CN116948694A - High-silicon naphtha processing system and processing technology - Google Patents
High-silicon naphtha processing system and processing technology Download PDFInfo
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- 229910052710 silicon Inorganic materials 0.000 title claims abstract description 94
- 239000010703 silicon Substances 0.000 title claims abstract description 92
- 238000012545 processing Methods 0.000 title claims abstract description 25
- 238000005516 engineering process Methods 0.000 title claims abstract description 16
- 238000006243 chemical reaction Methods 0.000 claims abstract description 107
- 239000007788 liquid Substances 0.000 claims abstract description 85
- 239000007791 liquid phase Substances 0.000 claims abstract description 77
- XUIMIQQOPSSXEZ-UHFFFAOYSA-N Silicon Chemical compound [Si] XUIMIQQOPSSXEZ-UHFFFAOYSA-N 0.000 claims abstract description 69
- 239000002994 raw material Substances 0.000 claims abstract description 61
- 239000003921 oil Substances 0.000 claims abstract description 48
- 229910052739 hydrogen Inorganic materials 0.000 claims abstract description 37
- 239000001257 hydrogen Substances 0.000 claims abstract description 37
- 238000010574 gas phase reaction Methods 0.000 claims abstract description 36
- UFHFLCQGNIYNRP-UHFFFAOYSA-N Hydrogen Chemical compound [H][H] UFHFLCQGNIYNRP-UHFFFAOYSA-N 0.000 claims abstract description 33
- 238000004939 coking Methods 0.000 claims abstract description 25
- 150000001336 alkenes Chemical class 0.000 claims abstract description 22
- JRZJOMJEPLMPRA-UHFFFAOYSA-N olefin Natural products CCCCCCCC=C JRZJOMJEPLMPRA-UHFFFAOYSA-N 0.000 claims abstract description 22
- 238000000926 separation method Methods 0.000 claims abstract description 19
- 238000005194 fractionation Methods 0.000 claims abstract description 12
- 239000003054 catalyst Substances 0.000 claims description 60
- 238000005984 hydrogenation reaction Methods 0.000 claims description 58
- 238000000034 method Methods 0.000 claims description 52
- 239000012071 phase Substances 0.000 claims description 42
- 230000008569 process Effects 0.000 claims description 41
- 239000007789 gas Substances 0.000 claims description 35
- 229910044991 metal oxide Inorganic materials 0.000 claims description 24
- -1 cyclic siloxane compounds Chemical class 0.000 claims description 22
- 239000003502 gasoline Substances 0.000 claims description 19
- 239000003795 chemical substances by application Substances 0.000 claims description 18
- 150000001993 dienes Chemical class 0.000 claims description 17
- 239000000463 material Substances 0.000 claims description 14
- IJGRMHOSHXDMSA-UHFFFAOYSA-N Atomic nitrogen Chemical compound N#N IJGRMHOSHXDMSA-UHFFFAOYSA-N 0.000 claims description 12
- PNEYBMLMFCGWSK-UHFFFAOYSA-N aluminium oxide Inorganic materials [O-2].[O-2].[O-2].[Al+3].[Al+3] PNEYBMLMFCGWSK-UHFFFAOYSA-N 0.000 claims description 12
- 150000004706 metal oxides Chemical class 0.000 claims description 12
- 238000010438 heat treatment Methods 0.000 claims description 9
- 239000002184 metal Substances 0.000 claims description 9
- 238000002156 mixing Methods 0.000 claims description 7
- 239000002210 silicon-based material Substances 0.000 claims description 7
- 150000002431 hydrogen Chemical class 0.000 claims description 6
- 229910052757 nitrogen Inorganic materials 0.000 claims description 6
- 238000007670 refining Methods 0.000 claims description 6
- NINIDFKCEFEMDL-UHFFFAOYSA-N Sulfur Chemical compound [S] NINIDFKCEFEMDL-UHFFFAOYSA-N 0.000 claims description 5
- 239000011148 porous material Substances 0.000 claims description 5
- 229910052717 sulfur Inorganic materials 0.000 claims description 5
- 239000011593 sulfur Substances 0.000 claims description 5
- 238000004523 catalytic cracking Methods 0.000 claims description 4
- 230000003111 delayed effect Effects 0.000 claims description 4
- 230000000694 effects Effects 0.000 claims description 4
- 150000003961 organosilicon compounds Chemical class 0.000 claims description 4
- 238000004230 steam cracking Methods 0.000 claims description 4
- 229910000077 silane Inorganic materials 0.000 claims description 3
- 239000012752 auxiliary agent Substances 0.000 claims description 2
- 229910052698 phosphorus Inorganic materials 0.000 claims description 2
- 125000005575 polycyclic aromatic hydrocarbon group Chemical group 0.000 claims description 2
- 150000004756 silanes Chemical class 0.000 claims description 2
- 239000000126 substance Substances 0.000 abstract description 4
- 239000002283 diesel fuel Substances 0.000 description 9
- VGGSQFUCUMXWEO-UHFFFAOYSA-N Ethene Chemical compound C=C VGGSQFUCUMXWEO-UHFFFAOYSA-N 0.000 description 5
- 239000005977 Ethylene Substances 0.000 description 5
- 238000002407 reforming Methods 0.000 description 5
- 229910021332 silicide Inorganic materials 0.000 description 5
- FVBUAEGBCNSCDD-UHFFFAOYSA-N silicide(4-) Chemical compound [Si-4] FVBUAEGBCNSCDD-UHFFFAOYSA-N 0.000 description 5
- 238000005265 energy consumption Methods 0.000 description 4
- 238000009616 inductively coupled plasma Methods 0.000 description 4
- 150000005673 monoalkenes Chemical class 0.000 description 4
- AQRLNPVMDITEJU-UHFFFAOYSA-N triethylsilane Chemical compound CC[SiH](CC)CC AQRLNPVMDITEJU-UHFFFAOYSA-N 0.000 description 4
- 238000005336 cracking Methods 0.000 description 3
- 230000008021 deposition Effects 0.000 description 3
- OKTJSMMVPCPJKN-UHFFFAOYSA-N Carbon Chemical compound [C] OKTJSMMVPCPJKN-UHFFFAOYSA-N 0.000 description 2
- 238000005481 NMR spectroscopy Methods 0.000 description 2
- VYPSYNLAJGMNEJ-UHFFFAOYSA-N Silicium dioxide Chemical compound O=[Si]=O VYPSYNLAJGMNEJ-UHFFFAOYSA-N 0.000 description 2
- 150000001343 alkyl silanes Chemical class 0.000 description 2
- 229910052799 carbon Inorganic materials 0.000 description 2
- 230000008859 change Effects 0.000 description 2
- 230000000052 comparative effect Effects 0.000 description 2
- 150000001875 compounds Chemical class 0.000 description 2
- 238000006477 desulfuration reaction Methods 0.000 description 2
- 230000023556 desulfurization Effects 0.000 description 2
- JJQZDUKDJDQPMQ-UHFFFAOYSA-N dimethoxy(dimethyl)silane Chemical compound CO[Si](C)(C)OC JJQZDUKDJDQPMQ-UHFFFAOYSA-N 0.000 description 2
- 239000012535 impurity Substances 0.000 description 2
- 230000009467 reduction Effects 0.000 description 2
- 229920006395 saturated elastomer Polymers 0.000 description 2
- SCPYDCQAZCOKTP-UHFFFAOYSA-N silanol Chemical compound [SiH3]O SCPYDCQAZCOKTP-UHFFFAOYSA-N 0.000 description 2
- 150000003377 silicon compounds Chemical class 0.000 description 2
- CZDYPVPMEAXLPK-UHFFFAOYSA-N tetramethylsilane Chemical compound C[Si](C)(C)C CZDYPVPMEAXLPK-UHFFFAOYSA-N 0.000 description 2
- 238000004400 29Si cross polarisation magic angle spinning Methods 0.000 description 1
- YZCKVEUIGOORGS-OUBTZVSYSA-N Deuterium Chemical compound [2H] YZCKVEUIGOORGS-OUBTZVSYSA-N 0.000 description 1
- 239000004593 Epoxy Substances 0.000 description 1
- 239000006173 Good's buffer Substances 0.000 description 1
- BLRPTPMANUNPDV-UHFFFAOYSA-N Silane Chemical compound [SiH4] BLRPTPMANUNPDV-UHFFFAOYSA-N 0.000 description 1
- YXLZUVKNJMLMFM-UHFFFAOYSA-N [Si].[SiH3]O[SiH3] Chemical compound [Si].[SiH3]O[SiH3] YXLZUVKNJMLMFM-UHFFFAOYSA-N 0.000 description 1
- 230000009471 action Effects 0.000 description 1
- 230000002411 adverse Effects 0.000 description 1
- 125000000217 alkyl group Chemical group 0.000 description 1
- 238000004458 analytical method Methods 0.000 description 1
- 210000004556 brain Anatomy 0.000 description 1
- 239000013064 chemical raw material Substances 0.000 description 1
- DBKNGKYVNBJWHL-UHFFFAOYSA-N chloro-dimethyl-octylsilane Chemical compound CCCCCCCC[Si](C)(C)Cl DBKNGKYVNBJWHL-UHFFFAOYSA-N 0.000 description 1
- 230000008878 coupling Effects 0.000 description 1
- 230000001808 coupling effect Effects 0.000 description 1
- 238000010168 coupling process Methods 0.000 description 1
- 238000005859 coupling reaction Methods 0.000 description 1
- 239000010779 crude oil Substances 0.000 description 1
- 230000009849 deactivation Effects 0.000 description 1
- 230000007547 defect Effects 0.000 description 1
- 230000001934 delay Effects 0.000 description 1
- 238000013461 design Methods 0.000 description 1
- 238000010586 diagram Methods 0.000 description 1
- 125000000118 dimethyl group Chemical group [H]C([H])([H])* 0.000 description 1
- YYLGKUPAFFKGRQ-UHFFFAOYSA-N dimethyldiethoxysilane Chemical compound CCO[Si](C)(C)OCC YYLGKUPAFFKGRQ-UHFFFAOYSA-N 0.000 description 1
- 238000007599 discharging Methods 0.000 description 1
- 238000004134 energy conservation Methods 0.000 description 1
- 238000011156 evaluation Methods 0.000 description 1
- 239000012530 fluid Substances 0.000 description 1
- 239000000295 fuel oil Substances 0.000 description 1
- 229910052736 halogen Inorganic materials 0.000 description 1
- 150000002367 halogens Chemical class 0.000 description 1
- UQEAIHBTYFGYIE-UHFFFAOYSA-N hexamethyldisiloxane Chemical compound C[Si](C)(C)O[Si](C)(C)C UQEAIHBTYFGYIE-UHFFFAOYSA-N 0.000 description 1
- 238000001225 nuclear magnetic resonance method Methods 0.000 description 1
- CXQXSVUQTKDNFP-UHFFFAOYSA-N octamethyltrisiloxane Chemical compound C[Si](C)(C)O[Si](C)(C)O[Si](C)(C)C CXQXSVUQTKDNFP-UHFFFAOYSA-N 0.000 description 1
- MSRJTTSHWYDFIU-UHFFFAOYSA-N octyltriethoxysilane Chemical compound CCCCCCCC[Si](OCC)(OCC)OCC MSRJTTSHWYDFIU-UHFFFAOYSA-N 0.000 description 1
- 238000005192 partition Methods 0.000 description 1
- 231100000572 poisoning Toxicity 0.000 description 1
- 230000000607 poisoning effect Effects 0.000 description 1
- 238000011160 research Methods 0.000 description 1
- 239000000377 silicon dioxide Substances 0.000 description 1
- 238000001179 sorption measurement Methods 0.000 description 1
- 238000001228 spectrum Methods 0.000 description 1
- VCZQFJFZMMALHB-UHFFFAOYSA-N tetraethylsilane Chemical compound CC[Si](CC)(CC)CC VCZQFJFZMMALHB-UHFFFAOYSA-N 0.000 description 1
- 125000003698 tetramethyl group Chemical group [H]C([H])([H])* 0.000 description 1
- INPZSKMAWFGEOP-UHFFFAOYSA-N tetrapropylsilane Chemical compound CCC[Si](CCC)(CCC)CCC INPZSKMAWFGEOP-UHFFFAOYSA-N 0.000 description 1
- 231100000331 toxic Toxicity 0.000 description 1
- 230000002588 toxic effect Effects 0.000 description 1
- 238000012546 transfer Methods 0.000 description 1
- AAPLIUHOKVUFCC-UHFFFAOYSA-N trimethylsilanol Chemical compound C[Si](C)(C)O AAPLIUHOKVUFCC-UHFFFAOYSA-N 0.000 description 1
Classifications
-
- 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
- C10G65/00—Treatment of hydrocarbon oils by two or more hydrotreatment processes only
- C10G65/02—Treatment of hydrocarbon oils by two or more hydrotreatment processes only plural serial stages only
- C10G65/04—Treatment of hydrocarbon oils by two or more hydrotreatment processes only plural serial stages only including only refining steps
- C10G65/06—Treatment of hydrocarbon oils by two or more hydrotreatment processes only plural serial stages only including only refining steps at least one step being a selective hydrogenation of the diolefins
-
- 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
- C10G65/00—Treatment of hydrocarbon oils by two or more hydrotreatment processes only
- C10G65/02—Treatment of hydrocarbon oils by two or more hydrotreatment processes only plural serial stages only
- C10G65/04—Treatment of hydrocarbon oils by two or more hydrotreatment processes only plural serial stages only including only refining steps
- C10G65/08—Treatment of hydrocarbon oils by two or more hydrotreatment processes only plural serial stages only including only refining steps at least one step being a hydrogenation of the aromatic hydrocarbons
Landscapes
- 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 high-silicon naphtha processing system and a processing technology, wherein the high-silicon naphtha processing system comprises a fixed bed reactor, a desilication reactor, a heat exchanger and a fractionation unit; the fixed bed reactor is sequentially provided with a gas phase reaction zone, a gas-liquid separation zone, a gas-liquid countercurrent zone, a gas-liquid contact zone and a liquid phase reaction zone from top to bottom. The high silicon naphtha processing technology comprises the following steps: hydrogen firstly enters a gas-liquid contact zone at the lower part of the fixed bed reactor, high-silicon naphtha and olefin-rich raw materials enter a gas-liquid separation zone from the upper part of the fixed bed reactor, the raw materials react in each reaction zone, and the desilicated product is mixed with the effluent at the bottom of the liquid phase reaction zone after heat exchange, and then stripping fractionation is carried out, so that a qualified product is obtained. The invention not only realizes the high-efficiency removal of silicon in the high-silicon raw material, but also can simultaneously treat the conventional coking raw material, thereby obtaining qualified chemical raw materials and automotive diesel products, and simultaneously prolonging the operation period of the coking oil processing device.
Description
Technical Field
The invention belongs to the field of clean oil refining, and particularly relates to a high-silicon naphtha processing technology.
Background
At present, the hydrogenation process for coking oil products has the problem of short operation period of the device. On one hand, the raw materials contain cyclosiloxane silicon species, and although the cyclosiloxane silicon species can be removed under certain conditions, silicon is easy to deposit to cause catalyst poisoning, on the other hand, the content of olefin in coked oil products is too high, and the violent heat release in the reaction process causes coking and carbon deposition on the surface of the catalyst and causes bed pressure drop. The two reasons cause that the operation period of the hydrogenation device of the coking oil products is obviously lower than that of other oil product hydrogenation devices.
With the surplus capacity of the fuel oil market in China, oil refining enterprises tend to convert naphtha into ethylene cracking raw materials and reforming raw materials through hydrogenation technology, and more values are created. The naphtha fraction is light and has a low impurity content, and generally has little processing difficulty. The conventional catalyst is adopted, and the requirements of reforming feeding and ethylene cracking raw materials can be met under certain process conditions. However, because of the complex sources of naphtha and the complex types of impurities therein, particularly the complex types of silicon-containing compounds, severe desilication conditions are required, which tend to adversely affect other components in the oil, and the current technology has difficulty in converting such difficult-to-remove, high-content silicon-containing oil products into acceptable reforming or ethylene feeds.
CN201911212127.9 discloses a grading method of a hydrofining catalyst of coker gasoline and a hydrofining method. Comprising the following steps: the method comprises the steps that a replaceable zone and a main reaction zone which are connected in series are sequentially arranged along the material flow direction, wherein the replaceable zone is sequentially filled with a diolefin saturated catalyst in a first reactor along the material flow direction, a second reactor is filled with a silicon catching agent, and the second reactor is a parallel double reactor which can be switched on line; the main reaction zone is sequentially filled with the dearsenization agent and the gasoline hydrogenation catalyst along the material flow direction, and the invention solves the problem of catalyst coking.
CN200810166851.8 discloses a method for prolonging the service life of a hydrotreating catalyst for coker gasoline and diesel. After being mixed at the inlet of the reactor, the coked gasoline and diesel oil and hydrogen enter a hydrogenation reactor from the upper part of the reactor and sequentially enter two reaction areas, wherein the first reaction area is filled with a desilication catalyst with high desilication activity and high Rong Gui capacity, and the second reaction area is filled with a conventional hydrotreating catalyst. Silicon in the coked gasoline and diesel oil is removed and deposited on the desilication catalyst before entering the conventional hydrogenation catalyst, so that the hydrotreating catalyst is protected from being polluted by silicon. The invention uses graded filling silicon catching agent to prolong the service life of hydrogenation main catalyst.
The technology of the patent is applied to the naphtha processing technology containing high content of difficult-to-remove silicon-containing compounds, and the obtained products can not meet the problems of high-quality reforming feed and ethylene cracking raw materials.
In view of the above, the problem of difficult processing of conventional coker oils and oils such as high silica brains is still to be solved, so as to provide more suitable reforming and ethylene raw materials for refineries.
Disclosure of Invention
Aiming at the defects of the prior art, the invention provides a high-silicon naphtha processing system and a processing technology, which match the characteristics of raw materials with the technological process, realize the high-efficiency removal of silicon in the high-silicon raw materials by virtue of zonal reaction control and fully utilize reaction heat, simultaneously process the conventional coking raw materials, obtain qualified chemical raw materials and automotive diesel products, and prolong the operation period of a coking oil processing device.
The existing naphtha desilication technology generally aims at desilication treatment of oil products with silicon content less than 100ppm and existing in the form of epoxy silane, and the silicon species are easier to remove under proper temperature and pressure. However, in recent years, as the quality of crude oil changes, the properties of naphtha raw materials also change, and the silicon content in some naphtha raw materials is up to 300-1000ppm, which is greatly beyond the conventional silicon content range, and the inventor finds that the types of silicon compounds therein also change, and more silicon compounds in alkylsilanes, silanol or silicon ethers are contained, and the silicon species are enriched in naphtha light components. After the raw oil is treated according to the desilication method and the process in the prior art, the silicon content is difficult to reach the standard.
The high-silicon naphtha processing system comprises a fixed bed reactor, a desilication reactor, a heat exchanger and a fractionation unit;
the fixed bed reactor is sequentially provided with a gas phase reaction zone, a gas-liquid separation zone, a gas-liquid countercurrent zone, a gas-liquid contact zone and a liquid phase reaction zone from top to bottom;
the gas-liquid separation zone is used for converting the materials entering the zone into gas-phase raw materials and liquid-phase raw materials; wherein the gas phase raw material enters a gas phase reaction zone, and the liquid phase raw material enters a gas-liquid countercurrent zone; the gas phase component is generally a fraction below 80 ℃, and the liquid phase component is generally a fraction above 80 ℃; the material comprises a high silicon naphtha raw material and an olefin-rich raw material in a certain proportion;
the gas phase reaction zone is used for carrying out low-temperature hydrogenation diene removal reaction on gas phase components from the gas-liquid separation zone;
the gas-liquid countercurrent zone is used for further hydrofining reaction of liquid phase raw materials and generally comprises cyclosiloxane silicide removal, hydrogenation saturation of mono-olefin and removal reaction of sulfur and nitrogen in naphtha fraction;
the gas-liquid contact zone is used for carrying out contact heat exchange on the hydrogen entering the zone and the liquid-phase effluent from the gas-liquid countercurrent zone, and mixing and dispersing the liquid-phase effluent and the hydrogen to form a hydrogen-dissolved stream; the gas-liquid contact area is communicated with a hydrogen source, and a mixed hydrogen dissolving component or equipment such as a membrane-tube type hydrogen dissolving component, a high-efficiency hydrogen mixer, a micro-bubble generator, a bubble fractal device and the like is generally arranged in the gas-liquid contact area;
the liquid phase reaction zone is used for carrying out deep hydrofining reaction on the hydrogen-dissolved material flow flowing out of the gas-liquid contact zone, and further carrying out deep desulfurization and dearomatization reaction;
the heat exchanger is used for exchanging heat between the bottom effluent of the liquid phase reaction zone and the gas phase component from the gas phase reaction zone for removing the diolefin;
the desilication reactor is used for desilication reaction of removing gas phase components of the diene after heat exchange;
the fractionating unit is used for stripping and fractionating the gas-phase effluent of the desilication reactor and the bottom effluent of the liquid-phase reaction zone after heat exchange; typically in the form of a fractionation column.
In the system of the invention, the top of the gas phase reaction zone is provided with a gas phase outlet, the bottom of the liquid phase reaction zone is provided with a liquid phase outlet, and effluent at the bottom of the liquid phase reaction zone is connected with an outlet pipeline of the desilication reactor after passing through a heat exchanger and then enters a fractionation unit together.
The invention relates to a high silicon naphtha processing technology, which comprises the following steps:
(1) The hydrogen firstly enters a gas-liquid contact zone at the lower part of the fixed bed reactor, is contacted and mixed with a liquid phase product flowing out of the bottom of a gas-liquid countercurrent zone, flows upwards and sequentially passes through the gas-liquid countercurrent zone, enters a gas phase reaction zone together with a gas phase component separated from a gas-liquid separation zone, and then flows out of the top of the fixed bed reactor along with a gas phase component from which diolefins are removed;
(2) High silicon naphtha and olefin-rich raw materials enter a gas-liquid separation zone from the upper part of a fixed bed reactor and are converted into gas-phase raw materials and liquid-phase raw materials under certain conditions; the gas phase raw material upwards enters a gas phase reaction zone to carry out hydrogenation and diene removal reaction; the liquid phase raw material enters into a gas-liquid countercurrent zone downwards to be in countercurrent contact with upward hydrogen to carry out hydrogenation refining reaction, liquid phase effluent obtained by hydrogenation enters into a gas-liquid contact zone to be in contact heat exchange with hydrogen, the obtained hydrogen-dissolved material flow enters into a liquid phase reaction zone to carry out deep hydrogenation refining reaction, gas phase components flowing out of the bottom effluent of the liquid phase reaction zone and the top of a fixed bed reactor are subjected to heat exchange by a heat exchanger, the gas phase components enter into a desilication reactor to carry out desilication, and the desilication product is mixed with the bottom effluent of the liquid phase reaction zone after heat exchange to carry out steam stripping fractionation to obtain a qualified product.
In the process, the hydrogen and the oil raw materials are fed at the medium temperature (100-150 ℃) of a heating furnace during starting, and the heating furnace is stopped after the bed layer is obviously heated, and the hydrogen and the oil raw materials are fed at the normal temperature (5-30 ℃).
In the process of the invention, the silicon content of the high silicon naphtha is 1-3000 mug/g, preferably 5-800 mug/g, the sulfur content is no more than 1000 mug/g, the nitrogen content is no more than 200 mug/g, the olefin content is 1-30 wt%, preferably 2-15 wt%,typically, the catalyst is at least one of straight run naphtha, coker naphtha, hydroupgraded naphtha, and the like; the high silicon naphtha contains at least one silicon-containing compound except cyclic siloxane compounds and silane compounds; wherein the cyclic siloxane compound is a cyclic siloxane bond (-Si-O-) n Organosilicon compounds having a main chain of n.gtoreq.3, e.g. of the formula (H) 2 SiO) n A compound in which n is not less than 3, or a compound in which H in the general formula is substituted with an alkyl group, a halogen or the like; the silane compound is Si-containing n H 2n+2 An organosilicon compound of the general formula n.gtoreq.1.
Further, the silicon-containing compounds in the high silicon naphtha include, but are not limited to, at least one of alkylsilanes, silanol, or silyl ether silicon-containing compounds.
Further, as a specific embodiment, the silicon-containing compound in the raw oil includes, but is not limited to, at least one of tetramethylsilane, triethylsilane, tetraethylsilane, tetrapropylsilane, trimethylsilanol, dimethoxydimethylsilane, diethoxydimethylsilane, hexamethyldisiloxane, dimethyloctylchlorosilane, octamethyltrisiloxane, and n-octyltriethoxysilane.
In the process of the present invention, the properties of the olefin-rich feedstock are as follows: the content of olefin is 5-60wt%, preferably 10-40wt%, the sulfur content is no more than 15000 mug/g, the nitrogen content is no more than 400 mug/g, and the polycyclic aromatic hydrocarbon content is no more than 50wt%. The olefin-rich feedstock generally comprises at least one gasoline or diesel fraction obtained by catalytic cracking, steam cracking, delayed coking, ebullated bed residuum hydrogenation, etc., and one or more of gasoline or diesel fractions obtained by catalytic cracking, steam cracking, delayed coking, ebullated bed residuum hydrogenation, etc., preferably the coker gasoline is present in an amount of not less than 60wt%.
In the process of the present invention, the proportion of the olefin-rich raw material to the total raw material (high silicon naphtha and olefin-rich raw material) is not less than 70 wt%, preferably 40 wt% to 60wt%.
In the process of the invention, the operation conditions of the gas-liquid separation zone are as follows: the pressure is 1.0-10.0 MPa, preferably 2.0-6.0 MPa, wherein the hydrogen partial pressure accounts for 45-80% of the total pressure; the feeding temperature is 50-240 ℃, preferably 100-150 ℃.
In the process of the invention, the operating conditions of the gas phase reaction zone are as follows: the pressure is 0.1-6.0 MPa, preferably 0.5-3.0 MPa, wherein the hydrogen partial pressure accounts for 40-70% of the total pressure; volume space velocity is 0.1-6.0 h -1 Preferably 0.5 to 3.0 hours -1 The method comprises the steps of carrying out a first treatment on the surface of the The reaction temperature is 80 to 200 ℃, preferably 120 to 160 ℃. The reaction zone is filled with a light distillate oil hydrogenation catalyst with low-temperature hydrogenation activity, wherein the active metal is VIB group metal oxide and/or VIII group metal oxide, the carrier is alumina or modified alumina, the content of the VIB group metal oxide is 5% -30%, preferably 5% -15%, and the content of the VIII group metal oxide is 1% -15%, preferably 2% -6% based on the weight of the catalyst; the specific surface area is 100-500m 2 /g, preferably 300-500m 2 Per gram, the pore volume is from 0.3 to 1.2mL/g, preferably from 0.4 to 0.8mL/g. Taking the total loading of the catalyst in the fixed bed reactor as a reference, the loading volume ratio of the catalyst in the gas phase reaction zone is 1% -50%, preferably 5% -35%;
the gas phase reaction zone is used for low-temperature diolefin removal of gas phase components from the gas-liquid separation zone, and at the moment, acyclic siloxanes in the high-silicon naphtha also enter the gas phase reaction zone due to lighter fractions, but the inventor researches find that the removal of the acyclic siloxanes needs higher reaction temperature, so that the silicide does not react in the gas phase reaction zone.
In the process of the invention, the operating conditions of the gas-liquid countercurrent zone are generally as follows: 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 5.0 hours -1 The method comprises the steps of carrying out a first treatment on the surface of the The reaction temperature is 200 to 400 ℃, preferably 220 to 300 ℃.
Silicon catching agent and light distillate oil hydrogenation catalyst are filled in the gas-liquid countercurrent zone in grading mode. The loading volume ratio of the two catalysts is 1:5~3:1. the active metal of the silicon catching agent is VIB group metal oxide or/and VIII group metal oxide, the carrier is alumina or alumina modified by an auxiliary agent, the content of the VIB group metal oxide is 5% -30%, preferably 5% -15%, and the content of the VIII group metal oxide is 1% -15%, preferably 2% -6% based on the weight of the silicon catching agent; comparison meterThe area is 100-500m 2 /g, preferably 300-500m 2 Per gram, the pore volume is from 0.3 to 1.2mL/g, preferably from 0.4 to 0.8mL/g. The light distillate hydrogenation catalyst is well known to those skilled in the art, and the active metal content can be adjusted according to actual needs, as the type of the catalyst in the gas phase reaction zone is the same. The filling volume ratio of the catalyst in the gas-liquid countercurrent zone is 1% -60%, preferably 5% -40%;
in the process, the operating condition of the gas-liquid contact zone is that the pressure is 3.0-10.0 MPa, preferably 5.0-8.0 MPa, wherein the hydrogen partial pressure accounts for 100% of the total pressure ratio; the feeding temperature is 100-300 ℃ at the initial stage of operation, preferably 160-260 ℃. After the device runs stably, normal-temperature feeding can be switched.
In the process of the present invention, the operating conditions of the liquid phase reaction zone are generally: the pressure is 1.0-8.0 MPa, preferably 3.0-7.0 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 0.1-6.0 h -1 Preferably 0.5 to 3.0 hours -1 The method comprises the steps of carrying out a first treatment on the surface of the The reaction temperature is 260 to 400 ℃, preferably 300 to 360 ℃. The catalyst loading volume ratio in the liquid phase reaction zone is 1% -80%, preferably 30% -60%. The liquid phase reaction zone is filled with a diesel hydrofining catalyst, which is well known to the person skilled in the art, wherein the active metal is a VIB group metal oxide or/and a VIII group metal oxide, the content of the VIB group metal oxide is 5% -30%, preferably 15% -25%, and the content of the VIII group metal oxide is 1% -15%, preferably 3% -8% based on the weight of the hydrogenation catalyst; the carrier is alumina or assistant modified alumina, the assistant is one or more of B, P, mg, zr or Si, and the weight of the assistant is 3-15%, preferably 3-10% of the weight of the carrier.
In the process, the temperature of the gas phase component subjected to heat exchange with the effluent of the liquid phase reaction zone is 200-340 ℃, and then the gas phase component enters a desilication reactor, wherein a silicon capturing agent is filled in the desilication reactor, and the type of the silicon capturing agent is the same as that of a gas-liquid countercurrent zone, and is not described again. The operating conditions of the desilication reactor are generally: the pressure is 0.2-4.0 MPa, preferably 1.0-3.0 MPa, and the volume space velocity is 0.1-4.0 h -1 Preferably 0.2 to 2.0 hours -1 The method comprises the steps of carrying out a first treatment on the surface of the The reaction temperature is 150 to 400 ℃, preferably 220 to 300 ℃.
Further, according to the silicon content in the high silicon naphtha oil, if the Si content is 500ppm or more, it is preferable to set two desilication reactors to be operated in an on-line switching.
The system and the process of the invention design a low-energy-consumption and low-cost process line aiming at the problems of difficult desilication of high-silicon naphtha and limited long-period operation of coking devices, can meet the processing of inferior oil products, and provide qualified chemical raw materials and automotive diesel products, and have the following advantages:
(1) According to the invention, by the combined action of the raw material feeding mode, the gas-liquid separation zone and the gas-liquid countercurrent zone, the diolefin and high-silicon substances in the raw material are enriched in the gas-phase reaction zone to carry out independent reaction, and the diolefin and the monoolefin are separated by cutting fractions, so that chain reaction caused by massive saturation of the monoolefin after the diolefin hydrogenation is exothermic is avoided, and the control of the temperature of a catalyst bed layer is facilitated. In addition, compared with the conventional coking oil hydrogenation process, the device for removing the diolefin by the pre-protection reactor can be omitted, and the process flow is simplified.
(2) By mixing high silicon naphtha with coker oil (an olefin-rich feedstock). Because the silicide in the high-silicon naphtha needs higher reaction temperature to be removed, the independent hydrogenation reaction condition is harsh and the energy consumption is high. By processing the oil with the coking oil, the saturated reaction heat of olefin in the coking oil can be fully utilized, so that the energy is saved and the consumption is reduced. However, if the conventional hydrogenation technology is used, the bed temperature rise of the coking oil with larger hydrogenation mainly occurs in the middle and lower parts of the reactor, and if the high-silicon species can be removed only in the middle and lower parts of the reactor, the high-silicon species can cover the surface of the main catalyst, so that the high-silicon species are quickly poisoned and deactivated. The invention utilizes the characteristic that the high silicon component is enriched in the light component, effectively separates the high silicon component through the arrangement of partition reaction, then increases the reaction temperature through heat exchange with the heavy component, and enters the gas phase reactor for independent hydrogenation, thereby utilizing the hydrogenation reaction heat of the coking oil, avoiding the toxic action on the main catalyst caused by the common reaction with the coking oil, and realizing the prolonging of the operation period of the device. Meanwhile, the heating furnace is only required to be started for heating the feed in the initial start-up period, and the heating furnace can be stopped after the temperature of the catalyst bed layer is obviously increased, so that the energy is further saved and the consumption is reduced.
(3) The invention can better distribute and balance the heat of the three reaction areas by arranging the gas-liquid countercurrent areas, thereby realizing coupling. The invention mainly generates a great deal of saturation of the mono-olefin in the gas-liquid countercurrent zone, is also a reactor with most obvious concentrated heat release, can rapidly carry a great deal of reaction heat to the gas-phase reaction zone and the liquid-phase reaction zone through the countercurrent flow of the gas phase and the liquid phase in the zone, meets the reaction temperature requirements of the two reaction zones, can better control the temperature of the catalyst bed layer in the gas-liquid countercurrent zone, and delays the coking deactivation of the surface of the catalyst. Meanwhile, the three reaction areas provided by the invention have the coupling effect, so that the reaction system is more stable. The gas-liquid countercurrent zone has the advantages that stable pressure control is required along with back mixing while the gas-liquid countercurrent zone is in countercurrent contact with the gas-liquid to strengthen the mass transfer process, the uppermost gas-phase reaction zone has a larger compressible gas-phase space, a good buffer effect is achieved on the stable bed pressure and the stable fluid flow state, the gas-phase flow rate and the liquid layer thickness of the gas-liquid countercurrent zone can be adjusted by flexibly controlling the outlet gas quantity, and the gas-speed range in which flooding occurs is improved. The effluent of the gas-liquid countercurrent zone is contacted with pure hydrogen through the gas-liquid contact zone, then is subjected to heat exchange and temperature reduction, and quickly enters the liquid phase reaction zone below, so that the hydrogen gas is prevented from escaping after the hydrogen mixing process of the conventional liquid phase hydrogenation process enters the reactor, meanwhile, the liquid phase reaction zone below is a liquid phase space, the physical distribution state of the outlet of the reactor can be well controlled, if the effluent does not have the area, the problem that the hydrogen gas is carried out of the reactor without reacting through the catalyst bed layer can be caused, a high-pressure separator in the conventional process flow can be omitted, and the flow is simplified.
Drawings
FIG. 1 is a schematic diagram of the high silicon naphtha processing process flow of the present invention.
In the figure: 1-raw materials, 2-hydrogen, 3-gas phase reaction zone, 4-gas-liquid separation zone, 5-gas-liquid countercurrent zone, 6-gas-liquid mixing zone, 7-liquid phase reaction zone, 8-gas phase reaction zone effluent, 9-liquid phase reaction zone effluent, 10-heat exchanger, 11-desilication reactor and 12-fractionation unit.
Detailed Description
The following non-limiting examples will enable those of ordinary skill in the art to more fully understand the invention and are not intended to limit the invention in any way.
The catalyst composition provided by the invention can be characterized by combining inductively coupled plasma ICP and XRF energy spectra, and the total content of the VIB group metal and the total content of the VIII group metal in the catalyst are characterized. The specific surface area and pore volume of the catalyst provided by the invention are analyzed by a nitrogen physical adsorption method. The silicon content and type in the raw oil and the product in the present invention are analyzed by ICP and nuclear magnetic resonance methods, wherein the silicon content in the raw oil and the product is analyzed by ICP, and the silicon type in the raw oil and the product is analyzed by nuclear magnetic resonance 29Si MAS NMR. The silicon type in the high silicon naphtha is tetramethyl silane, dimethoxy dimethyl silane and triethyl silane through nuclear magnetic resonance analysis; the silicon types in coker oils (olefin-rich feedstock) are dimethyl, trimethyl, tetramethyl, and the like.
Taking fig. 1 as an example, the implementation process of the processing technology of the high silicon naphtha of the invention is as follows: the reaction raw material 1 enters a hydrogenation reactor from a gas-liquid separation zone 4. In the gas-liquid separation zone 4, the gas phase and the liquid phase are separated. The gas phase flows upward into the gas phase reaction zone 3 and the liquid phase flows downward into the gas-liquid countercurrent zone 5. The hydrogen 2 enters the hydrogenation reactor in the gas-liquid mixing zone 6, and after being mixed and contacted with the liquid-phase material flowing downwards in the gas-liquid countercurrent zone 5, flows upwards to enter the gas-liquid countercurrent zone 5, and the liquid-phase material carrying the hydrogen flows downwards to enter the liquid-phase reaction zone 7.
The gas phase reaction takes place in the gas phase reaction zone 3, mainly with saturation of diolefins in the fraction below 80 ℃ to produce a gas phase reaction zone effluent 8. In the gas-liquid countercurrent zone 5, gas-liquid two-phase reaction occurs, the liquid phase is naphtha heavy component and diesel fraction flows downwards, the gas phase is hydrogen gas flows upwards, and the gas-liquid countercurrent contact occurs hydrodesulfurization and desilication reaction of naphtha fraction. The liquid phase material flow after the reaction in the gas-liquid countercurrent zone 5 flows downwards to enter a gas-liquid contact zone 6, and enters a liquid phase reaction zone 7 after contacting and exchanging heat with hydrogen, wherein the liquid phase reaction zone 7 is in liquid phase reaction, deep desulfurization and dearomatization of diesel oil fraction mainly occurs, and the effluent 9 of the liquid phase reaction zone flows out of the device, exchanges heat with the effluent 8 of the gas phase reaction zone through a heat exchanger 10, then enters a desilication reactor 11, and enters a subsequent fractionation unit 12 together with the effluent 9 of the liquid phase reaction zone after desilication, so that qualified chemical raw material naphtha and automotive diesel oil are obtained.
Examples 1 to 3
In the process flow of fig. 1, a 100mL fixed bed hydrogenation reactor, a heat exchanger, a desilication reactor and a fractionation unit are provided, respectively. Wherein, the 100mL fixed bed hydrogenation reactor is provided with a gas phase reaction zone, a gas-liquid countercurrent zone and a liquid phase reaction zone from top to bottom, and a catalyst bed layer is arranged in each of the gas phase reaction zone, the gas-liquid countercurrent zone and the liquid phase reaction zone. 20mL of light distillate oil hydrogenation catalyst A is filled in the gas phase reaction zone, a silicon capturing agent B and the light distillate oil hydrogenation catalyst A are filled in a gas-liquid countercurrent zone in a grading manner, the filling volumes are respectively 10mL and 20mL, and 50mL of diesel oil hydrofining catalyst C is filled in the liquid phase reaction zone. The desilication reactor was filled with 50mL of silicon trap B. The method adopts high silicon naphtha and coker gasoline and diesel mixed oil as raw materials, wherein the mass ratio of the high silicon naphtha to the coker gasoline to the coker diesel is 4:4:2. the catalyst properties are shown in Table 1, the raw oil properties are shown in Table 2, and the reaction conditions and results are shown in Table 3.
Comparative example 1
The conventional coking gasoline and diesel hydrogenation process is adopted, and a pre-protection reactor and a main reactor are respectively arranged in the process flow. The pre-protection reactor is filled with 20mL of light distillate oil hydrogenation catalyst A, and the main reactor is filled with 10mL of silicon catching agent B, 20mL of light distillate oil hydrogenation catalyst A and 50mL of diesel oil hydrogenation catalyst C in a grading manner. The raw material properties are the same as in example, the reaction conditions of the pre-protection reactor are the same as in example 3, the reaction conditions of the main reactor are the same as in example 3, the reaction conditions of the liquid phase reaction zone are the same as in example 3.
Comparative example 2
The high silicon naphtha and coked gasoline and diesel oil are adopted for hydrogenation respectively.
The high silicon naphtha is subjected to hydrodesiliconizing in a gas phase reactor, and 50mL of silicon catching agent B and 20mL of light distillate hydrogenation catalyst A are filled in the reactor. The high silicon naphtha raw material is heated to the temperature required by the reaction by adopting a heating furnace.
The hydrogenation of the coking gasoline and diesel adopts a conventional process, and a pre-protection reactor and a main reactor are respectively arranged in the process flow. The pre-protection reactor is filled with 20mL of light distillate oil hydrogenation catalyst A, and the main reactor is filled with 10mL of silicon catching agent B and 50mL of diesel oil hydrogenation catalyst C. The raw material properties were the same as in example, and the reaction conditions were the same as in example 3 in the liquid phase reaction zone.
TABLE 1 oil Properties of raw materials
TABLE 2 catalyst Properties
TABLE 3 Process conditions and results
The evaluation results in table 3 show that the high-silicon naphtha processing technology of the invention can simultaneously process the difficult-to-remove high-content silicide, control the silicon content in the refined naphtha to be less than 1.0 mug/g, better control and utilize the reaction heat, and has simple flow, energy conservation and consumption reduction. Meanwhile, the device of the invention has no obvious silicon deposition phenomenon during discharging agent, and no coking and carbon deposition phenomenon, which proves that the device can be operated for a long period. When the conventional coking gasoline and diesel hydrogenation process is adopted to process naphtha containing high silicon content, silicide can be effectively removed at the middle and lower parts of the reactor, so that the main catalyst is rapidly deactivated, and the refined product is difficult to reach the standard. The high-silicon naphtha and the coker gasoline and diesel oil are adopted for hydrogenation respectively, a heating furnace is required to be arranged in the high-silicon naphtha desilication reactor, the energy consumption is high, and the reaction heat of the coker oil cannot be fully utilized; the coking gasoline and diesel hydrogenation process needs to be provided with a pre-protection reactor, and the problem of centralized heat release due to olefin saturation is difficult to solve, and the problem of short operation period still exists.
Claims (17)
1. The high-silicon naphtha processing system is characterized by comprising a fixed bed reactor, a desilication reactor, a heat exchanger and a fractionation unit; the fixed bed reactor is sequentially provided with a gas phase reaction zone, a gas-liquid separation zone, a gas-liquid countercurrent zone, a gas-liquid contact zone and a liquid phase reaction zone from top to bottom;
the gas-liquid separation zone is used for converting the materials entering the zone into gas-phase raw materials and liquid-phase raw materials;
the gas phase reaction zone is used for carrying out low-temperature hydrogenation diene removal reaction on gas phase components from the gas-liquid separation zone;
the gas-liquid countercurrent zone is used for carrying out further hydrofining reaction on the liquid phase raw material;
the gas-liquid contact zone is used for carrying out contact heat exchange on the hydrogen entering the zone and the liquid-phase effluent from the gas-liquid countercurrent zone, and mixing and dispersing the liquid-phase effluent and the hydrogen to form a hydrogen-dissolved stream;
the liquid phase reaction zone is used for carrying out deep hydrofining reaction on the hydrogen-dissolved material flow flowing out of the gas-liquid contact zone;
the heat exchanger is used for exchanging heat between the bottom effluent of the liquid phase reaction zone and the gas phase component from the gas phase reaction zone for removing the diolefin;
the desilication reactor is used for desilication reaction of removing gas phase components of the diene after heat exchange;
the fractionating unit is used for stripping and fractionating the gas phase effluent of the desilication reactor and the bottom effluent of the liquid phase reaction zone after heat exchange.
2. The system according to claim 1, wherein: the top of the gas phase reaction zone is provided with a gas phase outlet, the bottom of the liquid phase reaction zone is provided with a liquid phase outlet, and effluent at the bottom of the liquid phase reaction zone is connected with an outlet pipeline of the desilication reactor after passing through a heat exchanger and then enters the fractionation unit together.
3. The system according to claim 1, wherein: the gas phase raw material enters a gas phase reaction zone, and the liquid phase raw material enters a gas-liquid countercurrent zone; the gas phase component is a fraction below 80 ℃, and the liquid phase component is a fraction above 80 ℃; the material comprises a certain proportion of high silicon naphtha raw material and olefin-rich raw material.
4. The system according to claim 1, wherein: the gas-liquid contact area is communicated with a hydrogen source, and a mixed hydrogen dissolving component or equipment is arranged in the gas-liquid contact area.
5. The high silicon naphtha processing technology is characterized by comprising the following steps: (1) The hydrogen firstly enters a gas-liquid contact zone at the lower part of the fixed bed reactor, is contacted and mixed with a liquid phase product flowing out of the bottom of a gas-liquid countercurrent zone, flows upwards and sequentially passes through the gas-liquid countercurrent zone, enters a gas phase reaction zone together with a gas phase component separated from a gas-liquid separation zone, and then flows out of the top of the fixed bed reactor along with a gas phase component from which diolefins are removed; (2) High silicon naphtha and olefin-rich raw materials enter a gas-liquid separation zone from the upper part of a fixed bed reactor and are converted into gas-phase raw materials and liquid-phase raw materials under certain conditions; the gas phase raw material upwards enters a gas phase reaction zone to carry out hydrogenation and diene removal reaction; the liquid phase raw material enters into a gas-liquid countercurrent zone downwards to be in countercurrent contact with upward hydrogen to carry out hydrogenation refining reaction, liquid phase effluent obtained by hydrogenation enters into a gas-liquid contact zone to be in contact heat exchange with hydrogen, the obtained hydrogen-dissolved material flow enters into a liquid phase reaction zone to carry out deep hydrogenation refining reaction, gas phase components flowing out of the bottom effluent of the liquid phase reaction zone and the top of a fixed bed reactor are subjected to heat exchange by a heat exchanger, the gas phase components enter into a desilication reactor to carry out desilication, and the desilication product is mixed with the bottom effluent of the liquid phase reaction zone after heat exchange to carry out steam stripping fractionation to obtain a qualified product.
6. The process according to claim 5, wherein: the hydrogen and oil raw materials are fed at the medium temperature of 100-150 ℃ in a heating furnace during starting, the heating furnace is stopped after the bed layer has obvious temperature rise, and the hydrogen and oil raw materials are fed at the normal temperature of 5-30 ℃.
7. The process according to claim 5, wherein: the silicon content of the high silicon naphtha is 1-3000 mu g/g, and the sulfur content is no more than1000 mug/g, nitrogen content not more than 200 mug/g, olefin content 1-30wt%, preferably 2-15wt%; the high silicon naphtha contains at least one silicon-containing compound except cyclic siloxane compounds and silane compounds; wherein the cyclic siloxane compound is a cyclic siloxane bond (-Si-O-) n An organosilicon compound wherein n is not less than 3 as a main chain; the silane compound is Si-containing n H 2n+2 An organosilicon compound of the general formula n.gtoreq.1.
8. The process according to claim 7, wherein: the high silicon naphtha is at least one of straight run naphtha, coker naphtha or hydro-modified naphtha.
9. The process according to claim 5, wherein: the properties of the olefin-rich feedstock are as follows: the content of olefin is 5-60wt%, preferably 10-40wt%, the sulfur content is no more than 15000 mug/g, the nitrogen content is no more than 400 mug/g, and the polycyclic aromatic hydrocarbon content is no more than 50wt%; the olefin-rich feedstock comprises at least one gasoline and diesel fraction obtained by catalytic cracking, steam cracking, delayed coking, ebullated bed residuum hydrogenation processes, and one or more of gasoline or diesel fractions obtained by catalytic cracking, steam cracking, delayed coking, ebullated bed residuum hydrogenation processes, preferably the coker gasoline content is not less than 60wt%.
10. The process according to claim 5, wherein: the proportion of the olefin-rich raw material to the total raw material (high silicon naphtha and olefin-rich raw material) is not less than 70 wt%, preferably 40 wt% to 60wt%.
11. The process according to claim 5, wherein: the operating conditions of the gas-liquid separation zone are as follows: the pressure is 1.0-10.0 MPa, wherein the hydrogen partial pressure accounts for 45% -80% of the total pressure proportion; the feeding temperature is 50-240 ℃, preferably 100-150 ℃.
12. The process according to claim 5, wherein: the saidThe operating conditions of the gas phase reaction zone are as follows: the pressure is 0.1-6.0 MPa, wherein the hydrogen partial pressure accounts for 40-70% of the total pressure proportion, and the volume space velocity is 0.1-6.0 h -1 The reaction temperature is 80-200 ℃; the reaction zone is filled with a light distillate oil hydrogenation catalyst with low-temperature hydrogenation activity, wherein active metal is VIB group metal oxide and/or VIII group metal oxide, a carrier is alumina or modified alumina, the weight of the catalyst is taken as a reference, the content of the VIB group metal oxide is 5% -30%, and the content of the VIII group metal oxide is 1% -15%; the specific surface area is 100-500m 2 Per gram, pore volume of 0.3-1.2mL/g; the total loading of the catalyst in the fixed bed reactor is taken as a reference, and the loading volume ratio of the catalyst in the gas phase reaction zone is 1% -50%.
13. The process according to claim 5, wherein: the operating conditions of the gas-liquid countercurrent 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, and the volume space velocity is 0.1-10.0 h -1 The reaction temperature is 200-400 ℃; the gas-liquid countercurrent zone is internally graded and filled with a silicon catching agent and a light distillate hydrogenation catalyst; the loading volume ratio of the two catalysts is 1:5~3:1, a step of; the active metal of the silicon catching agent is VIB group metal oxide and/or VIII group metal oxide, the carrier is alumina or alumina modified by auxiliary agent, the weight of the silicon catching agent is taken as the reference, the content of the VIB group metal oxide is 5-30%, the content of the VIII group metal oxide is 1-15%, and the specific surface area is 100-500m 2 And/g, wherein the pore volume is 0.3-1.2mL/g, and the catalyst filling volume ratio in the gas-liquid countercurrent zone is 1% -60%.
14. The process according to claim 5, wherein: the operating condition of the gas-liquid contact zone is that the pressure is 3.0-10.0 MPa, wherein the hydrogen partial pressure accounts for 100% of the total pressure ratio; the feeding temperature is 50-300 ℃ in the initial stage of start-up, and normal-temperature feeding is switched after the device stably operates.
15. The process according to claim 5, wherein: the operating conditions of the liquid phase reaction zone are as follows: the pressure is 1.0-8.0 MPa, which is a pure liquid phase reaction zone,the volume ratio of standard hydrogen to oil is 2-50, and the volume space velocity is 0.1-6.0 h -1 The reaction temperature is 260-400 ℃; the catalyst filling volume ratio in the liquid phase reaction zone is 1% -80%, the diesel hydrofining catalyst is filled in the liquid phase reaction zone, the active metal is VIB group metal oxide and/or VIII group metal oxide, the content of the VIB group metal oxide is 5% -30% and the content of the VIII group metal oxide is 1% -15% based on the weight of the hydrogenation catalyst; the carrier is alumina or assistant modified alumina, the assistant is one or more of B, P, mg, zr and Si, and the weight of the carrier is taken as a reference, and the assistant accounts for 3-15% of the weight of the carrier.
16. The process according to claim 5, wherein: the temperature of the gas phase component subjected to heat exchange with the effluent of the liquid phase reaction zone is 200-340 ℃, then the gas phase component enters a desilication reactor, a silicon capturing agent is filled in the desilication reactor, the type of the silicon capturing agent is the same as that of the gas-liquid countercurrent zone, and the operation conditions of the desilication reactor are as follows: the pressure is 0.2-4.0 MPa, and the volume space velocity is 0.1-4.0 h -1 The reaction temperature is 150-400 ℃.
17. The process according to claim 5, wherein: when the Si content in the high-silicon naphtha oil is more than 500ppm, two desilication reactors are arranged for on-line switching operation.
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| Publication number | Priority date | Publication date | Assignee | Title |
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| CN1552819A (en) * | 2003-05-31 | 2004-12-08 | 中国石油化工股份有限公司 | Light hydrocarbon hydrogenation method |
| CN101343566A (en) * | 2007-07-09 | 2009-01-14 | 中国石油化工股份有限公司 | Method for improving running period of hydrogenation plant for poor petroleum naphtha |
| CN101591565A (en) * | 2008-05-29 | 2009-12-02 | 中国石油化工股份有限公司 | A kind of hydrofinishing process of inferior patrol |
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| Publication number | Priority date | Publication date | Assignee | Title |
|---|---|---|---|---|
| CN1552819A (en) * | 2003-05-31 | 2004-12-08 | 中国石油化工股份有限公司 | Light hydrocarbon hydrogenation method |
| CN101343566A (en) * | 2007-07-09 | 2009-01-14 | 中国石油化工股份有限公司 | Method for improving running period of hydrogenation plant for poor petroleum naphtha |
| CN101591565A (en) * | 2008-05-29 | 2009-12-02 | 中国石油化工股份有限公司 | A kind of hydrofinishing process of inferior patrol |
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