CN110437867B - Method for producing high value-added product by using liquefied gas - Google Patents
Method for producing high value-added product by using liquefied gas Download PDFInfo
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- CN110437867B CN110437867B CN201810415637.5A CN201810415637A CN110437867B CN 110437867 B CN110437867 B CN 110437867B CN 201810415637 A CN201810415637 A CN 201810415637A CN 110437867 B CN110437867 B CN 110437867B
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- 238000004519 manufacturing process Methods 0.000 title claims abstract description 37
- 239000007789 gas Substances 0.000 claims abstract description 202
- 238000005899 aromatization reaction Methods 0.000 claims abstract description 98
- 239000000047 product Substances 0.000 claims abstract description 86
- 238000005804 alkylation reaction Methods 0.000 claims abstract description 53
- 230000003197 catalytic effect Effects 0.000 claims abstract description 49
- 238000000034 method Methods 0.000 claims abstract description 46
- 230000029936 alkylation Effects 0.000 claims abstract description 40
- 238000006356 dehydrogenation reaction Methods 0.000 claims abstract description 40
- 239000002994 raw material Substances 0.000 claims abstract description 40
- 238000006266 etherification reaction Methods 0.000 claims abstract description 38
- BZLVMXJERCGZMT-UHFFFAOYSA-N Methyl tert-butyl ether Chemical compound COC(C)(C)C BZLVMXJERCGZMT-UHFFFAOYSA-N 0.000 claims abstract description 35
- UFHFLCQGNIYNRP-UHFFFAOYSA-N Hydrogen Chemical compound [H][H] UFHFLCQGNIYNRP-UHFFFAOYSA-N 0.000 claims abstract description 26
- 239000001257 hydrogen Substances 0.000 claims abstract description 25
- 229910052739 hydrogen Inorganic materials 0.000 claims abstract description 25
- QQONPFPTGQHPMA-UHFFFAOYSA-N propylene Natural products CC=C QQONPFPTGQHPMA-UHFFFAOYSA-N 0.000 claims abstract description 24
- 125000004805 propylene group Chemical group [H]C([H])([H])C([H])([*:1])C([H])([H])[*:2] 0.000 claims abstract description 24
- 239000006227 byproduct Substances 0.000 claims abstract description 10
- 239000000463 material Substances 0.000 claims description 122
- OKKJLVBELUTLKV-UHFFFAOYSA-N Methanol Chemical compound OC OKKJLVBELUTLKV-UHFFFAOYSA-N 0.000 claims description 121
- ATUOYWHBWRKTHZ-UHFFFAOYSA-N Propane Chemical compound CCC ATUOYWHBWRKTHZ-UHFFFAOYSA-N 0.000 claims description 92
- 239000003054 catalyst Substances 0.000 claims description 78
- 229910052799 carbon Inorganic materials 0.000 claims description 57
- 238000006243 chemical reaction Methods 0.000 claims description 54
- 239000000203 mixture Substances 0.000 claims description 54
- OKTJSMMVPCPJKN-UHFFFAOYSA-N Carbon Chemical compound [C] OKTJSMMVPCPJKN-UHFFFAOYSA-N 0.000 claims description 52
- NNPPMTNAJDCUHE-UHFFFAOYSA-N isobutane Chemical compound CC(C)C NNPPMTNAJDCUHE-UHFFFAOYSA-N 0.000 claims description 52
- 239000001294 propane Substances 0.000 claims description 48
- 238000000926 separation method Methods 0.000 claims description 39
- IJDNQMDRQITEOD-UHFFFAOYSA-N n-butane Chemical compound CCCC IJDNQMDRQITEOD-UHFFFAOYSA-N 0.000 claims description 36
- 239000004215 Carbon black (E152) Substances 0.000 claims description 33
- VQTUBCCKSQIDNK-UHFFFAOYSA-N Isobutene Chemical compound CC(C)=C VQTUBCCKSQIDNK-UHFFFAOYSA-N 0.000 claims description 29
- 229930195733 hydrocarbon Natural products 0.000 claims description 28
- 239000001282 iso-butane Substances 0.000 claims description 26
- XLYOFNOQVPJJNP-UHFFFAOYSA-N water Substances O XLYOFNOQVPJJNP-UHFFFAOYSA-N 0.000 claims description 23
- 150000002430 hydrocarbons Chemical class 0.000 claims description 21
- 238000004821 distillation Methods 0.000 claims description 19
- 238000003808 methanol extraction Methods 0.000 claims description 18
- 238000011084 recovery Methods 0.000 claims description 16
- 239000001273 butane Substances 0.000 claims description 14
- OFBQJSOFQDEBGM-UHFFFAOYSA-N n-pentane Natural products CCCCC OFBQJSOFQDEBGM-UHFFFAOYSA-N 0.000 claims description 14
- QAOWNCQODCNURD-UHFFFAOYSA-N sulfuric acid group Chemical group S(O)(O)(=O)=O QAOWNCQODCNURD-UHFFFAOYSA-N 0.000 claims description 12
- 125000000383 tetramethylene group Chemical group [H]C([H])([*:1])C([H])([H])C([H])([H])C([H])([H])[*:2] 0.000 claims description 12
- 239000007788 liquid Substances 0.000 claims description 10
- 238000002156 mixing Methods 0.000 claims description 10
- 238000004523 catalytic cracking Methods 0.000 claims description 8
- 229910052804 chromium Inorganic materials 0.000 claims description 8
- 239000011651 chromium Substances 0.000 claims description 8
- 238000005984 hydrogenation reaction Methods 0.000 claims description 8
- 239000012752 auxiliary agent Substances 0.000 claims description 7
- PNEYBMLMFCGWSK-UHFFFAOYSA-N aluminium oxide Inorganic materials [O-2].[O-2].[O-2].[Al+3].[Al+3] PNEYBMLMFCGWSK-UHFFFAOYSA-N 0.000 claims description 6
- 229910052791 calcium Inorganic materials 0.000 claims description 6
- 238000004517 catalytic hydrocracking Methods 0.000 claims description 6
- 229910052802 copper Inorganic materials 0.000 claims description 6
- 229910052593 corundum Inorganic materials 0.000 claims description 6
- 229910052759 nickel Inorganic materials 0.000 claims description 6
- BASFCYQUMIYNBI-UHFFFAOYSA-N platinum Chemical group [Pt] BASFCYQUMIYNBI-UHFFFAOYSA-N 0.000 claims description 6
- 229910001845 yogo sapphire Inorganic materials 0.000 claims description 6
- 239000002808 molecular sieve Substances 0.000 claims description 5
- URGAHOPLAPQHLN-UHFFFAOYSA-N sodium aluminosilicate Chemical group [Na+].[Al+3].[O-][Si]([O-])=O.[O-][Si]([O-])=O URGAHOPLAPQHLN-UHFFFAOYSA-N 0.000 claims description 5
- KRHYYFGTRYWZRS-UHFFFAOYSA-N Fluorane Chemical compound F KRHYYFGTRYWZRS-UHFFFAOYSA-N 0.000 claims description 4
- 238000000605 extraction Methods 0.000 claims description 4
- 229910052749 magnesium Inorganic materials 0.000 claims description 4
- 229910052697 platinum Inorganic materials 0.000 claims description 4
- 229910052700 potassium Inorganic materials 0.000 claims description 4
- 239000003377 acid catalyst Substances 0.000 claims description 3
- 239000007795 chemical reaction product Substances 0.000 claims description 3
- VYZAMTAEIAYCRO-UHFFFAOYSA-N Chromium Chemical compound [Cr] VYZAMTAEIAYCRO-UHFFFAOYSA-N 0.000 claims description 2
- 239000002283 diesel fuel Substances 0.000 claims description 2
- 229910052750 molybdenum Inorganic materials 0.000 claims description 2
- 229910052720 vanadium Inorganic materials 0.000 claims description 2
- 238000005516 engineering process Methods 0.000 abstract description 13
- -1 polypropylene Polymers 0.000 abstract description 4
- 239000003254 gasoline additive Substances 0.000 abstract description 3
- 239000004743 Polypropylene Substances 0.000 abstract description 2
- 229920001155 polypropylene Polymers 0.000 abstract description 2
- 239000003921 oil Substances 0.000 description 25
- VXNZUUAINFGPBY-UHFFFAOYSA-N 1-Butene Chemical compound CCC=C VXNZUUAINFGPBY-UHFFFAOYSA-N 0.000 description 22
- IAQRGUVFOMOMEM-UHFFFAOYSA-N but-2-ene Chemical compound CC=CC IAQRGUVFOMOMEM-UHFFFAOYSA-N 0.000 description 14
- QWTDNUCVQCZILF-UHFFFAOYSA-N isopentane Chemical compound CCC(C)C QWTDNUCVQCZILF-UHFFFAOYSA-N 0.000 description 12
- RTZKZFJDLAIYFH-UHFFFAOYSA-N ether Substances CCOCC RTZKZFJDLAIYFH-UHFFFAOYSA-N 0.000 description 9
- 150000001336 alkenes Chemical class 0.000 description 7
- ORTVZLZNOYNASJ-UPHRSURJSA-N (z)-but-2-ene-1,4-diol Chemical compound OC\C=C/CO ORTVZLZNOYNASJ-UPHRSURJSA-N 0.000 description 6
- KAKZBPTYRLMSJV-UHFFFAOYSA-N Butadiene Chemical compound C=CC=C KAKZBPTYRLMSJV-UHFFFAOYSA-N 0.000 description 6
- AFABGHUZZDYHJO-UHFFFAOYSA-N dimethyl butane Natural products CCCC(C)C AFABGHUZZDYHJO-UHFFFAOYSA-N 0.000 description 6
- IAQRGUVFOMOMEM-ONEGZZNKSA-N trans-but-2-ene Chemical compound C\C=C\C IAQRGUVFOMOMEM-ONEGZZNKSA-N 0.000 description 6
- 238000005470 impregnation Methods 0.000 description 5
- JRZJOMJEPLMPRA-UHFFFAOYSA-N olefin Natural products CCCCCCCC=C JRZJOMJEPLMPRA-UHFFFAOYSA-N 0.000 description 5
- VGGSQFUCUMXWEO-UHFFFAOYSA-N Ethene Chemical compound C=C VGGSQFUCUMXWEO-UHFFFAOYSA-N 0.000 description 4
- 239000005977 Ethylene Substances 0.000 description 4
- 150000004945 aromatic hydrocarbons Chemical class 0.000 description 4
- 238000012856 packing Methods 0.000 description 4
- 238000002360 preparation method Methods 0.000 description 4
- NWUYHJFMYQTDRP-UHFFFAOYSA-N 1,2-bis(ethenyl)benzene;1-ethenyl-2-ethylbenzene;styrene Chemical compound C=CC1=CC=CC=C1.CCC1=CC=CC=C1C=C.C=CC1=CC=CC=C1C=C NWUYHJFMYQTDRP-UHFFFAOYSA-N 0.000 description 3
- 239000003729 cation exchange resin Substances 0.000 description 3
- 238000000975 co-precipitation Methods 0.000 description 3
- 238000010586 diagram Methods 0.000 description 3
- XNMQEEKYCVKGBD-UHFFFAOYSA-N dimethylacetylene Natural products CC#CC XNMQEEKYCVKGBD-UHFFFAOYSA-N 0.000 description 3
- 238000001035 drying Methods 0.000 description 3
- 239000004005 microsphere Substances 0.000 description 3
- TVMXDCGIABBOFY-UHFFFAOYSA-N octane Chemical compound CCCCCCCC TVMXDCGIABBOFY-UHFFFAOYSA-N 0.000 description 3
- FLTJDUOFAQWHDF-UHFFFAOYSA-N 2,2-dimethylhexane Chemical class CCCCC(C)(C)C FLTJDUOFAQWHDF-UHFFFAOYSA-N 0.000 description 2
- JVSWJIKNEAIKJW-UHFFFAOYSA-N 2-Methylheptane Chemical compound CCCCCC(C)C JVSWJIKNEAIKJW-UHFFFAOYSA-N 0.000 description 2
- IJGRMHOSHXDMSA-UHFFFAOYSA-N Atomic nitrogen Chemical compound N#N IJGRMHOSHXDMSA-UHFFFAOYSA-N 0.000 description 2
- 229910002651 NO3 Inorganic materials 0.000 description 2
- NHNBFGGVMKEFGY-UHFFFAOYSA-N Nitrate Chemical compound [O-][N+]([O-])=O NHNBFGGVMKEFGY-UHFFFAOYSA-N 0.000 description 2
- 150000001335 aliphatic alkanes Chemical class 0.000 description 2
- KDKYADYSIPSCCQ-UHFFFAOYSA-N but-1-yne Chemical compound CCC#C KDKYADYSIPSCCQ-UHFFFAOYSA-N 0.000 description 2
- 239000003795 chemical substances by application Substances 0.000 description 2
- 239000000470 constituent Substances 0.000 description 2
- 238000005336 cracking Methods 0.000 description 2
- 230000007547 defect Effects 0.000 description 2
- 239000008367 deionised water Substances 0.000 description 2
- 229910021641 deionized water Inorganic materials 0.000 description 2
- 238000007323 disproportionation reaction Methods 0.000 description 2
- 238000009826 distribution Methods 0.000 description 2
- 239000000945 filler Substances 0.000 description 2
- ATRRKUHOCOJYRX-UHFFFAOYSA-N Ammonium bicarbonate Chemical compound [NH4+].OC([O-])=O ATRRKUHOCOJYRX-UHFFFAOYSA-N 0.000 description 1
- VHUUQVKOLVNVRT-UHFFFAOYSA-N Ammonium hydroxide Chemical compound [NH4+].[OH-] VHUUQVKOLVNVRT-UHFFFAOYSA-N 0.000 description 1
- NHTMVDHEPJAVLT-UHFFFAOYSA-N Isooctane Chemical compound CC(C)CC(C)(C)C NHTMVDHEPJAVLT-UHFFFAOYSA-N 0.000 description 1
- 238000010521 absorption reaction Methods 0.000 description 1
- 230000032683 aging Effects 0.000 description 1
- 150000001345 alkine derivatives Chemical class 0.000 description 1
- 229910052782 aluminium Inorganic materials 0.000 description 1
- XAGFODPZIPBFFR-UHFFFAOYSA-N aluminium Chemical compound [Al] XAGFODPZIPBFFR-UHFFFAOYSA-N 0.000 description 1
- 235000012501 ammonium carbonate Nutrition 0.000 description 1
- 239000001099 ammonium carbonate Substances 0.000 description 1
- 235000011114 ammonium hydroxide Nutrition 0.000 description 1
- 238000004458 analytical method Methods 0.000 description 1
- 239000010692 aromatic oil Substances 0.000 description 1
- 230000015572 biosynthetic process Effects 0.000 description 1
- WFYPICNXBKQZGB-UHFFFAOYSA-N butenyne Chemical group C=CC#C WFYPICNXBKQZGB-UHFFFAOYSA-N 0.000 description 1
- 125000004432 carbon atom Chemical group C* 0.000 description 1
- 150000001768 cations Chemical class 0.000 description 1
- 239000012295 chemical reaction liquid Substances 0.000 description 1
- 238000003795 desorption Methods 0.000 description 1
- 230000000694 effects Effects 0.000 description 1
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- 239000004744 fabric Substances 0.000 description 1
- 238000011049 filling Methods 0.000 description 1
- 238000001914 filtration Methods 0.000 description 1
- 239000012467 final product Substances 0.000 description 1
- 239000000446 fuel Substances 0.000 description 1
- 238000010438 heat treatment Methods 0.000 description 1
- 125000004435 hydrogen atom Chemical group [H]* 0.000 description 1
- 229910052751 metal Inorganic materials 0.000 description 1
- 239000002184 metal Substances 0.000 description 1
- 230000004048 modification Effects 0.000 description 1
- 238000012986 modification Methods 0.000 description 1
- 229910052757 nitrogen Inorganic materials 0.000 description 1
- 238000000197 pyrolysis Methods 0.000 description 1
- 238000007670 refining Methods 0.000 description 1
- 239000011347 resin Substances 0.000 description 1
- 229920005989 resin Polymers 0.000 description 1
- 230000000087 stabilizing effect Effects 0.000 description 1
- 229910001220 stainless steel Inorganic materials 0.000 description 1
- 239000010935 stainless steel Substances 0.000 description 1
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- 238000003756 stirring Methods 0.000 description 1
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- C—CHEMISTRY; METALLURGY
- C01—INORGANIC CHEMISTRY
- C01B—NON-METALLIC ELEMENTS; COMPOUNDS THEREOF; METALLOIDS OR COMPOUNDS THEREOF NOT COVERED BY SUBCLASS C01C
- C01B3/00—Hydrogen; Gaseous mixtures containing hydrogen; Separation of hydrogen from mixtures containing it; Purification of hydrogen
- C01B3/02—Production of hydrogen or of gaseous mixtures containing a substantial proportion of hydrogen
- C01B3/22—Production of hydrogen or of gaseous mixtures containing a substantial proportion of hydrogen by decomposition of gaseous or liquid organic compounds
- C01B3/24—Production of hydrogen or of gaseous mixtures containing a substantial proportion of hydrogen by decomposition of gaseous or liquid organic compounds of hydrocarbons
- C01B3/26—Production of hydrogen or of gaseous mixtures containing a substantial proportion of hydrogen by decomposition of gaseous or liquid organic compounds of hydrocarbons using catalysts
-
- C—CHEMISTRY; METALLURGY
- C07—ORGANIC CHEMISTRY
- C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
- C07C41/00—Preparation of ethers; Preparation of compounds having groups, groups or groups
- C07C41/01—Preparation of ethers
- C07C41/05—Preparation of ethers by addition of compounds to unsaturated compounds
- C07C41/06—Preparation of ethers by addition of compounds to unsaturated compounds by addition of organic compounds only
-
- C—CHEMISTRY; METALLURGY
- C07—ORGANIC CHEMISTRY
- C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
- C07C5/00—Preparation of hydrocarbons from hydrocarbons containing the same number of carbon atoms
- C07C5/32—Preparation of hydrocarbons from hydrocarbons containing the same number of carbon atoms by dehydrogenation with formation of free hydrogen
- C07C5/327—Formation of non-aromatic carbon-to-carbon double bonds only
- C07C5/333—Catalytic processes
- C07C5/3335—Catalytic processes with metals
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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
- C10G50/00—Production of liquid hydrocarbon mixtures from lower carbon number hydrocarbons, e.g. by oligomerisation
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- C—CHEMISTRY; METALLURGY
- C07—ORGANIC CHEMISTRY
- C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
- C07C2523/00—Catalysts comprising metals or metal oxides or hydroxides, not provided for in group C07C2521/00
- C07C2523/70—Catalysts comprising metals or metal oxides or hydroxides, not provided for in group C07C2521/00 of the iron group metals or copper
- C07C2523/76—Catalysts comprising metals or metal oxides or hydroxides, not provided for in group C07C2521/00 of the iron group metals or copper combined with metals, oxides or hydroxides provided for in groups C07C2523/02 - C07C2523/36
- C07C2523/84—Catalysts comprising metals or metal oxides or hydroxides, not provided for in group C07C2521/00 of the iron group metals or copper combined with metals, oxides or hydroxides provided for in groups C07C2523/02 - C07C2523/36 with arsenic, antimony, bismuth, vanadium, niobium, tantalum, polonium, chromium, molybdenum, tungsten, manganese, technetium or rhenium
- C07C2523/85—Chromium, molybdenum or tungsten
- C07C2523/86—Chromium
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- 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
- Y02P20/10—Process efficiency
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- 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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Abstract
The invention provides a method for producing high value-added products by using liquefied gas, which adopts the process combination of etherification technology, alkylation technology, aromatization technology and catalytic dehydrogenation technology to directly produce MTBE, alkylate, aromatization oil and propylene, and simultaneously produces hydrogen as a byproduct. The MTBE, the alkylate and the aromatized oil can be used as high-octane gasoline additive components, propylene can provide raw materials for downstream polypropylene devices, and the byproduct hydrogen can be used by other hydrogen consumption devices, so that the additional value of the liquefied gas is further improved. The method has the advantages of simple process flow, low dry gas yield and high product yield, and improves the utilization rate of liquefied gas resources.
Description
Technical Field
The invention relates to a method for producing methyl tert-butyl ether (MTBE), alkylate, aromatized oil and propylene by using liquefied gas and producing hydrogen gas by-product. The method can utilize all components in the liquefied gas to a great extent and realize the maximum utilization of the liquefied gas resource through the process combination.
Background
A large amount of liquefied gas is produced as by-products of a catalytic cracking gas separation device, an ethylene cracking device and a light hydrocarbon recovery device in a refinery, and the liquefied gas mainly comprises propylene, propane, n-butane, isobutane, isobutene, 1-butene, 2-butene, a small amount of C-C and the like. At present, the industrial technology for comprehensively utilizing liquefied gas in China is relatively short, and besides the liquefied gas is utilized by some enterprises to produce methyl tert-butyl ether (MTBE), most of the liquefied gas is delivered as civil fuel, so that huge resource waste is caused. Therefore, the liquefied gas is converted into petrochemical products with high added values as much as possible, the liquefied gas resources can be fully utilized, and the energy utilization rate is improved.
CN201110319824.1 discloses a method for producing aromatic hydrocarbons and olefins from liquefied gas, which utilizes olefins in liquefied gas to perform aromatization reaction, then separates propane and butane in aromatization product, and dehydrogenates separated propane and butane to produce propylene and butylene products. The method has the defects that when the yield of the aromatic hydrocarbon is high, the yield of dry gas is relatively large, and the yield of olefin is not high.
CN201210186746.7 discloses a method for producing high octane clean gasoline by combining liquefied alkane aromatization and olefin aromatization. In the method, an alkane aromatization catalyst is filled at the upper end of a reactor, an olefin aromatization catalyst is filled at the lower end, and the final product is aromatic oil. The method has the defects of single product and the need of filling two different catalysts in one reactor, and is easy to cause the problem of difficult operation control.
CN201510699100.2 discloses an improved process for aromatization of liquefied gas. The method carries out aromatization reaction on an aromatization product again after the first aromatization reaction, and leads the product to produce propane as much as possible through two aromatization reactions. The method has the problems that although the purity of the propane is high after two times of aromatization reactions, the added value of the propane is not further improved, and the propane is still mixed with a certain proportion of C4 to prepare the liquefied gas for vehicles.
CN201010256506.0 discloses a method for utilizing mixed C4, which mainly utilizes mixed C4 to carry out etherification reaction with methanol in the presence of cation exchange resin to generate MTBE and mixed C4 without isobutene, the mixed C4 without isobutene is extracted and rectified to obtain mixed butane and mixed butylene, the mixed butane is catalytically converted to produce propane and aromatic hydrocarbon, and the mixed butylene and ethylene are subjected to disproportionation reaction to produce propylene.
CN201210137002.6 discloses a method for improving the utilization value of mixed C4, which comprises the steps of using pyrolysis C4 as a raw material, carrying out selective hydrogenation to remove alkyne, hydrogenating vinyl acetylene and 1-butyne to generate 1, 3-butadiene and 1-butene, carrying out extractive distillation on the product, and separating out 1, 3-butadiene. And mixing the residual C-C material flow with the C-C of the refinery to carry out hydroisomerization, isomerizing 1-butene in the C-C mixture to obtain 2-butene, and separating the product to obtain the isobutene product. The rest material flow is subjected to disproportionation reaction to produce propylene, and the unreacted ethylene and carbon four after separation are subjected to full hydrogenation reaction to be used as ethylene cracking materials.
CN201611002809.3 discloses a method for preparing clean gasoline by aromatizing ether hydrocarbon four and naphtha, which comprises the steps of mixing dehydrated and dealkalized nitrogen naphtha with ether hydrocarbon four after methanol removal, heating and reacting, separating and refining reaction products by a light and heavy separation tank, a reaction liquid separation tank, a heavy aromatics removal tower, an absorption and desorption tower and a stabilizing tower, and finally obtaining aromatized gasoline.
CN201620571171.4 discloses a low-carbon hydrocarbon fluidized bed aromatization device. The fluidized bed device is used for aromatization reaction of low-carbon hydrocarbon, and the low-carbon hydrocarbon is firstly converted into olefin, and then aromatization reaction is carried out. Has the disadvantages of sending dehydrogenation products to aromatization, complex process flow and large investment on devices.
CN201610420079.2 discloses a low-carbon hydrocarbon fluidized bed aromatization device and application thereof. The method comprises the steps of firstly converting low-carbon hydrocarbon into olefin, and then carrying out aromatization reaction; or dehydrogenating the hydrocarbon separated by aromatization and then dearomatizing. When the yield of the aromatic hydrocarbon is high, the yield of the dry gas is relatively high.
As is clear from an analysis of the above-mentioned patent documents, the utilization of liquefied gas resources is basically involved, but the main purpose of the utilization is mostly to produce etherified gasoline from isobutylene in mixed carbon four, or further produce aromatized gasoline from mixed carbon four, and chemical products such as 1, 3-butadiene, 1-butene and 2-butene. The method has the problems that the utilization value of each component in the liquefied gas is not exploited as much as possible, and the comprehensive utilization effect is poor.
Disclosure of Invention
The invention aims to provide a method for producing a high value-added product by using liquefied gas, which fully utilizes the liquefied gas to produce a process combination of MTBE, alkylate, aromatized oil and propylene and coproducing hydrogen. The method realizes the effective utilization of the liquefied gas resource and improves the added value of the liquefied gas resource.
The process flow mainly comprises a depropanization unit, a gas separation unit, an etherification unit, an alkylation unit, an aromatization unit and a catalytic dehydrogenation unit, and mainly comprises the following steps:
the propane removal unit comprises a depropanizing tower and is used for separating a liquefied gas raw material 1 (material flow a) into propane 1 (material flow b) and propane removal mixed carbon four (material flow c).
And separating the liquefied gas raw material 2 (material flow d) through a gas separation unit to obtain propylene (material flow e), propane 2 (material flow f) and gas separation mixed C4 (material flow g).
Mixing the gas component mixed carbon four (material flow g) with methanol (material flow h), preheating after mixing, then sending to an etherification reactor, enabling an etherification product (material flow i) to pass through a catalytic distillation tower to obtain unreacted carbon four (material flow j) and MTBE (material flow k), enabling the unreacted carbon four (material flow j) to enter a methanol extraction tower, and performing methanol extraction by using water as an extractant to obtain etherified carbon four (material flow l) and a mixture of water and methanol (material flow m).
And fourthly, dividing the propane-removed mixed carbon four (material flow c) obtained in the step S into two material flows, namely material flow c1 and material flow c2 in proportion. Wherein the stream c1 is mixed with the four carbon atoms (stream l) after the ether, and the mixture is jointly sent to an alkylation reactor, and isobutane and butene are subjected to alkylation reaction under the action of a catalyst. After the reaction is finished, sending the alkylate (stream n) into an alkylate separation tower to obtain alkylate (stream o) and alkylate tail gas (stream p).
Sending the material flow c2 in the step four and the alkylation tail gas (material flow p) to an aromatization reactor, carrying out an aromatization reaction on butane under the action of a catalyst, and sending an aromatization product (material flow q) to an aromatization product separation tower after the reaction is finished to obtain aromatization oil (material flow r) and aromatization tail gas (material flow s).
Sixthly, sending propane 1 (stream b) obtained in the step, propane 2 (stream f) obtained in the step and the aromatization tail gas (stream s) obtained in the step to a catalytic dehydrogenation unit together, and carrying out dehydrogenation reaction on propane under the action of a catalyst to obtain propylene (stream t) and hydrogen as a byproduct.
Wherein: the sum of isobutane and normal butane in liquefied gas raw material 1 is not less than 40%, preferably not less than 45%. The isobutene content of the liquefied gas raw material 2 is not less than 10%, preferably not less than 11%, and the isobutane content is not less than 13%, preferably not less than 15%.
The liquefied gas raw material 1 is preferably one or two of light hydrocarbon liquefied gas and hydrogen recovery liquefied gas.
The liquefied gas raw material 2 is preferably refinery catalytic cracking liquefied gas.
The light hydrocarbon liquefied gas is a product obtained by one or a mixture of more of isomerized dry gas, residual oil hydrogenation stripping tower top gas, diesel oil hydrogenation stripping tower top gas, hydrocracking stripping tower top gas, naphtha hydrogenation reaction products, hydrocracking stripping tower top liquid, hydrocracking fractionating tower top liquid and reformed liquefied gas through a light hydrocarbon recovery device. The content of isobutane in the light hydrocarbon liquefied gas is preferably not less than 25 percent, more preferably not less than 30 percent; the n-butane content is not less than 30%, preferably not less than 35%.
The hydrogen recovery liquefied gas is the liquefied gas produced by a hydrogen recovery device by taking one or two of catalytic dry gas and raffinate oil as raw materials. The content of isobutane in the hydrogen-recovered liquefied gas is preferably not less than 28 percent, more preferably not less than 30 percent; the n-butane content is not less than 12%, preferably not less than 15%.
The specific implementation method can be as follows:
the method comprises the steps of sending a liquefied gas raw material 1 (material flow a) into a depropanizing tower, wherein the depropanizing tower adopts a plate tower, and the separation products of the liquefied gas raw material 1 are propane 1 (material flow b) and depropanized mixed carbon four (material flow c). The yield of propane is not less than 99%, preferably not less than 99.5%.
And step II, sending the liquefied gas raw material 2 (material flow d) to a gas separation unit for separation to obtain propylene (material flow e), propane (material flow f) and gas separation mixed carbon four (material flow g). The yield of the gas-component mixed C4 (material flow g) is not less than 99 percent, and is preferably 99.5 percent.
And step three, the etherification unit consists of an etherification reactor, a catalytic distillation tower and a methanol extraction tower. The etherification reactor adopts a fixed bed reactor, and the number of the fixed bed reactors is one or two. When two fixed bed reactors are used, the distribution mode is series connection or parallel connection, and the distribution mode is realized by the following connection mode: the upper parts and the upper parts, the lower parts and the lower parts, and the upper parts and the lower parts of the two fixed bed reactors are sequentially connected by valves and pipelines, and the two fixed bed reactors are connected in series or in parallel by switching the valves, and a schematic diagram is shown in figure 2. When the two fixed bed reactors are connected in series, the depth of the etherification reaction can be increased, and when the two fixed bed reactors are connected in parallel, the feeding can be switched on line, so that the catalyst can be replaced on line. The etherification reactor and the catalytic distillation tower are respectively filled with etherification catalysts, and etherification is carried out under the conditions that the reaction temperature is 10-80 ℃, preferably 25-50 ℃, the reaction pressure is 0.01-3.0 MPa, preferably 0.05-1.5 MPa, and the alcohol-hydrocarbon ratio (mass ratio of methanol to hydrocarbon feed) is 0.1-1.0, preferably 0.15-0.5. The mixture of gas-mixed C.sub.four (stream g) and methanol (stream h) is fed from the bottom of the fixed bed, and the etherification product (stream i) at the top of the fixed bed is introduced into the catalytic distillation column. The catalytic distillation tower is filled with a filler besides a macroporous cation exchange resin catalyst. Therefore, etherification reaction and separation of etherification products can be simultaneously carried out in the catalytic distillation tower, unreacted carbon four (material flow j) is obtained at the top of the catalytic distillation tower, MTBE (material flow k) is obtained at the bottom of the catalytic distillation tower, and the purity of the MTBE is not lower than 98%. Because unreacted carbon four (material flow j) obtained at the top of the catalytic distillation tower contains a certain amount of methanol, the unreacted carbon four (material flow j) needs to be introduced into a methanol extraction tower for methanol extraction, water is used as an extracting agent, the feeding temperature of the extraction tower is 20-70 ℃, preferably 30-60 ℃, and the pressure is 0.1-2.0 MPa, preferably 0.5-1.5 MPa. After extraction with methanol, ethereal carbon four (stream l) and a mixture of water and methanol (stream m) are obtained. The methanol content of the mixture of water and methanol is controlled to be not higher than 30%, preferably not higher than 25%.
The etherification catalyst is used in the field, such as a macroporous cation exchange resin catalyst, the mass total exchange capacity is 2-10 mmol/g, the water content is 1-9%, and the wet apparent density is 0.6-1.5 g/ml.
The present invention is not limited to the type of methanol extraction column, and may be a plate column or a packed column.
The alkylation reaction unit in the fourth step includes an alkylation reactor and an alkylate separation tower. Dividing the depropanized mixed carbon four (material flow c) obtained in the depropanizing tower in the step into two material flows according to a certain proportion, namely material flow c1 and material flow c 2. The material flow c1 accounts for 25-35%, preferably 27-30% of the total mass of the depropanized mixed carbon four (material flow c). And c1, mixing the stream c and the etherified carbon four (stream l) in the step three, sending the mixture into an alkylation reactor, carrying out alkylation reaction on isobutane and butene under the action of a catalyst under the conditions that the reaction temperature is 0.1-20 ℃, preferably 2-10 ℃, the reaction pressure is 0-10 MPa, preferably 0.5-2.0 MPa, and sending the alkylation product (stream n) into an alkylation product separation tower after the reaction is finished to obtain alkylate oil (stream o) and alkylation tail gas (stream p). The major constituents of the alkylate are Trimethylpentanes (TMP) and (DMH), the Research Octane Number (RON) is about 95, and the major constituent of the alkylation tail gas (stream p) is butane.
The catalyst in the alkylation reactor is an alkylation catalyst used in the art, preferably a liquid acid catalyst such as sulfuric acid or hydrofluoric acid.
The aromatization reaction unit in the step fifthly comprises an aromatization reactor and an aromatization product separation tower. The material flow c2 and the alkylation tail gas (material flow p) are sent to an aromatization reactor together, under the action of a catalyst, the reaction temperature is 300-600 ℃, the preferable temperature is 350-550 ℃, the pressure is 0.05-4.0 MPa, the preferable pressure is 0.08-4.0 MPa, and the liquid hourly space velocity is 0.01-10 h-1Preferably 0.05 to 5 hours-1Under the condition, butane is subjected to aromatization reaction, and the content of aromatization oil in an aromatization product (material flow q) is not less than 22%. After the aromatization product is separated, aromatization oil (material flow r) and aromatization tail gas (material flow s) are obtained, and the Research Octane Number (RON) of aromatization oil can be up to above 90.
The catalyst in the aromatization reactor is an aromatization catalyst used in the art, preferably a molecular sieve catalyst such as ZSM-5 or a metal modified HZSM-5 molecular sieve catalyst.
The sixth step is a catalytic dehydrogenation unit, and the propane 1 (material flow b) obtained in the first step, the propane 2 (material flow f) obtained in the second step and the aromatization tail gas (material flow s) obtained in the second step are jointly sent to the catalytic dehydrogenation unit. The main component of the aromatization tail gas (stream s) is propane.
The catalyst used in the catalytic dehydrogenation reactor is a dehydrogenation catalyst of the art, such as a platinum or chromium based catalyst.
Dehydrogenation catalyst with Al2O3The carrier is loaded with one or more active components selected from Pt, Cr, Mo and V, and one or more auxiliary agents selected from Ni, Cu, Ca, Mg and K. The catalyst is preferably prepared by a coprecipitation method or an impregnation method. When the catalyst is prepared by adopting a coprecipitation method, nitrate with a proper ratio is dissolved in deionized water, ammonia water/ammonium carbonate is dropwise added under vigorous stirring, and aging is carried out for 3-48 h. Then filtering, washing, drying and roasting to form the catalyst. When the catalyst is prepared by adopting an impregnation method,dissolving nitrate of aluminum and an active component in a proper proportion in a certain amount of deionized water, adding a carrier, slowly drying at 40-150 ℃, and then drying, pre-decomposing and roasting to obtain the catalyst. The dehydrogenation catalyst comprises 50-90% of a carrier, 1-40% of active component elements and 0.1-30% of auxiliary agent elements.
The catalytic dehydrogenation reaction conditions recommended by the invention are as follows: the reaction temperature is 400-800 ℃, preferably 500-700 ℃, the reaction pressure is 0-1 MPa, preferably 0.1-0.5 MPa, and the feeding airspeed is 1.0-6.0 h-1Preferably 1.0 to 4.0 hours-1. In the process, the best propane conversion rate is 30-50%, and the propylene selectivity is 60-80%.
The catalytic dehydrogenation reactor can be a fixed bed reactor or a fluidized bed reactor.
The catalyst used in the catalytic dehydrogenation unit may be an extruded catalyst suitable for fixed bed use or a microspherical catalyst suitable for fluidized bed use. The dehydrogenation reactor can be a fixed bed reactor or a fluidized bed reactor according to different catalyst forms.
The method adopts the process combination of etherification technology, alkylation technology, aromatization technology and catalytic dehydrogenation technology, can directly produce MTBE, alkylate, aromatization oil and propylene, and can produce hydrogen as a byproduct. The method can utilize all components in the liquefied gas to a great extent, and realizes the maximum utilization of the liquefied gas resource through process combination. The invention is characterized in that:
the invention can fully utilize rich mixed C4 in liquefied gas, etherify isobutene and methanol to generate MTBE, and alkylate oil and aromatized butane to generate aromatized oil, wherein the three products can be used as high-octane gasoline additive components.
The invention utilizes propane separated from the liquefied gas of the light hydrocarbon liquefied gas and the liquefied gas of the hydrogen recovery device and the aromatization tail gas produced by the aromatization unit to carry out dehydrogenation reaction to generate propylene, which can provide raw materials for a downstream polypropylene device and simultaneously produce a certain amount of hydrogen as a byproduct, and the hydrogen as the byproduct can be used by other hydrogen consuming devices, thereby improving the added value of the liquefied gas.
The invention combines the etherification technology, the alkylation technology, the aromatization technology and the catalytic dehydrogenation technology, has simple process flow, low dry gas yield and high product yield, and realizes the optimal utilization of liquefied gas resources.
Drawings
FIG. 1 is a process flow diagram of the application of the present invention.
In fig. 1: 1-gas separation unit, 2-etherification reactor, 3-catalytic distillation tower, 4-methanol extraction tower, 5-alkylation reactor, 6-alkylation product separation tower, 7-aromatization reactor, 8-aromatization product separation tower, 9-depropanization tower and 10-catalytic dehydrogenation unit.
Stream a-liquefied gas feed 1; stream b-propane 1; stream c-depropanized mixed C4; stream d-liquefied gas feed 2; stream e-propylene; stream f-propane 2; material flow g-gas component mixed carbon four; stream h-methanol; a stream i-etherification product; stream j-unreacted carbon four; a stream k-MTBE; stream l-ether post carbon four; stream m-a mixture of water and methanol; stream n-alkylate; stream o-alkylate; stream p-alkylation tail gas; a stream q-aromatization product; stream r-aromatized oil; stream s-aromatization tail gas; stream t-propene.
FIG. 2 is a schematic diagram of the connection of an etherification unit according to the present invention using two fixed bed reactors.
In the figure, 1# and 2# are fixed bed reactors, respectively.
Detailed Description
Example 1
The catalytic cracking liquefied gas, the light hydrocarbon recovery liquefied gas and the hydrogen recovery liquefied gas are used as raw materials, and the composition of the raw materials is shown in table 1. Sending catalytic cracking liquefied gas, namely liquefied gas raw material 2 (material flow d) into a gas separation unit 1, obtaining propylene (material flow e), propane 2 (material flow f) and gas separation mixed carbon four (material flow g) after passing through the gas separation unit 1, wherein the composition of the gas separation mixed carbon four (material flow g) is shown in a table 2, mixing light hydrocarbon recovery liquefied gas and hydrogen recovery liquefied gas into liquefied gas raw material 1 (material flow a), sending the liquefied gas raw material 1 (material flow a) into a depropanizing tower 9, obtaining propane 1 (material flow b) and depropanization mixed carbon four (material flow c), and the composition of the material flow c is shown in a table 3.
TABLE 1 raw material composition
TABLE 2 composition of gas-component mixed C4
TABLE 3 composition of depropanized C4 mixtures
| Mixed carbon four components | Content/wt% |
| Isobutane | 45.42 |
| N-butane | 49.88 |
| N-butene | 0.13 |
| Isobutene | 0.42 |
| Butene of trans-butene | 0.11 |
| Cis-butenediol | 0.23 |
| C5+ (containing isopentane) | 3.89 |
| Total up to | 100.00 |
Mixing the gas-component mixed carbon four (material flow g) and methanol (material flow h) and sending the mixture to an etherification reactor 2, wherein the etherification reactor 2 adopts a fixed bed reactor, and the etherification reaction is carried out under the conditions of the reaction temperature of 40 ℃, the reaction pressure of 0.8MPa and the alcohol-hydrocarbon ratio of 0.45, and the properties of the methanol (material flow h) are shown in a table 4. The etherified product (stream i) after the reaction is fed to a catalytic distillation column 3.
Macroporous cation resin catalysts are respectively filled in the etherification reactor 2 and the catalytic distillation tower 3. The catalyst in the etherification reactor 2 is filled in sections, so that an expansion space of the catalyst in a reaction state is reserved, and the formation of hot spots is well avoided. The catalytic distillation tower 3 is filled with filler besides the etherification catalyst. The catalyst in the catalytic distillation column 3 was packed in the packing unit of example 1 in ZL 201520508723.2. The total mass exchange capacity of the etherification catalyst is 2-10 mmol/g, the water content is 1-9%, and the wet apparent density is 0.6-1.5 g/ml.
Unreacted carbon four (material flow j) is obtained at the top of the catalytic distillation tower 3, and MTBE (material flow k) is obtained at the bottom of the tower. And (3) feeding the unreacted carbon four (material flow j) into a methanol extraction tower 4 for methanol extraction, wherein the feeding temperature is 45 ℃, the pressure is 0.8MPa, the methanol extraction tower 4 adopts a plate tower, and the extracting agent is water. The top of the methanol extraction tower 4 is ether carbon four (material flow l), the content of methanol in the ether carbon four (material flow l) is not higher than 1%, the bottom of the tower is a mixture of water and methanol (material flow m), and the content of methanol in the mixture of water and methanol (material flow m) is 18.6%.
TABLE 4 methanol compliance with GB 338-
The depropanized mixed C4 (material flow c) is divided into two material flows according to a certain proportion: stream c1 and stream c2, the ratio of stream c1 to stream c2 in this example being 1:2.6, i.e. stream c1 represents 27.8% of the depropanized C4 (stream c). Wherein stream c1 is mixed with the post-etherification carbon four (stream l) and is jointly fed to the alkylation reactor 5 for alkylation.
The catalytic system in the alkylation reactor 5 is sulfuric acid, the reaction feeding is 8 ℃, the reaction pressure is 1MPa, isobutane and butylene are subjected to alkylation reaction, the conversion rate of butylene is not lower than 98.5%, an alkylation product (material flow n) is introduced into an alkylation product separation tower 6 after the reaction is finished, the alkylation product separation tower 6 adopts a plate tower, the top of the tower is alkylation tail gas (material flow p), the main component of the alkylation tail gas (material flow p) is butane, and the butane content is 74%. The bottom of the tower is alkylate oil (flow o), and the main components of the alkylate oil (flow o) are trimethylpentane and dimethylhexane.
Mixing the alkylation tail gas (stream p) and the stream c2 and sending the mixture into an aromatization reactor 7, wherein the aromatization catalyst is the aromatization catalyst SIHZSM-5(A) -01 prepared in the catalyst preparation example 7 in the patent CN1586721A, the reaction temperature is 430 ℃, the reaction pressure is 3.0MPa, and the space velocity is 2.5h-1Under the conditions of (1), carrying out aromatization reaction. The aromatization product (material flow q) is separated by an aromatization product separation tower 8 to obtain aromatization oil (material flow r) and aromatization tail gas (material flow s), wherein the propane content in the aromatization tail gas (material flow s) is 96.8 percent.
Propane 1 (stream b), propane 2 (stream f) and aromatization tail gas (stream s) are jointly fed into the catalytic dehydrogenation unit 10 for catalytic dehydrogenation reaction. The catalytic dehydrogenation unit 10 employs a fixed bed reactor, and is filled with a catalytic dehydrogenation catalyst. Catalytic dehydrogenation catalyst with Al2O3The catalyst is a carrier, carries Cr, Ni, Cu and Ca elements, is prepared by a coprecipitation method, and has the catalyst carrier content of 85 percent, the active component content of 10 percent and the auxiliary agent content of 5 percent. The catalytic dehydrogenation reaction is carried out at the temperature of 605 ℃, the reaction pressure of 0.2MPa and the space velocity of 2.5h-1Under the conditions of (1) by catalytic dehydrationThe dehydrogenation product obtained in the hydrogen unit 10 is mixed with the propylene obtained in the gas separation unit 1 (stream e) to obtain propylene (stream t).
The mass yield of each component of the system in the process is as follows: 4.77% of dry gas, 35.98% of propylene, 17.66% of propane, 10.57% of MTBE, 24.50% of alkylate oil and 6.27% of aromatized oil, wherein hydrogen in the dry gas accounts for 79.04 vol%.
Example 2
The catalytic cracking liquefied gas (liquefied gas raw material 2 (material flow d) and the light hydrocarbon recovery liquefied gas (liquefied gas raw material 1 (material flow a)) are used as raw materials, the composition of the raw materials is shown in table 1, and the properties of the methanol (material flow h) raw material are shown in table 4. The process flow in example 2 is the same as in example 1. The composition of the gas-divided carbon four (stream g) is shown in Table 5, and the composition of the depropanized carbon four (stream c) is shown in Table 6.
TABLE 5 composition of gas mixture carbon four
| Mixed carbon four components | Content/wt% |
| Isobutane | 28.00 |
| N-butane | 6.40 |
| N-butene | 14.00 |
| Isobutene | 25.00 |
| Butene of trans-butene | 14.00 |
| Cis-butenediol | 8.00 |
| C5+ (containing isopentane) | 4.60 |
| Total up to | 100.00 |
TABLE 6 composition of depropanized C4 mixtures
| Mixed carbon four components | Content/wt% |
| Isobutane | 44.00 |
| N-butane | 51.00 |
| N-butene | 0.10 |
| Isobutene | 0.50 |
| Butene of trans-butene | 0.18 |
| Cis-butenediol | 0.25 |
| C5+ (containing isopentane) | 4.00 |
| Total up to | 100.00 |
Example 2 differs from example 1 in several respects:
the etherification reaction is carried out at the temperature of 48 ℃, the reaction pressure of 1.0MPa and the alcohol-hydrocarbon ratio of 0.48. The catalyst in the catalytic distillation tower 3 is packed in a way that the catalyst is wrapped by a special cloth bag and then is packed into the packing member of ZL201520508723.2 in example 1. In the methanol extraction tower 4, the feeding temperature of unreacted carbon four (material flow j) is 48 ℃, the pressure is 0.9MPa, the tower bottom is a mixture of water and methanol (material flow m), and the content of methanol in the mixture of water and methanol (material flow m) is 17.9 percent.
② the catalytic system in the alkylation reactor 5 is sulfuric acid, the reaction feeding temperature is 5 ℃, and the reaction pressure is 0.8 MPa. Isobutane and butylene are subjected to alkylation reaction, the conversion rate of the butylene is not lower than 99%, and an alkylation tail gas (material flow p) is obtained from the top of the alkylation product separation tower 6, wherein the butane content is 71%.
③ the aromatization catalyst adopts the aromatization catalyst SIHZSM-5(A) -02 prepared in the catalyst preparation example 7 in the patent CN1586721A, the reaction temperature is 440 ℃, the reaction pressure is 2.5MPa, and the space velocity is 2.5h-1Under the conditions of (1), carrying out aromatization reaction. The aromatization product (material flow q) is separated by an aromatization product separation tower 8 to obtain aromatization oil (material flow r) and aromatization tail gas (material flow s), wherein the propane content in the aromatization tail gas (material flow s) is 96 percent.
And fourthly, the catalytic dehydrogenation unit 10 adopts a fluidized bed reactor, and the catalyst is a microsphere catalyst. With Al2O3The carrier is loaded with Cr, Ni, Cu, Ca and K elements and prepared by an impregnation method, wherein the carrier content is 82%, the active component content is 15% and the auxiliary agent content is 3%. Under the conditions that the reaction temperature is 600 ℃, the reaction pressure is 0.15MPa,the space velocity is 2.0h-1The dehydrogenation reaction is carried out under the conditions of (1).
The mass yield of each component of the system in the process is as follows: 4.45% of dry gas, 36.55% of propylene, 16.47% of propane, 11.30% of MTBE, 25.14% of alkylate and 6.09% of aromatized oil, wherein hydrogen in the dry gas accounts for 79.05 vol%.
Example 3
The catalytic cracking liquefied gas (liquefied gas raw material 2 (material flow d)) and the hydrogen recovery liquefied gas (liquefied gas raw material 1 (material flow a)) are used as raw materials, the composition of the raw materials is shown in table 1, and the properties of the methanol (material flow h) raw material are shown in table 4. The process flow in example 3 is the same as in example 1. The composition of the gas-divided carbon four (stream g) is shown in Table 7, and the composition of the depropanized carbon four (stream c) is shown in Table 8. Example 3 differs from example 1 in several respects:
the etherification reaction is carried out at the temperature of 48 ℃, the reaction pressure of 0.9MPa and the alcohol-hydrocarbon ratio of 0.4. The catalyst was packed in a special stainless steel mesh structure, and then packed in a catalyst packing member of ZL201520508723.2 in example 1, and packed in the catalytic distillation column 3. In the methanol extraction tower 4, the feeding temperature of unreacted carbon four (material flow j) is 50 ℃, the pressure is 0.7MPa, the tower bottom is a mixture of water and methanol (material flow m), and the content of methanol in the mixture of water and methanol (material flow m) is 22.2 percent.
② the catalytic system in the alkylation reactor 5 is sulfuric acid, the reaction feeding temperature is 4 ℃, and the reaction pressure is 0.9 MPa. Isobutane and butylene are subjected to alkylation reaction, the conversion rate of the butylene is not lower than 99%, and an alkylation tail gas (material flow p) is obtained at the top of the alkylation product separation tower 6, wherein the butane content is 35%.
③ the aromatization catalyst adopts the aromatization catalyst SIHZSM-5(A) -03 prepared in the catalyst preparation example 7 in the patent CN1586721A, the reaction temperature is 435 ℃, the reaction pressure is 3.0MPa, and the space velocity is 2.0h-1Under the conditions of (1), carrying out aromatization reaction. The aromatization product (material flow q) is separated by an aromatization product separation tower 8 to obtain aromatization oil (material flow r) and aromatization tail gas (material flow s), wherein the propane content in the aromatization tail gas (material flow s) is 95 percent.
TABLE 7 composition of gas mixture carbon four
TABLE 8 composition of depropanized C4 mixtures
| Mixed carbon four components | Content/wt% |
| Isobutane | 44.80 |
| N-butane | 49.47 |
| N-butene | 0.20 |
| Isobutene | 0.58 |
| Butene of trans-butene | 0.25 |
| Cis-butenediol | 0.40 |
| C5+ (containing isopentane) | 4.30 |
| Total up to | 100.00 |
The mass yield of each component of the system in the process is as follows: 2.76% of dry gas, 36.60% of propylene, 10.23% of propane, 15.87% of MTBE, 28.38% of alkylate oil and 3.16% of aromatized oil, wherein the hydrogen in the dry gas accounts for 79.05 vol%.
Example 4
The starting material of example 4 was the same as that of example 1: the catalytic cracking liquefied gas is liquefied gas raw material 2 (material flow d), and the light hydrocarbon recovery liquefied gas and the hydrogen recovery liquefied gas are mixed into liquefied gas raw material 1 (material flow a). The composition of the feed is shown in Table 1, and the properties of the methanol (stream h) feed are shown in Table 4. The process flow in example 4 is the same as in example 1. The composition of the gas-divided carbon four (stream g) is shown in Table 9, and the composition of the depropanized carbon four (stream c) is shown in Table 10. Example 4 differs from example 1 in several respects:
TABLE 9 composition of gas mixture carbon four
| Mixed carbon four components | Content/wt% |
| Isobutane | 28.70 |
| N-butane | 6.22 |
| N-butene | 13.80 |
| Isobutene | 20.80 |
| Butene of trans-butene | 14.80 |
| Cis-butenediol | 9.90 |
| C5+ (containing isopentane) | 5.78 |
| Total up to | 28.70 |
TABLE 10 composition of depropanized C4 mixtures
| Mixed carbon four components | Content/wt% |
| Isobutane | 44.83 |
| N-butane | 50.22 |
| N-butene | 0.17 |
| Isobutene | 0.44 |
| Butene of trans-butene | 0.14 |
| Cis-butenediol | 0.29 |
| C5+ (containing isopentane) | 3.90 |
| Total up to | 100.00 |
The etherification reaction is carried out at the temperature of 45 ℃, the reaction pressure of 1.0MPa and the alcohol-hydrocarbon ratio of 0.5. The catalyst was packed in the catalytic distillation column 3 in the same manner as in example 1 of ZL201520508723.2, and the catalyst was directly packed in the packing member. In the methanol extraction tower 4, the feeding temperature of unreacted carbon four (material flow j) is 42 ℃, the pressure is 1.0MPa, the tower bottom is a mixture of water and methanol (material flow m), and the content of methanol in the mixture of water and methanol (material flow m) is 20.3 percent.
② the catalytic system in the alkylation reactor 5 is sulfuric acid, the reaction feeding temperature is 8 ℃, and the reaction pressure is 1.0 MPa. Isobutane and butylene are subjected to alkylation reaction, the conversion rate of the butylene is not lower than 99%, and an alkylation tail gas (material flow p) is obtained at the top of the alkylation product separation tower 6, wherein the content of the butane is 68%.
③ the aromatization catalyst adopts the aromatization catalyst SIHZSM-5(A) -04 prepared in the catalyst preparation example 7 in the patent CN1586721A, the reaction temperature is 430 ℃, the reaction pressure is 2.5MPa, and the space velocity is 2.0h-1Under the conditions ofAnd (4) carrying out aromatization reaction. The aromatization product (material flow q) is separated by an aromatization product separation tower 8 to obtain aromatization oil (material flow r) and aromatization tail gas (material flow s), wherein the propane content in the aromatization tail gas (material flow s) is 9.58 percent.
And fourthly, the catalytic dehydrogenation unit 10 adopts a fluidized bed reactor, and the catalyst is a microsphere catalyst. With Al2O3The carrier is loaded with Cr, Ni, Cu, Ca, Mg and K elements and prepared by an impregnation method, wherein the carrier content is 84%, the active component content is 12% and the auxiliary agent content is 4%. The reaction temperature is 595 ℃, the reaction pressure is 0.2MPa, and the space velocity is 2.0h-1The dehydrogenation reaction is carried out under the conditions of (1).
The mass yield of each component of the system in the process is as follows: 5.16% of dry gas, 36.47% of propylene, 17.30% of propane, 10.47% of MTBE, 24.26% of alkylate and 6.35% of aromatized oil, wherein hydrogen in the dry gas accounts for 82.32 vol%.
The above examples show that the method of the present invention can produce a certain proportion of gasoline additive components, and also produce propylene by dehydrogenating propane, thereby increasing the added value of propane and producing hydrogen as a byproduct. The residual propane is less, and the propane separated from the product can also be recycled to a propane dehydrogenation unit for further dehydrogenation.
The present invention is capable of other embodiments, and various changes and modifications may be made by one skilled in the art without departing from the spirit and scope of the invention as defined in the appended claims.
Claims (30)
1. A method for producing high value-added products by using liquefied gas is characterized in that the process flow mainly comprises a depropanization unit, a gas separation unit, an etherification unit, an alkylation unit, an aromatization unit and a catalytic dehydrogenation unit, and comprises the following steps:
the propane removing unit comprises a depropanizing tower and is used for separating a liquefied gas raw material 1 into propane 1 and a depropanized mixed carbon four;
separating the liquefied gas raw material 2 through a gas separation unit to obtain propylene, propane 2 and gas separation mixed carbon four;
mixing the gas component mixed C4 with methanol, preheating after mixing, conveying to an etherification reactor, allowing an etherification product to pass through a catalytic distillation tower to obtain unreacted C4 and MTBE, allowing the unreacted C4 to enter a methanol extraction tower, and performing methanol extraction by using water as an extractant to obtain a mixture of etherified C4, water and methanol;
dividing the depropanized mixed carbon four obtained in the step I into two streams of material flow, namely material flow c1 and material flow c2, wherein the material flow c1 is mixed with the etherified carbon four, the mixture is jointly sent to an alkylation reactor, isobutane and butylene are subjected to alkylation reaction under the action of a catalyst, and after the reaction is finished, an alkylation product is sent to an alkylation product separation tower, so that alkylate oil and alkylation tail gas are obtained;
sending the material flow c2 in the step four and the alkylation tail gas to an aromatization reactor, carrying out an aromatization reaction on butane under the action of a catalyst, and sending an aromatization product to an aromatization product separation tower after the reaction is finished to obtain aromatization oil and aromatization tail gas;
sending propane 1 obtained in the step I, propane 2 obtained in the step II and aromatization tail gas obtained in the step I to a catalytic dehydrogenation unit together, and carrying out dehydrogenation reaction on propane under the action of a catalyst to obtain propylene and hydrogen as a byproduct;
wherein: the sum of isobutane and normal butane in the liquefied gas raw material 1 is not less than 40 percent; the isobutene content in the liquefied gas raw material 2 is not less than 10 percent, and the isobutane content is not less than 13 percent.
2. The method for producing high added-value products by using liquefied gas according to claim 1, wherein the sum of isobutane and n-butane in the liquefied gas raw material 1 is not less than 45%; the isobutene content in the liquefied gas raw material 2 is not lower than 11%, and the isobutane content is not lower than 15%.
3. The method for producing high added-value products by using liquefied gas as claimed in claim 1, wherein liquefied gas raw material 1 is one or both of light hydrocarbon liquefied gas and hydrogen-recovered liquefied gas.
4. The method for producing high added-value products by using liquefied gas as claimed in claim 3, wherein the light hydrocarbon liquefied gas is: and (3) a product obtained by one or a mixture of more of isomerized dry gas, residual oil hydrogenation stripping tower top gas, diesel oil hydrogenation stripping tower top gas, hydrocracking stripping tower top gas, naphtha hydrogenation reaction products, hydrocracking stripping tower top liquid, hydrocracking fractionating tower top liquid and reformed liquefied gas through a light hydrocarbon recovery device.
5. The method for producing high added-value products by using liquefied gas according to claim 3, wherein the liquefied gas recovered by using hydrogen is the liquefied gas produced by using one or both of catalytic dry gas and raffinate oil as raw materials and using a hydrogen recovery device.
6. The method for producing high added-value products by using liquefied gas as claimed in claim 3, wherein the content of isobutane in the light hydrocarbon liquefied gas is not less than 25% and the content of n-butane in the light hydrocarbon liquefied gas is not less than 30%.
7. The method for producing high added-value products by using liquefied gas as claimed in claim 6, wherein the content of isobutane in the light hydrocarbon liquefied gas is not less than 30% and the content of n-butane in the light hydrocarbon liquefied gas is not less than 35%.
8. The method for producing high added-value products using liquefied gas according to claim 3, wherein the content of isobutane in the hydrogen-recovered liquefied gas is not less than 28% and the content of n-butane in the hydrogen-recovered liquefied gas is not less than 12%.
9. The method for producing high added-value products using liquefied gas according to claim 8, wherein the content of isobutane in the hydrogen-recovered liquefied gas is not less than 30% and the content of n-butane in the hydrogen-recovered liquefied gas is not less than 15%.
10. The method for producing high added-value products using liquefied gas according to claim 1, wherein the liquefied gas raw material 2 is a catalytic cracking liquefied gas.
11. The method for producing high added-value products by using liquefied gas as claimed in claim 1, wherein the material flow c1 accounts for 25-35% of the total mass of the depropanized C4 mixture.
12. The method for producing high added-value products by using liquefied gas as claimed in claim 11, wherein the material flow c1 accounts for 27-30% of the total mass of the depropanized C4 mixture.
13. The method for producing high added-value products by using liquefied gas according to claim 1, wherein the etherification reactor is a fixed bed reactor, and the number of the fixed bed reactors is one or two.
14. The method for producing high added-value products by using liquefied gas according to claim 1, wherein the etherification reaction temperature is 10 to 80 ℃, the reaction pressure is 0.01 to 3.0MPa, and the mass ratio of methanol to hydrocarbon feedstock is 0.1 to 1.0.
15. The method for producing high added-value products by using liquefied gas according to claim 14, wherein the etherification reaction temperature is 25 to 50 ℃, the reaction pressure is 0.05 to 1.5MPa, and the mass ratio of methanol to hydrocarbon feedstock is 0.15 to 0.5.
16. The method for producing high added-value products by using liquefied gas according to claim 1, wherein the methanol content in the mixture of water and methanol obtained from the methanol extraction tower is not higher than 30%.
17. The method for producing high added-value products using liquefied gas according to claim 16, wherein the methanol content in the mixture of water and methanol obtained in the methanol extraction tower is not higher than 25%.
18. The method for producing high added-value products by using liquefied gas according to claim 1, wherein the feeding temperature of the extraction tower is 20-70 ℃, and the pressure is 0.1-2.0 MPa.
19. The method for producing high added-value products by using liquefied gas as claimed in claim 18, wherein the feeding temperature of the extraction tower is 30-60 ℃ and the pressure is 0.5-1.5 MPa.
20. A method for producing high added-value products using liquefied gas according to claim 1, wherein the catalyst in the alkylation reactor is a liquid acid catalyst.
21. A method for producing high added-value products using liquefied gas according to claim 20, wherein the liquid acid catalyst is sulfuric acid or hydrofluoric acid.
22. The method for producing high value-added products by using liquefied gas according to claim 1, wherein the alkylation reaction temperature is 0.1-20 ℃ and the reaction pressure is 0-10 MPa.
23. The method for producing high value-added products by using liquefied gas according to claim 22, wherein the alkylation reaction temperature is 2 to 10 ℃ and the reaction pressure is 0.5 to 2.0 MPa.
24. The method for producing high added-value products using liquefied gas according to claim 1, wherein the catalyst in the aromatization reactor is a molecular sieve catalyst.
25. A method for producing high added-value products using liquefied gas according to claim 24, wherein the molecular sieve catalyst is ZSM-5 or a metal-modified HZSM-5 molecular sieve catalyst.
26. Method for producing high added value products using liquefied gas according to claim 1, characterized in thatThe aromatization reaction temperature is 300-600 ℃, and the reaction pressure is 0.05-4.0 MPa; the liquid hourly space velocity is 0.01-10 h-1。
27. The method for producing high added-value products by using liquefied gas according to claim 26, wherein the aromatization reaction temperature is 350 to 550 ℃, and the reaction pressure is 0.08 to 4.0 MPa; the liquid hourly space velocity is 0.05-5 h-1。
28. The method for producing high added-value products using liquefied gas according to claim 1, wherein the catalyst used in the catalytic dehydrogenation reactor is a platinum-based or chromium-based catalyst.
29. The method for producing high added-value products using liquefied gas according to claim 1, wherein the reactor of the catalytic dehydrogenation unit is a fixed bed reactor or a fluidized bed reactor.
30. The method for producing high added-value products using liquefied gas according to claim 1, wherein the dehydrogenation catalyst is Al2O3The carrier is used, the loaded active component is one or more of Pt, Cr, Mo and V elements, and the loaded auxiliary agent is one or more of Ni, Cu, Ca, Mg and K elements; the carrier content is 50-90%, the active component element content is 1-40%, and the auxiliary agent element content is 0.1-30%.
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