effect Effects 0.000 description 2
- 230000001276 controlling effect Effects 0.000 description 2
- VLXBWPOEOIIREY-UHFFFAOYSA-N dimethyl diselenide Natural products C[Se][Se]C VLXBWPOEOIIREY-UHFFFAOYSA-N 0.000 description 2
- 150000002431 hydrogen Chemical class 0.000 description 2
- 239000012535 impurity Substances 0.000 description 2
- 230000000977 initiatory effect Effects 0.000 description 2
- 239000012528 membrane Substances 0.000 description 2
- PXHVJJICTQNCMI-UHFFFAOYSA-N nickel Substances [Ni] PXHVJJICTQNCMI-UHFFFAOYSA-N 0.000 description 2
- 229910052757 nitrogen Inorganic materials 0.000 description 2
- 239000011148 porous material Substances 0.000 description 2
- 230000002035 prolonged effect Effects 0.000 description 2
- 238000005486 sulfidation Methods 0.000 description 2
- RHZKOFJYQGKKAO-UHFFFAOYSA-N 3-(fluoromethyl)but-3-enyl phosphono hydrogen phosphate Chemical compound OP(O)(=O)OP(O)(=O)OCCC(=C)CF RHZKOFJYQGKKAO-UHFFFAOYSA-N 0.000 description 1
- QTBSBXVTEAMEQO-UHFFFAOYSA-M Acetate Chemical compound CC([O-])=O QTBSBXVTEAMEQO-UHFFFAOYSA-M 0.000 description 1
- QGZKDVFQNNGYKY-UHFFFAOYSA-O Ammonium Chemical compound [NH4+] QGZKDVFQNNGYKY-UHFFFAOYSA-O 0.000 description 1
- 239000004215 Carbon black (E152) Substances 0.000 description 1
- 229910002651 NO3 Inorganic materials 0.000 description 1
- NHNBFGGVMKEFGY-UHFFFAOYSA-N Nitrate Chemical compound [O-][N+]([O-])=O NHNBFGGVMKEFGY-UHFFFAOYSA-N 0.000 description 1
- 229910019142 PO4 Inorganic materials 0.000 description 1
- VYPSYNLAJGMNEJ-UHFFFAOYSA-N Silicium dioxide Chemical compound O=[Si]=O VYPSYNLAJGMNEJ-UHFFFAOYSA-N 0.000 description 1
- QAOWNCQODCNURD-UHFFFAOYSA-L Sulfate Chemical compound [O-]S([O-])(=O)=O QAOWNCQODCNURD-UHFFFAOYSA-L 0.000 description 1
- NINIDFKCEFEMDL-UHFFFAOYSA-N Sulfur Chemical compound [S] NINIDFKCEFEMDL-UHFFFAOYSA-N 0.000 description 1
- 238000009825 accumulation Methods 0.000 description 1
- 239000003463 adsorbent Substances 0.000 description 1
- 230000002411 adverse Effects 0.000 description 1
- APUPEJJSWDHEBO-UHFFFAOYSA-P ammonium molybdate Chemical compound [NH4+].[NH4+].[O-][Mo]([O-])(=O)=O APUPEJJSWDHEBO-UHFFFAOYSA-P 0.000 description 1
- 229940010552 ammonium molybdate Drugs 0.000 description 1
- 235000018660 ammonium molybdate Nutrition 0.000 description 1
- 239000011609 ammonium molybdate Substances 0.000 description 1
- 150000003863 ammonium salts Chemical class 0.000 description 1
- 239000002518 antifoaming agent Substances 0.000 description 1
- 230000001174 ascending effect Effects 0.000 description 1
- QVGXLLKOCUKJST-UHFFFAOYSA-N atomic oxygen Chemical compound [O] QVGXLLKOCUKJST-UHFFFAOYSA-N 0.000 description 1
- 238000009835 boiling Methods 0.000 description 1
- QGJOPFRUJISHPQ-NJFSPNSNSA-N carbon disulfide-14c Chemical group S=[14C]=S QGJOPFRUJISHPQ-NJFSPNSNSA-N 0.000 description 1
- 238000001833 catalytic reforming Methods 0.000 description 1
- 230000000295 complement effect Effects 0.000 description 1
- 230000009849 deactivation Effects 0.000 description 1
- 230000007547 defect Effects 0.000 description 1
- 238000010586 diagram Methods 0.000 description 1
- 238000009826 distribution Methods 0.000 description 1
- 238000001914 filtration Methods 0.000 description 1
- 238000005187 foaming Methods 0.000 description 1
- 238000010574 gas phase reaction Methods 0.000 description 1
- 238000010438 heat treatment Methods 0.000 description 1
- 239000011261 inert gas Substances 0.000 description 1
- 230000005764 inhibitory process Effects 0.000 description 1
- 239000000203 mixture Substances 0.000 description 1
- KBJMLQFLOWQJNF-UHFFFAOYSA-N nickel(ii) nitrate Chemical compound [Ni+2].[O-][N+]([O-])=O.[O-][N+]([O-])=O KBJMLQFLOWQJNF-UHFFFAOYSA-N 0.000 description 1
- 230000003647 oxidation Effects 0.000 description 1
- 238000007254 oxidation reaction Methods 0.000 description 1
- 229910052760 oxygen Inorganic materials 0.000 description 1
- 239000001301 oxygen Substances 0.000 description 1
- NBIIXXVUZAFLBC-UHFFFAOYSA-K phosphate Chemical compound [O-]P([O-])([O-])=O NBIIXXVUZAFLBC-UHFFFAOYSA-K 0.000 description 1
- 239000010452 phosphate Substances 0.000 description 1
- 230000001012 protector Effects 0.000 description 1
- 239000002516 radical scavenger Substances 0.000 description 1
- 230000002441 reversible effect Effects 0.000 description 1
- 239000012266 salt solution Substances 0.000 description 1
- 238000009738 saturating Methods 0.000 description 1
- 229910052814 silicon oxide Inorganic materials 0.000 description 1
- 239000002210 silicon-based material Substances 0.000 description 1
- 150000003384 small molecules Chemical class 0.000 description 1
- 238000002791 soaking Methods 0.000 description 1
- 238000001179 sorption measurement Methods 0.000 description 1
- 239000011593 sulfur Substances 0.000 description 1
- 238000004227 thermal cracking Methods 0.000 description 1
- 230000009466 transformation Effects 0.000 description 1
Images
Classifications
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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
- C10G67/00—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one process for refining in the absence of hydrogen only
- C10G67/02—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one process for refining in the absence of hydrogen only plural serial stages only
-
- 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
- C10G2300/4006—Temperature
-
- 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
- C10G2300/4012—Pressure
-
- 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
- C10G2300/4018—Spatial velocity, e.g. LHSV, WHSV
-
- 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
- C10G2300/70—Catalyst aspects
-
- Y—GENERAL TAGGING OF NEW TECHNOLOGICAL DEVELOPMENTS; GENERAL TAGGING OF CROSS-SECTIONAL TECHNOLOGIES SPANNING OVER SEVERAL SECTIONS OF THE IPC; TECHNICAL SUBJECTS COVERED BY FORMER USPC CROSS-REFERENCE ART COLLECTIONS [XRACs] AND DIGESTS
- Y02—TECHNOLOGIES OR APPLICATIONS FOR MITIGATION OR ADAPTATION AGAINST CLIMATE CHANGE
- Y02P—CLIMATE CHANGE MITIGATION TECHNOLOGIES IN THE PRODUCTION OR PROCESSING OF GOODS
- Y02P20/00—Technologies relating to chemical industry
- Y02P20/50—Improvements relating to the production of bulk chemicals
- Y02P20/52—Improvements relating to the production of bulk chemicals using catalysts, e.g. selective catalysts
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- Chemical & Material Sciences (AREA)
- Oil, Petroleum & Natural Gas (AREA)
- Engineering & Computer Science (AREA)
- Chemical Kinetics & Catalysis (AREA)
- General Chemical & Material Sciences (AREA)
- Organic Chemistry (AREA)
- Production Of Liquid Hydrocarbon Mixture For Refining Petroleum (AREA)
Abstract
The invention discloses a coking naphtha processing method, which comprises the following steps: (1) coking naphtha feedstock to a diolefin saturation treatment; (2) The reactor is internally provided with a first hydrogenation reaction zone, a second hydrogenation reaction zone and a third hydrogenation reaction zone from top to bottom; hydrogen enters from the second hydrogenation reaction zone and the third hydrogenation reaction zone; the material obtained in the step (1) is separated into a gas phase component and a liquid phase component through flash evaporation, the gas phase component flows upwards to enter a first hydrogenation reaction zone, and the obtained first reaction product is condensed and circulated to a second hydrogenation reaction zone; (3) The liquid phase component enters a second hydrogenation reaction zone and the first reaction product which is condensed and circulated back to carry out hydrogenation reaction, and the obtained second reaction product flows downwards to enter a third hydrogenation reaction zone to carry out liquid phase hydrogenation reaction, so that the product is obtained. According to the invention, through changing the passing path of the raw materials in the reactor and matching the proper catalyst in the corresponding reaction zone, the temperature rise of the catalyst bed in the reactor is effectively controlled, and the raw materials meeting the requirement of downstream ethylene cracking can be produced.
Description
Technical Field
The invention belongs to the fields of oil refining and chemical industry, and particularly relates to a coking naphtha processing method.
Background
The naphtha fraction can be used as ethylene raw material, synthetic ammonia raw material, catalytic reforming raw material, etc. after further hydrofining, and can also be used as blending component of vehicle gasoline. In recent years, with the continuous transformation of the oil refining industry in China to the chemical industry, the ethylene production demand is obviously increased year by year. Ethylene is not only the chemical with the largest yield in the world, but also one of the main raw materials in the petrochemical industry, and the ethylene production capacity has become a standard for the level of the national chemical industry. Therefore, the method continuously widens the naphtha hydrogenation raw material, optimizes the naphtha hydrogenation technology, supports the raw material requirements of the chemical industry in China, and is one of the important points of technical development in the current oil refining field.
Coker naphtha is one of the products of delayed coker units, not only is the sulfur and nitrogen content relatively high, but also the olefin content is high (about 40%) as a result of being a thermal cracking reaction product. Meanwhile, as a silicon-containing antifoaming agent is often added to the delayed coking device for foaming inhibition, the delayed coking device can follow the coked naphtha fraction, so that the coked naphtha contains 5-20 ppm of Si. At present, two reactors are generally connected in series in an industrial coking naphtha hydrogenation device, a pre-protection reactor is used for removing diolefin at a low temperature of about 120 ℃, then impurities such as monoolefin, si, S, N and the like are removed in a main reactor, and the outlet and inlet temperatures of the main reactor are generally 220 ℃ and 360 ℃. The following problems are often encountered in the operation of current industrial devices: (1) unstable heat exchange temperature control problem: before entering the pre-protection reactor, the raw materials are generally subjected to heat exchange and temperature rise in a heat exchanger, wherein the heat exchange flow is that after the heat exchange between the material at the outlet of the pre-protection reactor (about 130 ℃) and the material at the outlet of the main reactor (about 360 ℃), the heat exchange is carried out between the material and the raw materials, so that the raw materials are subjected to heat exchange from normal temperature to the temperature (about 120 ℃) required by the inlet of the pre-protection reactor. If the temperature control is unstable in the two heat exchange processes, the diene in the raw materials is easily coked in the heat exchanger without entering the pre-protection reactor, and the smooth operation of the whole process flow is affected. (2) problem of heat release of mono-olefins in the main reactor: after the diolefin is removed from the coked naphtha at a low temperature in the pre-protection reactor, a large amount of monoolefin is still contained, the olefin releases heat in a large amount in the hydrogenation reaction process of the main reactor, so that the temperature of a catalyst bed is quickly increased, carbon deposition is formed on the surface of the catalyst, if the raw material contains a small amount of oxygen, the condition of coking and carbon deposition of the olefin is more serious, at the moment, the catalyst can only play an active role at a higher temperature, and vicious circle causes quick deactivation of the catalyst and forced shutdown of bed pressure drop. (3) The problem of Si deposition of the main catalyst is that cyclosiloxane contained in coked naphtha is a macromolecular species and is easy to block catalyst pore channels. At present, the industrial device generally adopts a method of filling silicon capturing agent in the first bed layer of the reactor, and the macromolecular silicon-containing compound is converted into small molecules and then adsorbed and removed under the condition of relatively low reaction temperature at the upper part of the reactor. The method for filling the silicon catching agent can relatively reduce the filling amount of the main catalyst in the reactor, which is equivalent to increasing the volume space velocity of the main catalyst under the condition of unchanged treatment amount, and has adverse effect on hydrofining reaction.
CN102051202a discloses a coked naphtha silicon capturing agent, which increases the acidity of the catalyst surface by adding silicon oxide, and improves the silicon capturing effect. However, to substantially avoid Si deposition on the surface of the main catalyst, a higher loading amount of the silicon scavenger is still required, which correspondingly reduces the loading of the main catalyst and affects the refining effect.
CN108431180a discloses a method and system for producing olefins and aromatics from coker naphtha by removing siliceous species from the coker naphtha by adsorption, filtration and/or membrane separation to avoid its impact on the main catalyst, but the manner of physical removal affects the long cycle operation of the unit with saturation of the adsorbent or membrane separation system.
Disclosure of Invention
Aiming at the defects of the prior art, the invention provides a coking naphtha processing method. According to the invention, through setting different reaction areas, the passing path of raw materials in the reactor is changed, and meanwhile, a proper catalyst is matched in the corresponding reaction areas, so that the loading amount of the silicon capturing agent is reduced, the processing amount of the device is improved, the temperature rise of a catalyst bed layer in the reactor is effectively controlled, the running period of the device can be prolonged, and raw materials meeting the requirement of downstream ethylene cracking can be produced.
The coking naphtha processing method of the invention comprises the following steps:
(1) Coking naphtha raw material to carry out diene saturation treatment;
(2) The fixed bed reactor is internally provided with a first hydrogenation reaction zone, a flash evaporation zone, a second hydrogenation reaction zone and a third hydrogenation reaction zone from top to bottom in sequence; hydrogen enters from between the second hydrogenation reaction zone and the third hydrogenation reaction zone, and passes through the second hydrogenation reaction zone, the flash evaporation zone and the first hydrogenation reaction zone in turn upwards; the material obtained in the step (1) enters a flash evaporation zone and is separated into a gas phase component and a liquid phase component through flash evaporation, the gas phase component flows upwards to enter a first hydrogenation reaction zone above the flash evaporation zone for hydrogenation reaction, and the obtained first reaction product is condensed and circulated to a second hydrogenation reaction zone below the flash evaporation zone;
(3) The liquid phase component flows downwards to enter a second hydrogenation reaction zone and is subjected to hydrogenation reaction together with the first reaction product which is condensed and circulated, so that a second reaction product is obtained;
(4) The second reaction product flows downwards and mixes the hydrogen with the dissolution part flowing upwards, and enters a third hydrogenation reaction zone below the second hydrogenation reaction zone to carry out liquid phase hydrogenation reaction, and the obtained third reaction product flows out as a product.
In the method, the distillation range of the coked naphtha raw material is generally 30-230 ℃, S is not more than 5000 mug/g, N is not more than 300 mug/g, and olefin is not more than 50%.
In the method, coked naphtha raw material enters a pretreatment reactor for diene saturation treatment after heat exchange; the heat exchange of the coked naphtha raw material refers to heat exchange between the raw material at normal temperature (10-30 ℃) and liquid phase components (generally about 150-200 ℃) flowing out of the bottom of the fixed bed reactor, and the raw material after heat exchange can meet the feeding requirement of the pretreatment reactor (generally about 100-140 ℃). The hydrogen is fed under normal temperature, and plays roles of providing hydrogen needed by reaction and controlling the temperature rise of the catalyst bed.
In the process of the present invention, the pretreatment reactor is mainly used for saturating the diolefins in the feedstock at low temperature. The pretreatment reactor is generally filled with a catalyst with a diene hydrogenation saturation function, such as SHT-1 catalyst developed by FRIPP, and the reaction conditions in the pretreatment reactor are as follows: the temperature is 120-150 ℃, and the pressure is 2.0-5.0 MPa.
In the method, the flash evaporation zone is used for separating a fraction below 140-150 ℃ from the raw material, the fraction enters the first hydrogenation reaction zone as a gas phase component, and the fraction above 150 ℃ enters the second reaction zone as a liquid phase component. The operation condition of the flash evaporation zone is that the pressure is 1.0-6.0 MPa, preferably 2.0-4.0 MPa, wherein the hydrogen partial pressure accounts for 45-80% of the total pressure proportion; the feeding temperature is 100-200 ℃, preferably 130-180 ℃; hydrogen oil volume ratio 10: 1-800: 1, preferably 100: 1-400: 1.
in the method, the first hydrogenation reaction zone is used for carrying out olefin removal, desulfurization and denitrification reactions on gas phase components with the temperature of 140-150 ℃, and a semi-vulcanized hydrogenation catalyst is filled in the first hydrogenation reaction zone. The catalyst loading volume ratio in the first hydrogenation reaction zone is 1% -60%, preferably 5% -40% based on the total catalyst loading in the reactor.
The operating conditions of the first hydrogenation reaction zone are generally: the pressure is 1.0-8.0 MPa, preferably 2.0-6.0 MPa, wherein the hydrogen partial pressure accounts for 50-90% of the total pressure; volume space velocity is 0.1-10.0 h -1 Preferably 0.5 to 6.0 hours -1 The method comprises the steps of carrying out a first treatment on the surface of the The reaction temperature is 80-180 ℃, preferably 100-150 ℃; hydrogen oil volume ratio 10: 1-400: 1, preferably 100: 1-200: 1.
in the method, the second hydrogenation reaction zone is used for carrying out olefin removal, desulfurization and denitrification reactions on the liquid phase component subjected to flash evaporation and the first reaction product condensed and circulated from the first hydrogenation reaction zone. The second hydrogenation reaction zone is graded and filled with a silicon capturing agent and a semi-vulcanized hydrogenation catalyst, wherein the silicon capturing agent can be commercial silicon capturing agent such as FHRS series silicon capturing agent developed by FRIPP or can be prepared according to the prior art. The catalyst loading volume ratio in the second hydrogenation reaction zone is 1-80%, preferably 10-50% based on the total catalyst loading in the reactor. Wherein, the volume ratio of the silicon capturing agent to the semi-vulcanized hydrogenation catalyst is 5: 1-1: 5. the second hydrogenation reaction zone generally operates under the following conditions: the pressure is 1.0-8.0 MPa, preferably 2.0-6.0 MPa, wherein the hydrogen partial pressure accounts for 40-70% of the total pressure; volume space velocity is 0.1-10.0 h -1 Preferably 0.5 to 6.0 hours -1 The method comprises the steps of carrying out a first treatment on the surface of the The reaction temperature is 80-180 ℃, preferably 100-150 ℃; hydrogen oil volume ratio 10: 1-400: 1 preferably 100: 1-200: 1.
in the method of the invention, the third hydrogenation reaction zone is used for mixing and dissolving part of the second reaction product of hydrogen to carry out the complementary refining reaction, and the third hydrogenation reaction zone is filled with the semi-vulcanized catalyst. The catalyst loading volume ratio of the third hydrogenation reaction zone is 1-80%, preferably 20-50% based on the total catalyst loading in the reactor. Since the third hydrogenation reaction zone is subjected to a liquid phase reaction, the sulfidation degree of the semi-sulfidation catalyst of the third reaction zone (+4 valence state of the group VIB metal to the molar ratio of the total amount of the group VIB metals) is preferably 80% or more. The operating conditions of the third hydrogenation reaction zone are generally as follows: the pressure is 1.0-10.0 MPa, preferably 3.0-8.0 MPa, which is a pure liquid phase reaction zone, and the volume ratio of standard hydrogen to oil is 2-50, preferably 10-30; volume space velocity is 0.1-8.0 h -1 Preferably 0.5 to 6.0 hours -1 The method comprises the steps of carrying out a first treatment on the surface of the The reaction temperature is 80-180 ℃, preferably 100-150 ℃; hydrogen oil volume ratio 10: 1-600: 1, preferably 100: 1-300: 1.
in the method of the present invention, it is further preferred that the molar ratio of "+4 valence state group VIB metals to the total amount of group VIB metals of the semi-sulfided hydrogenation catalyst packed in the first hydrogenation reaction zone, the second hydrogenation reaction zone and the third hydrogenation reaction zone" increases in sequence.
In the method of the invention, the semi-vulcanized hydrogenation catalyst comprises a VIB group metal sulfide, a VIII group metal oxide and an alumina carrier; based on the total weight of the catalyst, the content of the VIB group metal sulfide is 5-20wt%, the content of the VIII group metal oxide is 3-10wt%, and the balance is an alumina carrier; wherein the group VIB metal is selected from at least one of Mo and W, and the group VIII metal is selected from at least one of Co and Ni; wherein the molar ratio of the +4 valent VIB group metal to the total VIB group metal is 60-90% when the hydrogenation catalyst (oxidation state) is analyzed by XPS energy spectrum.
The preparation method of the semi-vulcanized hydrogenation catalyst comprises the following steps:
(1) Impregnating an alumina carrier with an impregnating solution containing a VIB group metal, then drying, and vulcanizing the dried material;
(2) Impregnating the catalyst vulcanized in the step (1) with an impregnating solution containing a metal of the VIII group, and then drying and roasting the catalyst in an inert atmosphere to obtain the semi-vulcanized hydrogenation catalyst.
The preparation method of the impregnating solution of the VIB metal in the step (1) is well known to those skilled in the art, for example, phosphate or ammonium salt solution is generally adopted, the mass concentration of the impregnating solution is 0.1 g/mL-2.4 g/mL, and an equal volume impregnation mode can be adopted. The group VIB metal is preferably Mo and/or W.
The drying conditions in the step (1) are as follows: the drying temperature is 100-220 ℃ and the drying time is 2-6 hours.
The vulcanization treatment in the step (1) is well known to those skilled in the art, and dry vulcanization or wet vulcanization is generally adopted, wherein the dry vulcanizing agent is hydrogen sulfide, and the wet vulcanizing agent is carbon disulfide, dimethyl disulfide and the like; the vulcanization pressure is 3.5-6.5MPa, the vulcanization temperature is 220-380 ℃, and the vulcanization time is 3-10h.
The preparation method of the impregnation liquid of the group VIII metal in the step (2) is well known to those skilled in the art, for example, solutions such as acetate, nitrate, sulfate and the like are generally adopted, the mass concentration of the impregnation liquid is 0.14g/mL to 1.4g/mL, and an isovolumetric impregnation mode can be adopted, wherein the group VIII metal is Ni and/or Co.
The inert atmosphere in the step (2) is N 2 One or more of inert gases; the drying temperature is 40-120 ℃ and the drying time is 2-12 hours; the roasting temperature is 240-550 ℃ and the roasting time is 3-8 hours.
The invention also provides a coker naphtha processing system, which comprises a pretreatment reactor and a fixed bed hydrogenation reactor which are connected in series; wherein the fixed bed hydrogenation reactor is internally provided with a first hydrogenation reaction zone, a flash evaporation zone, a second hydrogenation reaction zone and a third hydrogenation reaction zone from top to bottom in sequence; the flash evaporation zone is provided with a liquid-phase material inlet, and a hydrogen inlet is arranged between the second hydrogenation reaction zone and the third hydrogenation reaction zone; the bottom of the fixed bed hydrogenation reactor is provided with a product outlet.
In the system, heat exchange equipment is arranged between a coking naphtha raw material inlet pipeline and a product outlet pipeline arranged at the bottom of the fixed bed hydrogenation reactor and used for exchanging heat between a product and a raw material, and the raw material after heat exchange enters the pretreatment reactor.
Compared with the prior art, the invention has the following advantages:
(1) By reducing the temperature of the reaction raw material inlet and matching with flash evaporation, the material rich in mono-olefins is ingeniously divided into two components of gas and liquid, and flows upwards and downwards respectively, so that the concentration of olefins in the material entering the first hydrogenation reaction zone and the second hydrogenation reaction zone is obviously reduced. The refined components in the first hydrogenation reaction zone are condensed and then recycled to the second hydrogenation reaction zone, so that the olefin concentration of the materials in the second hydrogenation reaction zone can be further diluted, and the temperature rise of a bed layer can be effectively controlled. Meanwhile, by combining a semi-vulcanized catalyst with good low-temperature activity and stable initial activity, the feeding temperature (lower than 200 ℃) at a reaction inlet is greatly reduced compared with that of the conventional technology (about 220 ℃) at the inlet, the liquid-phase material subjected to flash evaporation can be kept to react in a liquid-phase state, and the concentrated and severe heat release of olefin caused by the excessive initial activity is avoided, so that the temperature of a catalyst bed is difficult to control.
(2) The normal-temperature hydrogen enters from the lower part of the second hydrogenation reaction zone, can exchange heat with the liquid phase material flowing into the second hydrogenation reaction zone, can carry heat generated by the reaction to the first hydrogenation reaction zone and then flows out from the top of the reactor, and plays a role in integrally controlling the temperature rise of the catalyst bed of the reactor. The invention can realize that the temperature rise of the whole reaction is controlled within 50 ℃, and compared with the temperature rise of about 150 ℃ at the inlet and outlet of the conventional coking naphtha hydrofining reactor, the temperature rise is obviously reduced.
(3) The flash evaporation separation temperature of the raw material inlet is controlled below 150 ℃, so that silicon-containing species-cyclosiloxane (boiling point is above 150 ℃) in the raw material does not enter the first hydrogenation reaction zone along with gas phase, and the silicon capturing agent does not need to be filled in the first hydrogenation reaction zone. In addition, because the temperature of the whole catalyst bed can be effectively controlled, the silicon-containing species descending along with the liquid phase can not reach the highest reaction rate, and the catalyst pore channels can not be blocked after hydrogenation reaction occurs in a large amount, so that the second hydrogenation reaction zone can also reduce the filling amount of the silicon capturing agent by about 50 percent compared with a conventional naphtha hydrogenation device.
(4) Because the impurity content in the material entering the third hydrogenation reaction zone is obviously reduced, the effect of deep supplementary refining can be achieved and no large temperature rise occurs. Because low-temperature hydrogen enters from the second hydrogenation reaction zone and the third hydrogenation reaction zone, the hydrogen partial pressure at the bottom of the reactor can be ensured to be relatively high, which is beneficial to the occurrence of deep hydrogenation reaction. Meanwhile, the generated gas-phase light hydrocarbons flow out from the first hydrogenation reaction zone, and the liquid phase entering the reaction zone III hardly contains low-molecular hydrocarbons and hydrogen sulfide, so that the solubility of hydrogen in the liquid phase is increased, and the hydrogenation reaction is further promoted. The ascending hydrogen plays roles of balancing the temperature distribution in the whole reactor and stripping hydrogen sulfide, so that the accumulation of hydrogen sulfide in the reactor can be avoided, the hydrogen sulfide is carried to the surfaces of semi-sulfided catalysts filled in the first hydrogenation reaction zone and the second hydrogenation reaction zone, the catalysts are gradually activated, and long-period operation is ensured. The invention realizes the efficient and stable operation of the coking naphtha hydrogenation device through the adjustment of the process and the cooperation of the catalyst.
Drawings
FIG. 1 is a schematic diagram of a coker naphtha processing method of the present invention.
In the figure: 1-raw material, 2-heat exchanger, 3-pretreatment reactor, 4-fixed bed hydrogenation reactor (main reactor), 5-flash evaporation zone, 6-hydrogen, 7-first reaction product, 8-first hydrogenation reaction zone, 9-second hydrogenation reaction zone, 10-third hydrogenation reaction zone, 11-second hydrogenation reaction zone inlet, 12-refined naphtha product and 13-condenser.
Detailed Description
The invention will now be described in more detail with reference to the accompanying drawings and examples, which are not intended to limit the invention thereto.
Taking the figure 1 as an example, the implementation process of the coking naphtha processing method of the invention comprises the following steps: the reaction raw material 1 is subjected to heat exchange with a refined naphtha product 12 at the bottom of the main reactor through a heat exchanger 2, enters a pretreatment reactor 3 to remove the diene in the reaction raw material, and then enters a main reactor 4. In the flash zone 5 the gas-liquid separation is into a gas phase and a liquid phase. The gas phase flows upward into the first hydrogenation reaction zone 8 and the liquid phase flows downward into the second hydrogenation reaction zone 9. The hydrogen 6 enters the main reactor 4 between the second hydrogenation reaction zone 9 and the third hydrogenation reaction zone 10, and after being mixed and contacted with the liquid phase material flowing downwards in the second hydrogenation reaction zone 9, the excessive hydrogen continues to flow upwards to enter the second hydrogenation reaction zone 9, and the liquid phase material dissolved and carrying the hydrogen flows downwards to enter the third hydrogenation reaction zone 10.
The first hydrogenation reaction zone 8 is subjected to gas phase reaction, and mainly subjected to hydrodesulfurization, denitrification and olefin saturation reactions of fractions below 150 ℃ in coker naphtha to generate a first reaction product 7, and the first reaction product is recycled to the inlet 11 of the second hydrogenation reaction zone through a condenser 13. In the second hydrogenation reaction zone 9, gas-liquid two-phase reaction occurs, the liquid phase is the fraction above 150 ℃ in the coked naphtha flows downwards, the gas phase is the hydrogen gas flows upwards, and the gas-liquid reverse contact occurs desilication, desulfurization, denitrification and olefin saturation reactions. The hydrogen sulfide and low molecular hydrocarbon generated by the reaction flow upwards along with the gas phase material flow into the second hydrogenation reaction zone 9 and the first hydrogenation reaction zone 8, and finally flow out of the device from the top of the reactor. The hydrogenated liquid phase flow flows downwards to enter a third hydrogenation reaction zone 10, the third hydrogenation reaction zone 10 is in liquid phase reaction, and liquid phase components and dissolved hydrogen undergo deep supplementary refining reaction to generate refined naphtha products 12.
Example 1
Ammonium heptamolybdate is regulated by ammonia water until the ammonium heptamolybdate is completely dissolved, so as to prepare impregnating solution with the concentration of 0.18g/mL, then the impregnating solution is impregnated into an alumina carrier by an isovolumetric impregnation method, and the alumina carrier is dried for 4 hours at 140 ℃. Then adopting DMDS to carry out vulcanization treatment, wherein the vulcanization pressure is 4.0MPa, the vulcanization temperature is 320 ℃, the vulcanization time is 4 hours, and then the vulcanization is carried out under N 2 Cooling to room temperature in the atmosphere to obtain MoS 2 /Al 2 O 3 。
Dissolving cobalt nitrate in deionized water to prepare an impregnating solution with the concentration of 0.15g/mL, and impregnating the solution into MoS by adopting an isovolumetric impregnation method 2 /Al 2 O 3 On, then at N 2 Drying at 60deg.C for 4 hr, and calcining at 320 deg.C for 3 hr to obtain MoS 2 -CoO/Al 2 O 3 Catalyst M1.
Example 2
Regulating ammonium heptamolybdate to be completely dissolved by ammonia waterPreparing an impregnating solution with the concentration of 0.18g/mL, impregnating the alumina carrier by adopting an isovolumetric impregnation method, and drying the alumina carrier for 4 hours at 140 ℃. Then use a solution containing 1.8% H 2 S, hydrogen is vulcanized, the vulcanization temperature is 330 ℃, the vulcanization pressure is 5.0MPa, the vulcanization time is 4 hours, and then the vulcanization is carried out on N 2 Cooling to room temperature in the atmosphere to obtain MoS 2 /Al 2 O 3 。
Dissolving nickel nitrate in deionized water to prepare an impregnating solution with the concentration of 0.16g/mL, and impregnating the solution into MoS by adopting an isovolumetric impregnation method 2 /Al 2 O 3 On, then at N 2 Drying at 80deg.C for 4 hr, and calcining at 300deg.C for 3 hr to obtain MoS 2 -NiO/Al 2 O 3 Catalyst M2.
Example 3
Ammonium metatungstate is regulated to be completely dissolved by ammonia water to prepare impregnating solution with the concentration of 0.18g/mL, and then the impregnating solution is impregnated into an alumina carrier by an isovolumetric impregnation method, and is dried for 4 hours at 140 ℃. Then adopting DMDS to carry out vulcanization treatment, wherein the vulcanization temperature is 340 ℃, the vulcanization pressure is 5.0MPa, the vulcanization time is 4 hours, and then the vulcanization is carried out under N 2 Cooling to room temperature in the atmosphere to obtain MoS 2 /Al 2 O 3 。
Dissolving cobalt nitrate in deionized water to prepare an impregnating solution with the concentration of 0.16g/mL, and impregnating the solution into MoS by adopting an isovolumetric impregnation method 2 /Al 2 O 3 On, then at N 2 Drying at 90deg.C for 4 hr, and calcining at 300deg.C for 3 hr to obtain MoS 2 -CoO/Al 2 O 3 Catalyst M3.
Comparative example 1
Dissolving ammonium molybdate and cobalt nitrate in deionized water, soaking the mixture in an equal volume to an alumina carrier, drying the alumina carrier at 140 ℃ for 3 hours, and roasting the alumina carrier at 380 ℃ for 4 hours to obtain MoO 3 -CoO/Al 2 O 3 Catalyst D1.
Examples 4 to 6
The hydrogenation reaction system of the embodiment adopts a pretreatment reactor and a main reactor to be connected in series, and the raw materials and the effluent at the bottom of the main reactor enter the pretreatment reactor after heat exchange. The pre-protector is a 100mL pilot plant, and 50mL of catalyst is filled; the main reactor was a 100mL pilot plant. The pretreatment reactor is provided with a bed layer filled with SHT-1 catalyst developed by FRIPP. The first, second and third hydrogenation reaction areas of the main reactor are respectively provided with a catalyst bed layer. Filling a semi-vulcanized catalyst M1 in a first hydrogenation reaction zone, and filling a silicon capturing agent and a semi-vulcanized catalyst M2 in a second hydrogenation reaction zone in a grading manner, wherein the silicon capturing agent is FHRS-2 silicon capturing agent developed by FRIPP, and the filling volume ratio of the silicon capturing agent to the semi-vulcanized catalyst is 1:2. the third hydrogenation reaction zone is filled with a semi-sulfided catalyst M3. The catalyst volume filling ratio of the first hydrogenation reaction zone, the second hydrogenation reaction zone and the third hydrogenation reaction zone is 3:3:4. coked naphtha is used as a raw material, the catalyst properties are shown in table 1, the raw material oil properties are shown in table 2, and the reaction process conditions and results are shown in table 3.
Comparative example 2
The pretreatment reactor is connected with the main reactor in series by adopting a conventional coking naphtha hydrogenation system, and is a 100mL fixed bed hydrogenation reactor. The raw materials enter a pretreatment reactor after being heated by a heating furnace. The pretreatment reactor was equipped with a bed packed with a SHT-1 catalyst developed by 50mL FIPP. The main reactor is provided with three beds corresponding to the first, second and third hydrogenation reaction areas respectively. The first hydrogenation reaction zone is filled with FHRS-2 silicon capture agent developed by FRIPP, the second and third hydrogenation reaction zones are filled with conventional hydrofining catalyst D1, the total catalyst filling volume of the main reactor is the same as that of the embodiment, and the catalyst filling volume ratio of the three reaction zones is 3:4:3. coked naphtha is used as a raw material, the catalyst properties are shown in table 1, the raw material oil properties are shown in table 2, and the reaction process conditions and results are shown in table 3.
TABLE 1 catalyst physicochemical Properties
TABLE 2 oil Properties of raw materials
TABLE 3 hydrogenation process conditions and results
As can be seen from Table 3, compared with the conventional fixed bed hydrogenation process, the method can use coked naphtha as a raw material to produce qualified ethylene raw material, and the reaction temperature rise is reduced by nearly 100 ℃ compared with the conventional fixed bed technology, which is very beneficial to the long-period stable operation of the device. In addition, the refined oil still has certain Si content, which indicates that Si in the raw material is not deposited on the surface of the main catalyst in a large amount, and the service life of the catalyst is also prolonged.
Claims (19)
1. A coker naphtha processing method characterized by comprising the following steps: (1) Carrying out a pretreatment reaction of diene saturation on coker naphtha raw material; (2) The fixed bed reactor is internally provided with a first hydrogenation reaction zone, a flash evaporation zone, a second hydrogenation reaction zone and a third hydrogenation reaction zone from top to bottom in sequence; hydrogen enters from between the second hydrogenation reaction zone and the third hydrogenation reaction zone, and passes through the second hydrogenation reaction zone, the flash evaporation zone and the first hydrogenation reaction zone in turn upwards; the material obtained in the step (1) enters a flash evaporation zone at a feeding temperature lower than 200 ℃ and is separated into a gas phase component and a liquid phase component through flash evaporation, the flash evaporation separation temperature is controlled below 150 ℃, and the gas phase component flows upwards to enter a first hydrogenation reaction zone above the flash evaporation zone to carry out hydrodeolefine, desulfuration and denitrification reactions, so as to obtain a first reaction product; (3) The liquid phase component and the first reaction product recycled by condensation flow downwards to enter a second hydrogenation reaction zone for hydrodeolefination, desulfurization and denitrification reaction, and a second reaction product is obtained; (4) The second reaction product flows downwards and mixes the hydrogen with the dissolution part flowing upwards, and enters a third hydrogenation reaction zone below the second hydrogenation reaction zone to carry out liquid phase hydrogenation reaction, and the obtained third reaction product flows out as a product; the first hydrogenation reaction zone, the second hydrogenation reaction zone and the third hydrogenation reaction zone are filled with a semi-vulcanized hydrogenation catalyst at the same time, and when the catalyst is analyzed by XPS energy spectrum, the +4 valence state VIB metal accounts for 60% -90% of the total mole ratio of the VIB metal; the second hydrogenation reaction zone is graded filled with a silicon catching agent and a semi-vulcanized hydrogenation catalyst;
the molar ratio of +4 valent VIB group metal in the semi-vulcanized catalyst in the third hydrogenation reaction zone to the total VIB group metal is more than 80%;
the molar ratio of the +4 valence state of the VIB metal of the semi-vulcanized hydrogenation catalyst filled in the first hydrogenation reaction zone, the second hydrogenation reaction zone and the third hydrogenation reaction zone to the total amount of the VIB metal is sequentially increased;
the semi-vulcanized hydrogenation catalyst comprises a VIB group metal sulfide, a VIII group metal oxide and an alumina carrier; based on the total weight of the catalyst, the content of the VIB group metal sulfide is 5-20wt%, the content of the VIII group metal oxide is 3-10wt%, and the balance is an alumina carrier; wherein the group VIB metal is selected from at least one of Mo and W, and the group VIII metal is selected from at least one of Co and Ni;
a process for preparing a semi-sulfided hydrogenation catalyst comprising the following: impregnating an alumina carrier with an impregnating solution containing a VIB group metal, then carrying out primary drying, and carrying out vulcanization treatment on the dried material; impregnating the catalyst after vulcanization with an impregnating solution containing a group VIII metal, and then drying and roasting for the second time in an inert atmosphere to obtain the semi-vulcanized hydrogenation catalyst.
2. The method according to claim 1, characterized in that: the distillation range of the coked naphtha raw material is 30-230 ℃, S is no more than 5000 mug/g, N is no more than 300 mug/g, and olefin is no more than 50%.
3. The method according to claim 1, characterized in that: the pretreatment reaction conditions are as follows: the temperature is 120-150 ℃ and the pressure is 2.0-5.0 MPa.
4. The method according to claim 1, characterized in that: the flash evaporation zone is used for separating a fraction below 140-150 ℃ from the raw material, and the fraction enters the first hydrogenation reaction zone as a gas phase component, and the fraction above 150 ℃ enters the second reaction zone as a liquid phase component; the flash zone operating conditions were: the pressure is 1.0-6.0 MPa, wherein the hydrogen partial pressure accounts for 45-80% of the total pressure proportion; the feeding temperature is 100-200 ℃; hydrogen oil volume ratio 10: 1-800: 1.
5. the method according to claim 1, characterized in that: the flash zone operating conditions were: the pressure is 2.0-4.0 MPa, wherein the hydrogen partial pressure accounts for 45-80% of the total pressure proportion; the feeding temperature is 130-180 ℃; hydrogen oil volume ratio 100:1 to 400:1.
6. the method according to claim 1, characterized in that: the first hydrogenation reaction zone is used for carrying out olefin removal, desulfurization and denitrification on gas phase components with the temperature of 140-150 ℃; the operating conditions of the first hydrogenation reaction zone are as follows: the pressure is 1.0-8.0 MPa, wherein the hydrogen partial pressure accounts for 50-90% of the total pressure proportion; the volume airspeed is 0.1-10.0 h < -1 >; the reaction temperature is 80-180 ℃; hydrogen oil volume ratio 10:1 to 400:1.
7. the method according to claim 6, wherein: the operating conditions of the first hydrogenation reaction zone are as follows: the pressure is 2.0-6.0 MPa, wherein the hydrogen partial pressure accounts for 50-90% of the total pressure proportion; the volume airspeed is 0.5-6.0 h < -1 >; the reaction temperature is 100-150 ℃; hydrogen oil volume ratio 100:1 to 200:1.
8. the method according to claim 1, characterized in that: the second hydrogenation reaction zone is used for carrying out olefin removal, desulfurization and denitrification reactions on the liquid phase component subjected to flash evaporation and the first reaction product condensed and circulated from the first hydrogenation reaction zone; the volume ratio of the silicon capturing agent to the semi-vulcanized hydrogenation catalyst in the second hydrogenation reaction zone is 5:1 to 1:5.
9. the method according to claim 8, wherein: the operating conditions of the second hydrogenation reaction zone are as follows: the pressure is 1.0-8.0 MPa, wherein the hydrogen partial pressure accounts for 40-70% of the total pressure proportion; the volume airspeed is 0.1-10.0 h < -1 >; the reaction temperature is 80-180 ℃; hydrogen oil volume ratio 10:1 to 400:1.
10. the method according to claim 9, wherein: the operating conditions of the second hydrogenation reaction zone are as follows: the pressure is 2.0-6.0 MPa, wherein the hydrogen partial pressure accounts for 40-70% of the total pressure proportion; the volume airspeed is 0.5-6.0 h < -1 >; the reaction temperature is 100-150 ℃; hydrogen oil volume ratio 100:1 to 200:1.
11. the method according to claim 1, characterized in that: the third hydrogenation reaction zone is used for mixing and dissolving part of the second reaction product of hydrogen to carry out supplementary refining reaction; the operating conditions of the third hydrogenation reaction zone are as follows: the pressure is 1.0-10.0 MPa, which is a pure liquid phase reaction zone, and the volume ratio of standard hydrogen to oil is 2-50; the volume airspeed is 0.1-8.0 h < -1 >; the reaction temperature is 80-180 ℃; hydrogen oil volume ratio 10:1 to 600:1.
12. the method according to claim 11, wherein: the operating conditions of the third hydrogenation reaction zone are as follows: the pressure is 3.0-8.0 MPa; the volume ratio of the standard hydrogen to the oil is 10-30; the volume airspeed is 0.5-6.0 h < -1 >; the reaction temperature is 100-150 ℃; hydrogen oil volume ratio 100:1 to 300:1.
13. the method according to claim 1, characterized in that: the mass concentration of the impregnating solution of the VIB group metal for preparing the semi-vulcanized hydrogenation catalyst is 0.1 g/mL-2.4 g/mL.
14. The method according to claim 1, characterized in that: the first drying conditions for preparing the semi-sulfided hydrogenation catalyst are: the drying temperature is 100-220 ℃ and the drying time is 2-6 hours.
15. The method according to claim 1, characterized in that: the vulcanization treatment for preparing the semi-vulcanized hydrogenation catalyst adopts dry vulcanization or wet vulcanization, wherein the dry vulcanizing agent is hydrogen sulfide, and the wet vulcanizing agent is carbon disulfide or dimethyl disulfide; the vulcanization pressure is 3.5-6.5MPa, the vulcanization temperature is 220-380 ℃, and the vulcanization time is 3-10h.
16. The method according to claim 1, characterized in that: the mass concentration of the impregnating solution of the VIII group metal for preparing the semi-vulcanized hydrogenation catalyst is 0.14 g/mL-1.4 g/mL.
17. The method according to claim 1, characterized in that: preparing a semi-vulcanized hydrogenation catalyst, wherein the inert atmosphere is N2; the second drying temperature is 40-120 ℃ and the drying time is 2-12 hours; the roasting temperature is 240-550 ℃ and the roasting time is 3-8 hours.
18. A coker naphtha processing system implementing the method of any one of claims 1-17, wherein: the system comprises a pretreatment reactor and a fixed bed hydrogenation reactor which are connected in series; wherein the fixed bed hydrogenation reactor is internally provided with a first hydrogenation reaction zone, a flash evaporation zone, a second hydrogenation reaction zone and a third hydrogenation reaction zone from top to bottom in sequence; the flash evaporation zone is provided with a liquid-phase material inlet, and a hydrogen inlet is arranged between the second hydrogenation reaction zone and the third hydrogenation reaction zone; the bottom of the fixed bed hydrogenation reactor is provided with a product outlet.
19. The processing system of claim 18, wherein: a heat exchange device is arranged between a coking naphtha raw material inlet pipeline and a product outlet pipeline arranged at the bottom of the fixed bed hydrogenation reactor and is used for heat exchange between a product and a raw material, and the raw material after heat exchange enters the pretreatment reactor.
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| CN108014781A (en) * | 2016-10-31 | 2018-05-11 | 中国石油化工股份有限公司 | A kind of hydrogenation catalyst and its preparation method and application |
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| CN108014781A (en) * | 2016-10-31 | 2018-05-11 | 中国石油化工股份有限公司 | A kind of hydrogenation catalyst and its preparation method and application |
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Effective date of registration: 20231129 Address after: 100728 No. 22 North Main Street, Chaoyang District, Beijing, Chaoyangmen Patentee after: CHINA PETROLEUM & CHEMICAL Corp. Patentee after: Sinopec (Dalian) Petrochemical Research Institute Co.,Ltd. Address before: 100728 No. 22 North Main Street, Chaoyang District, Beijing, Chaoyangmen Patentee before: CHINA PETROLEUM & CHEMICAL Corp. Patentee before: DALIAN RESEARCH INSTITUTE OF PETROLEUM AND PETROCHEMICALS, SINOPEC Corp. |