CN115725330A - Catalytic conversion method for producing low-carbon olefin - Google Patents
Catalytic conversion method for producing low-carbon olefin Download PDFInfo
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- 230000008569 process Effects 0.000 claims description 30
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Abstract
A catalytic conversion method for producing low-carbon olefins comprises the steps of contacting heavy hydrocarbon oil with a cracking catalyst, and sequentially carrying out catalytic cracking reaction in a first reactor and a third reactor; the isobutane-rich C4 hydrocarbon fraction, the isobutane-poor C4 hydrocarbon fraction and the optional light gasoline fraction are contacted with a catalyst, and catalytic cracking reaction is carried out in a second reactor and a third reactor in sequence; and separating oil gas and the catalyst at the outlet of the third reactor to obtain a spent catalyst and a reaction product, introducing the reaction product into a fractionating device for fractionation, and returning the spent catalyst to the first reactor and the second reactor for recycling after steam stripping and regeneration. The method provided by the invention has higher propylene yield.
Description
Technical Field
The invention relates to a catalytic conversion method for producing low-carbon olefin by heavy hydrocarbon oil.
Background
Ethylene and propylene are basic chemical raw materials, and at present, the ethylene and the propylene mainly come from steam thermal cracking of ethane, propane, butane, liquefied Petroleum Gas (LPG), condensate oil, naphtha, hydrocracking tail oil, crude diesel oil and the like and catalytic cracking of hydrocarbon oil. With the use of new light feedstocks for steam cracking, product distribution will change, e.g. ethane will be used as the steam cracking feedstock, and the ethylene proportion in the product will increase significantly and propylene will decrease compared to naphtha. Under the background, the catalytic conversion of heavy hydrocarbon oil to produce low-carbon olefin, especially propylene, is an effective supplementary measure for preparing ethylene by steam thermal cracking.
CN1004878B discloses a catalytic cracking method (DCC) for preparing lower olefins, in which petroleum fractions, residual oil or crude oil with different boiling ranges are used as raw materials, and a catalytic conversion reaction is carried out in a fluidized bed or moving bed reactor by using a solid acid catalyst at a temperature of 500-650 ℃ and a pressure of 1.5-3X 10 5 Pa, weight space velocity of 0.2-20h -1 The catalyst after reaction is returned to the reactor for recycling after being burnt and regenerated, and the catalyst has the catalyst oil ratio of 2-12. Compared with the conventional catalytic cracking and steam thermal cracking method, the method can obtain more propylene and butylene, and the total yield of the propylene and the butylene can reach about 40 percent. The yield of the C4 hydrocarbon fraction is about 20%, and if the part of the C4 hydrocarbon fraction is further converted, the yield of propylene and ethylene of DCC can be greatly improved.
When the heavy hydrocarbon oil is catalytically cracked to produce ethylene and propylene, a considerable part of C4 hydrocarbon fraction is obtained, wherein the part of C4 hydrocarbon fraction contains olefin and alkane, the olefin in the heavy hydrocarbon oil can be converted into propylene at present through recycling or single catalytic cracking, but the alkane is not utilized basically, and in the process of producing low-carbon olefin from the heavy hydrocarbon oil, the alkane proportion in the produced C4 hydrocarbon fraction is increased along with the increase of the proportion of the hydrogenated heavy hydrocarbon oil. How to efficiently convert C4 hydrocarbon fractions generated in the catalytic cracking process of heavy hydrocarbon oil is a problem which needs to be solved urgently in the catalytic cracking process of low-carbon olefins of the heavy hydrocarbon oil at present. In the prior art, the utilization efficiency of the C4 hydrocarbon fraction is not high, and a new process needs to be developed to improve the conversion efficiency of the whole C4 hydrocarbon fraction generated in the process of producing low-carbon olefins from heavy hydrocarbon oil, so that the propylene yield in the process of producing low-carbon olefins from heavy hydrocarbon oil is improved.
Disclosure of Invention
One of the purposes of the present invention is to provide a catalytic conversion method for producing low carbon olefins from heavy hydrocarbon oil, specifically, the present invention improves the propylene yield in the process of producing low carbon olefins from heavy hydrocarbon oil by improving the conversion efficiency of the whole C4 hydrocarbon fraction produced in the process of producing low carbon olefins from heavy hydrocarbon oil, and more specifically, the present invention improves the conversion rate of C4 alkanes by dividing the C4 hydrocarbon fraction into an isobutane-rich C4 hydrocarbon fraction and an isobutane-poor C4 hydrocarbon fraction to enrich alkanes in the isobutane-rich C4 fraction and thus improving the conversion severity of the isobutane-rich C4 hydrocarbon fraction; the olefins are enriched in the isobutane-depleted C4 hydrocarbon fraction and the olefins therein are converted under suitable conditions, thereby increasing the conversion efficiency of the entire C4 hydrocarbon fraction. The method has high yield of ethylene and propylene.
It is a further object of the present invention to provide a catalytic converter for use in the above method.
The invention provides a catalytic conversion method for producing low-carbon olefin, which comprises the following steps:
(1) Heavy hydrocarbon oil contacts with a cracking catalyst in a first reactor to carry out catalytic cracking reaction;
(2) The light hydrocarbon and the cracking catalyst are contacted in a second reactor to carry out catalytic cracking reaction; the light hydrocarbon is an isobutane-rich C4 hydrocarbon fraction, an isobutane-lean C4 hydrocarbon fraction and an optional light gasoline fraction;
(3) Introducing the catalyst and the oil-gas mixture reacted in the first reactor in the step (1) and the catalyst and the oil-gas mixture reacted in the second reactor in the step (2) into a third reactor for continuous reaction, then separating to obtain a spent catalyst and a reaction product, introducing the reaction product into a fractionating device for fractionation, and returning the spent catalyst to the first reactor and the second reactor for recycling after steam stripping and regeneration.
The reaction temperature of the first reactor is 550-580 deg.C, preferably 555-575 deg.C, the agent-oil ratio is 6-15, preferably 8-12, and the oil-gas retention time is 0.1-5 seconds, preferably 1-4 seconds.
In the second reactor, the reaction temperature of the isobutane-rich C4 hydrocarbon fraction is 660-690 ℃, the agent-oil ratio is 60-150, and the oil-gas retention time is 0.1-0.5 second; the reaction temperature of the isobutane-depleted C4 hydrocarbon fraction is 630-660 ℃, the agent-oil ratio is 20-40, and the oil gas retention time is 0.5-1.0 second; the reaction temperature of the light gasoline is 600-630 ℃, the agent-oil ratio is 12-30, and the oil-gas retention time is 0.5-1.5 seconds.
The reaction temperature of the third reactor is 550-590 ℃, preferably 555-585 ℃, and the weight hourly space velocity is 1h -1 -20h -1 Preferably for 2h -1 -15h -1 。
The heavy hydrocarbon oil is one or a mixture of several of vacuum gas oil, atmospheric residue oil, vacuum residue oil, hydrogenation vacuum gas oil, hydrogenation atmospheric residue oil, hydrocracking tail oil and deasphalted oil.
The isobutane-rich C4 hydrocarbon fraction and the isobutane-lean C4 hydrocarbon fraction are derived from the C4 hydrocarbon fraction of the fractionating device, and can also be derived from the C4 hydrocarbon fractions in other catalytic cracking devices, MTO devices, thermal cracking devices and the like. The C4 hydrocarbon fraction separation conditions are as follows: the theoretical plate number is 100-200, the reflux ratio is 0.10-0.30, the tower pressure is 0.4-0.8MPa, the tower top temperature is 60-70 ℃, and the tower kettle temperature is 75-85 ℃.
The isobutane content of the isobutane-rich C4 hydrocarbon fraction is not less than 50 wt%, and the olefin content of the isobutane-poor C4 hydrocarbon fraction is not less than 60 wt%;
the olefin content in the light gasoline fraction is 30-90 wt%, and the final distillation point of the light gasoline fraction is not more than 75 ℃, preferably 60-65 ℃. The light gasoline fraction comes from the catalytic conversion device of the invention and also from other catalytic cracking devices, MTO, coking devices and the like.
The weight ratio of the isobutane-rich C4 hydrocarbon fraction to the heavy hydrocarbon oil is 0.01-0.07; the weight ratio of the isobutane-depleted C4 hydrocarbon fraction to the heavy hydrocarbon oil is 0.05-0.18; the weight ratio of the light gasoline fraction to the heavy hydrocarbon oil is 0-0.15.
Preferably, one or more of the isobutane-rich C4 hydrocarbon fraction, the isobutane-poor C4 hydrocarbon fraction and the light gasoline fraction enter a second reactor after being in contact reaction with hydrogen on a nickel-containing catalyst at a reaction temperature of 20-90 ℃, a reaction pressure of 0.2-1 MPa and a molar ratio of hydrogen to olefin of 1-5.
The first reactor is preferably a riser reactor, the second reactor is preferably a riser reactor, and the third reactor is preferably a fluidized bed reactor.
Introducing into the first reactor a diluent in a weight ratio to said heavy hydrocarbon oil of from 0.01 to 0.5, preferably from 0.05 to 0.2; introducing into the second reactor a diluent in a weight ratio to said light hydrocarbons of from 0.01 to 0.5, preferably from 0.05 to 0.4; the diluent is selected from one or more of water vapor, low-carbon alkane and nitrogen.
The cracking catalyst containing the medium pore zeolite contains 1-60 wt% of zeolite, 5-99 wt% of refractory inorganic oxide and 0-70 wt% of clay based on the total weight of the catalyst; the zeolite contains 50-100 wt% of zeolite with MFI structure and 0-50 wt% of large pore zeolite, based on the total weight of the zeolite mixture.
The invention provides a hydrocarbon catalytic conversion device for producing low-carbon olefin, which comprises: the device comprises a riser reactor (1), a riser reactor (2), a fluidized bed reactor (3), a settler (4) and a stripper (5), wherein the stripper (5) is positioned below the fluidized bed reactor (3), and outlets of the riser reactor (1) and the fluidized bed reactor (2) are communicated with any position of the fluidized bed reactor (3).
The settler (4) is coaxial with the fluidized bed reactor (3) and is positioned right above the fluidized bed reactor; the stripper (5) is coaxial with the fluidized bed reactor (3) and is positioned right below the fluidized bed reactor; the riser reactor (1) extends through the interior of the stripper (5) into the bottom of the fluidized bed reactor, and the riser reactor (2) is located outside the stripper and extends through the interior of the settler (4) into the upper part of the fluidized bed reactor.
The catalytic conversion method provided by the invention achieves the purpose of increasing the yield of the low-carbon olefin by optimizing the catalytic process and the catalytic conditions and matching with corresponding catalysts, and has the advantages of wide raw oil application range, capability of simultaneously using various raw oils and high process flexibility. Compared with the existing hydrocarbon catalytic conversion, the hydrocarbon catalytic conversion method provided by the invention has higher yield of low-carbon olefin, especially propylene.
The catalytic cracking reaction device provided by the invention has high process flexibility, adopts a multi-zone coupling and zone regulation method, and is beneficial to improving the yield of propylene. And the isobutane-rich isobutane-C4 hydrocarbon fraction rich in isobutane is introduced to the bottom of the second reactor, so that on one hand, the conversion of isobutane is promoted by utilizing high-scale reaction conditions, and meanwhile, a small amount of carbon deposit of the catalyst can cover a part with strong acidity neutrality, on the other hand, an upward reaction material flow contains a large amount of positive carbon ions, and further, the reaction environment of the isobutane-rich isobutane-C4 hydrocarbon fraction poor in olefin is improved. Thereby greatly improving the conversion efficiency of the whole C4 hydrocarbon fraction.
In the context of the present invention, unless otherwise specified, the term "lower olefins" refers to C2-C3 olefins. C4 represents the number of carbon atoms in the hydrocarbon molecule of 4, and so on.
Drawings
FIG. 1 is a schematic view of a hydrocarbon catalytic conversion apparatus according to the present invention. Wherein:
1-heavy hydrocarbon oil cracking riser reactor (first reactor),
11-transfer line for transferring the regenerated catalyst to riser reactor 1,
12-regenerated catalyst flow control valve on the delivery line 11,
13-raw oil feed nozzle of riser reactor 1,
14-outlet distributor of riser reactor 1,
2-light hydrocarbon cracking riser reactor (second reactor),
21-riser reactor 2a transfer line for transferring regenerated catalyst,
22-a regenerated catalyst flow control valve on the delivery pipe 21,
23-isobutane-rich C4 hydrocarbon fraction feed nozzle of riser reactor 2,
24-isobutane-depleted C4 hydrocarbon fraction feed nozzle of riser reactor 2,
25-light gasoline feed nozzle of riser reactor 2,
3-fluidized bed reactor (third reactor),
4-a settler is arranged in the reaction tank,
41-a first-stage cyclone separator,
42-a two-stage cyclone separator, wherein,
43-Large oil-gas line at the outlet of the settler,
5-a stripper for the steam from the waste gas,
51-baffles in the stripper 5,
52-spent catalyst transfer line of stripper 5,
53-spent catalyst flow control valve on the delivery line 52.
6-a regenerator, wherein the gas is introduced into the reactor,
61-main wind into the regenerator line,
62-flue gas out of regenerator line.
Detailed Description
In the catalytic conversion method for producing low-carbon olefins from heavy hydrocarbon oil, the first reactor can be a riser reactor, a fluidized bed reactor, a down-flow conveyor line reactor or a composite reactor formed by a plurality of reactors connected in series and/or in parallel, and the reactors can be divided into two or more reaction zones according to requirements. The first reactor is preferably a riser reactor, and the riser reactor is one or more of an equal-diameter riser reactor, an equal-linear-speed riser reactor and a variable-diameter riser reactor. One or more first reactors may be included in the heavy hydrocarbon oil conversion process of the present invention. The reaction temperature of the first reactor is 550-580 ℃ (the outlet temperature of the reactor), preferably 555-575 ℃; the agent-to-oil ratio (weight ratio of catalyst introduced into the first reactor to heavy hydrocarbon oil) is from 6 to 15, preferably from 8 to 12; the oil gas residence time is 0.1-5 seconds, preferably 1-4 seconds, and the absolute pressure of the reaction zone is 0.15-0.30 MPa (outlet pressure), preferably 0.17-0.24 MPa. And (2) injecting a diluent into the first reactor in the step (1) to reduce the partial pressure of the hydrocarbon oil raw material, wherein the diluent is selected from one or more of steam, C1-C4 low-carbon alkane and nitrogen, and the weight ratio of the diluent to the heavy hydrocarbon oil is 0.01-0.5, preferably 0.05-0.20. The diluent is preferably water vapour.
In the catalytic conversion method for producing low-carbon olefins from heavy hydrocarbon oil, the second reactor can be a riser reactor, a fluidized bed reactor, a down-flow conveyor line reactor or a composite reactor formed by a plurality of reactors connected in series and/or in parallel, wherein each reactor can be divided into two or more reaction zones according to requirements. The preferable second reactor is a riser reactor, which can be one riser reactor, or a plurality of riser reactors connected in parallel or in series, wherein the riser reactor is one or more of an equal-diameter riser reactor, an equal-linear-speed riser reactor and a variable-diameter riser reactor. One or more secondary reactors may be included in the hydrocarbon conversion process of the present invention. The light hydrocarbon and the cracking catalyst are contacted in a second reactor to carry out catalytic cracking reaction; the light hydrocarbon is isobutane-rich C4 hydrocarbon fraction, isobutane-poor C4 hydrocarbon fraction and optional light gasoline fraction, the reaction temperature of the isobutane-rich C4 hydrocarbon fraction is 660-690 ℃, the catalyst-oil ratio is 60-150, and the oil-gas retention time is 0.1-0.5 second; the reaction temperature of the isobutane-depleted C4 hydrocarbon fraction is 630-660 ℃, the agent-oil ratio is 20-40, and the oil gas retention time is 0.5-1.0 second; the reaction temperature of the light gasoline is 600-630 ℃, the agent-oil ratio is 12-30, and the oil-gas retention time is 0.5-1.5 seconds; the absolute pressure of the reaction zone of the second reactor is between 0.15MPa and 0.30MPa (outlet pressure). And (3) injecting a diluent into the second reactor in the step (2) to reduce the partial pressure of the hydrocarbon oil raw material, wherein the diluent is selected from one or more of water vapor, C1-C4 low-carbon alkane and nitrogen, and the weight ratio of the diluent to the light hydrocarbon is 0.01-0.5, preferably 0.05-0.40. The diluent is preferably water vapour.
In the catalytic conversion method for producing low-carbon olefins from heavy hydrocarbon oil, the third reactor is preferably a fluidized bed reactor, and can be one or a plurality of fluidized bed reactors connected in parallel or in series; the fluidized bed reactor is selected from one or more of a bulk fluidized bed reactor, a bubbling bed reactor, a turbulent bed reactor, a fast bed reactor and a dense phase fluidized bed reactor. The reaction temperature of the third reactor is 550-590 ℃ (bed temperature), preferably 555-585 ℃; weight hourly space velocity of 1h -1 -20h -1 Preferably 2h -1 -15h -1 (ii) a The absolute pressure of the reaction zone is 0.15MPa-0.30MPa.
In the hydrocarbon catalytic conversion method for producing propylene and light aromatic hydrocarbon by using heavy hydrocarbon oil, the heavy hydrocarbon oil in the step (1) is one or a mixture of several of vacuum gas oil, atmospheric residue oil, vacuum residue oil, hydrogenation vacuum gas oil, hydrogenation atmospheric residue oil, hydrocracking tail oil and deasphalted oil.
In the catalytic conversion method for producing low-carbon olefins from heavy hydrocarbon oil, the light hydrocarbons in the step (2) are isobutane-rich C4 hydrocarbon fraction, isobutane-poor C4 hydrocarbon and optional light gasoline fraction. Wherein the isobutane-rich and isobutane-lean C4 hydrocarbon fractions are derived from the C4 hydrocarbon fractions of the fractionation unit according to the present invention. The C4 hydrocarbon fraction separation conditions are as follows: the theoretical plate number is 100-200, the reflux ratio is 0.10-0.30, the tower pressure is 0.4-0.8MPa, the tower top temperature is 60-70 ℃, and the tower kettle temperature is 75-85 ℃. The isobutane content of the isobutane-rich C4 hydrocarbon fraction is not less than 50 wt%, and the olefin content of the isobutane-poor C4 hydrocarbon fraction is not less than 60 wt%; wherein the final distillation point of the medium and light gasoline fraction is not more than 75 ℃, and preferably 60-65 ℃. The light gasoline fraction has an olefin content of from 30 to 90% by weight, preferably from 40 to 90% by weight, more preferably from 50 to 90% by weight; the light gasoline fraction and the C4 hydrocarbon fraction can be light gasoline fraction and C4 hydrocarbon fraction obtained by the fractionating device of the method of the invention, and can also be from other cracking devices, for example, the light gasoline fraction can be catalytic cracking light gasoline, steam cracking light gasoline, MTO light gasoline and coking light gasoline, and the C4 hydrocarbon fraction can be C4 hydrocarbon fraction generated by catalytic cracking, steam cracking, coking and MTO. Preferably, the light hydrocarbons comprise a light gasoline fraction and a C4 hydrocarbon fraction obtained by fractionation in the fractionation system of the present process.
In the catalytic conversion method for producing low-carbon olefins from heavy hydrocarbon oil, provided by the invention, the obtained oil-gas mixture is preferably introduced into a fractionation device, and is fractionated to obtain low-carbon olefins, C4 hydrocarbon fractions, light gasoline, heavy gasoline, light oil, heavy oil and other low-molecular saturated hydrocarbons, wherein the obtained C4 hydrocarbon fractions and optional light gasoline fractions can be partially or completely returned to the bottom of a second reactor for catalytic cracking reaction. The light hydrocarbons preferably do not contain dienes or alkynes; when the cracked light gasoline and the C4 hydrocarbon fraction are used as light hydrocarbon raw materials, the light gasoline is preferably subjected to selective hydrogenation to remove dienes and alkynes therein and then introduced into a second reactor for reaction, and the C4 hydrocarbon fraction is preferably subjected to selective hydrogenation to remove dienes and alkynes therein and then introduced into different positions of the second reactor for reaction after separation. The method comprises the following steps of carrying out contact reaction on C4 hydrocarbons, catalytic cracked gasoline whole fraction or cut gasoline fraction such as light gasoline fraction with the distillation range of 30-75 ℃ and C4 hydrocarbon fraction with hydrogen on a nickel-containing catalyst at the reaction temperature of 20-90 ℃, the reaction pressure of 0.2-1 MPa and the molar ratio of hydrogen to olefin of 1-5.
In the hydrocarbon catalytic conversion method for producing the low-carbon olefin, the cracking catalyst in the step (1), the step (2) and the step (3) is the same cracking catalyst, and the cracking catalyst contains 1-60 wt% of zeolite mixture, 5-99 wt% of heat-resistant inorganic oxide and 0-70 wt% of clay based on the total weight of the cracking catalyst, wherein the zeolite mixture contains 50-100 wt% of zeolite with an MFI structure and 0-50 wt% of large-pore zeolite based on the total weight of the zeolite mixture. Preferably, the cracking catalyst contains 10-50% of the zeolite mixture, 10-70% of the refractory inorganic oxide, and 0-60% of the clay, based on the total weight of the cracking catalyst.
The zeolite with MFI structure is high-silicon zeolite with pentasil structure selected from one or more of ZSM-5 and ZRP series zeolite, and the preferred zeolite with MFI structure is one or more of rare earth-containing ZRP zeolite (CN 1052290A, CN1058382A, US 5232675), phosphorus-containing ZRP zeolite (CN 1194181A, US 5951963), phosphorus-and rare earth-containing ZRP zeolite (CN 1147420A), phosphorus-and alkaline earth metal-containing ZRP zeolite (CN 1211469A, CN1211470A, US 6080698) and phosphorus-and transition metal-containing ZRP zeolite (CN 1465527A, CN 1611299A). The large-pore zeolite is selected from one or more of Y-type zeolite, HY-type zeolite and USY-type zeolite. Zeolites and macros of the MFI structureThe porous zeolite may be commercially available or prepared by various methods known in the art, and will not be described herein. The heat-resistant inorganic oxide is selected from SiO 2 And/or Al 2 O 3 (ii) a The clay is selected from kaolin and/or halloysite.
The catalytic cracking device provided by the invention comprises a first riser reactor, a second riser reactor, a stripper and a settler, wherein the stripper is positioned below the third reactor, the outlet of the first riser reactor and the outlet of the second riser reactor are communicated with any position of the fluidized bed reactor, and preferably, the stripper is coaxial with the third reactor and positioned below the third reactor. Wherein, the first and the second riser reactors are selected from at least any one of circular tubes with equal diameter, frustum-shaped cylinders or a combination formed by connecting 1-6 sections of straight cylinders with different diameters through reducer sections; the fluidized bed reactor is at least any one of a cylinder body with equal diameter, a frustum-shaped cylinder body or a combination body formed by connecting 1-6 sections of straight cylinder bodies with different diameters through a reducing section.
The catalytic cracking device can also comprise a catalyst regenerator, and the carbon deposited catalyst after steam stripping is conveyed to the catalyst regenerator through a catalyst conveying passage; the regenerated catalyst is supplied to the catalytic cracking reactor via at least one catalyst transfer path, which is well known to those skilled in the art and will not be described in further detail herein.
The method provided by the present invention will be further described with reference to the accompanying drawings, but the present invention is not limited thereby.
One embodiment of the catalytic cracking unit of the present invention is shown in fig. 1, and although the schematic diagram is a simplified flow scheme, it does not affect the understanding of the present invention by those skilled in the art. The device comprises a riser reactor 1 (a first reactor), a riser reactor 2 (a second reactor), a fluidized bed reactor 3 (a third reactor), a settler 4 and a stripper 5, wherein the settler 4, the stripper 5, the riser reactor 1 and the fluidized bed reactor 3 are coaxial, the stripper 5 is positioned below the fluidized bed reactor 3, an outlet of the riser reactor 1 is communicated with any position of the bottom of the fluidized bed reactor 3, and an outlet of the riser reactor 2 is communicated with any position of the upper part of the fluidized bed reactor 3.
Heavy hydrocarbon oil raw material is preheated to 180-340 ℃, then is sprayed into the riser reactor 1 together with steam through a nozzle 13, and contacts and reacts with the thermal regeneration catalyst from the pipeline 11 under the conditions that the temperature is 550-580 ℃, preferably 555-575 ℃, the pressure is 0.15-0.30 MPa, preferably 0.18-0.28 MPa (absolute pressure), the weight ratio of the catalyst to the hydrocarbon raw material is 6-15, the oil gas residence time is 0.1-5 seconds, preferably 1-4 seconds. The reactant stream and catalyst are introduced into the fluidized bed reactor 3 to continue the reaction.
The isobutane-rich C4 hydrocarbon fraction from the separating device of the invention is introduced from the bottom of the riser reactor 2, and contacts with the hot regenerated catalyst to react at the lower part of the riser reactor 2 under the conditions of 660-690 ℃, 60-150 of agent-oil ratio and 0.1-0.5 second of oil gas retention time. The isobutane-poor C4 hydrocarbon fraction from the separation device is introduced into the back part of the isobutane-rich C4 hydrocarbon fraction reaction section of the riser reactor 2, and is in contact reaction with the catalyst in the middle part of the riser reactor 2 under the conditions of 630-660 ℃, 20-40 of agent-oil ratio and 0.5-1 second of oil gas retention time. The light gasoline fraction from the separating device is introduced into the riser reactor 2 at the back of the reaction section of the isobutane-poor C4 hydrocarbon fraction, and is in contact reaction with the catalyst at the upper part of the riser reactor 2 under the conditions of 600-630 ℃, 12-30 of agent-oil ratio and 0.5-1.5 seconds of oil-gas retention time. The absolute pressure of the riser reactor 2 is 0.15MPa-0.30MPa. The reaction oil gas and the catalyst in the riser reactor 2 directly enter the upper part of the fluidized bed reactor 3 and continue to react in the fluidized bed reactor 3. In order to improve the yield of the low-carbon olefins, particularly the propylene, and the yield of the light aromatic hydrocarbons, the invention preferably comprises the steps of subjecting the C4 hydrocarbon fraction rich in olefins and the light gasoline component (the final distillation point is not more than 65 ℃, preferably 60-65 ℃) obtained by the fractionating device to selective hydrogenation reaction on a nickel-containing catalyst at a reaction temperature of 20-90 ℃ and a reaction pressure of 0.2-1 MPa, wherein the molar ratio of hydrogen to olefins is 1-5. Isobutane-rich C4 hydrocarbon fraction is injected through nozzle 23 into the columnThe bottom of riser reactor 2 is contacted and reacted with hot regenerated catalyst from line 21. The isobutane-depleted C4 hydrocarbon fraction is injected into the lower middle portion of riser reactor 2 through nozzle 24 and reacts in contact with the hot catalyst from the isobutane-enriched C4 hydrocarbon fraction reaction zone. The light gasoline is sprayed into the middle upper part of the riser reactor 2 through a nozzle 25 and contacts and reacts with the hot catalyst from the isobutane-lean C4 hydrocarbon fraction reaction section. The nozzle 23 is located above the pre-lift section of the second reactor of the riser. The reaction material flow and the catalyst from the riser reactor 2 enter a fluidized bed reactor 3 without separation, the temperature is 550-590 ℃, preferably 555-585 ℃, the reaction pressure is 0.15-0.30 MPa, preferably 0.18-0.28 MPa (absolute pressure), and the weight hourly space velocity is 1h -1 -20h -1 Preferably for 2h -1 -15h -1 Under the conditions of (1). After the reaction in the fluidized bed reactor 3, the oil gas and the catalyst carried by the oil gas enter a settler, most of the catalyst returns to the fluidized bed reactor through sedimentation, the oil gas and a small amount of the catalyst carried by the oil gas enter a primary cyclone separator 41 and a secondary cyclone separator 42, the catalyst is separated from the oil gas, the separated catalyst returns to the fluidized bed 3, and the oil gas enters a fractionating device (not shown in the figure) through a large oil gas pipeline 43. The spent catalyst enters a stripper 5 from the bottom of a fluidized bed 3 to strip oil gas carried by the spent catalyst, enters a regenerator 6 through a spent agent conveying pipeline 52 to be regenerated, and the regenerated hot catalyst is respectively conveyed to the riser reactor 1 and the riser reactor 2 through regenerant conveying pipelines 11 and 12 to be reused. Gas (including carbon dioxide, carbon monoxide, dry gas and liquefied gas), light gasoline, heavy gasoline, cracked light oil and heavy oil are obtained in a fractionating device. The gas product is separated by separation technology well known to those skilled in the art to obtain low-carbon olefins such as propylene and a C4 hydrocarbon fraction, and the C4 hydrocarbon fraction is further separated to obtain an isobutane-rich C4 hydrocarbon fraction and an isobutane-depleted C4 hydrocarbon fraction.
The following examples further illustrate the process but are not intended to limit it. The test was carried out on a medium-sized test apparatus, the flow of which is shown in the drawing, wherein the inner diameter (diameter) of the riser reactor 1 was 18 mm and the height was 5m, the inner diameter (diameter) of the riser reactor 2 was 12 mm and the height was 6 m, the inner diameter (diameter) of the fluidized bed reactor 3 was 64 mm and the height was 0.4 m, and the inner diameter of the settler was 300 mm.
Preparation of the catalysts used in the examples: in 175 kg of alumina sol (product of Qilu division, petrochemical catalyst Co., ltd., china, al) 2 O 3 11.4%, pH: 2-3), adding 62.5 kg of kaolin (industrial product of Suzhou china clay company, solid content is 80%), 33.3 kg of USY slurry (trade mark DASY, produced by Chinese petrochemical catalysis Co., ltd., solid content is 30%), stirring for 90 minutes, adding 66.7 kg of ZSM-5 zeolite slurry (trade mark ZRP, produced by Chinese petrochemical catalysis Co., ltd., solid content is 30%) and 162.5 kg of decationized water, homogenizing, spray drying, washing the obtained sample until the pH value is close to 6, drying, and roasting at 500 ℃ for 3 hours to obtain a catalyst sample.
The catalyst used contained 10 wt% USY zeolite, 20 wt% ZSM-5 zeolite, 50 wt% kaolin and 20 wt% alumina binder, based on the total weight of the catalyst. The catalyst was aged at 800 ℃ for 10 hours in an atmosphere of 100% steam, and the amount of the catalyst contained in the apparatus (system catalyst inventory) was about 60 kg.
Example 1
Heavy hydrocarbon oil (the properties of which are shown in table 1) is introduced into the riser reactor 1, and after the heavy hydrocarbon oil is contacted with the hot catalyst from the regenerator 6 for reaction, the reaction oil gas and the catalyst do not need to be separated, and the heavy hydrocarbon oil is introduced into the fluidized bed reactor 3 for continuous reaction. Isobutane-rich C4 hydrocarbon fraction, isobutane-poor C4 hydrocarbon fraction and light gasoline fraction (the compositions of isobutane-rich C4 hydrocarbon fraction and isobutane-poor C4 hydrocarbon fraction are shown in table 2, the units are mass percent, and the properties of light gasoline are shown in table 5) are respectively introduced from the bottom, middle lower part and middle upper part of the riser reactor 2, and contact with a hot catalyst to react, and the reaction oil gas and the catalyst are introduced into the fluidized bed reactor 3 to continue to react. The oil gas and the catalyst pass through an outlet of the fluidized bed reactor 3, most of the catalyst in the settler returns to the fluidized bed reactor 3 through sedimentation, the oil gas carrying a small amount of catalyst enters a cyclone separator, and the separated catalyst returns to the fluidized bed. The spent catalyst in the fluidized bed reactor 3 enters a stripper 5, and is conveyed to a regenerator 6 for regeneration after oil gas carried by the catalyst is stripped. And the oil gas obtained by separation enters a fractionating device. The weight ratios of the isobutane-rich C4 hydrocarbon fraction, isobutane-lean C4 hydrocarbon fraction and light gasoline introduced into the riser reactor 2 to the heavy hydrocarbon oil introduced into the riser reactor 1 were 0.026, 0.074, and 0.08, respectively, and the reaction conditions and the reaction results are shown in table 6. The theoretical plate number of the C4 hydrocarbon fraction separation is 160, the reflux ratio is 0.20, the tower pressure is 0.7MPa, the tower top temperature is 68 ℃, and the tower kettle temperature is 80 ℃.
Example 2
The procedure of this example is the same as example 1 except that light gasoline (properties are shown in Table 5) is subjected to selective hydrogenation reaction on a catalyst (product name: RDD-1, produced by Zhongpetrochemical ChangLing catalyst division) at a reaction temperature of 40 ℃ and a reaction pressure of 0.5MPa and a molar ratio of hydrogen to olefin of 4, and diolefin and alkyne are converted into olefin and introduced into the riser reactor 2. C4 hydrocarbon fractions (composition shown in table 2) are subjected to selective hydrogenation reaction on a catalyst (product brand is RDD-1, produced by Zhongpetrochemical ChangLing catalyst division) at a reaction temperature of 30 ℃ and a reaction pressure of 0.45MPa, with a molar ratio of hydrogen to olefins of 4, dienes and alkynes are converted into olefins, and then the olefins are separated to obtain isobutane-rich C4 hydrocarbon fractions and isobutane-poor C4 hydrocarbon fractions, and then the fractions are introduced into the riser reactor 2 from different positions. The isobutane-rich C4 hydrocarbon fraction, isobutane-lean C4 hydrocarbon fraction (composition of C4 hydrocarbon fraction, isobutane-rich C4 hydrocarbon fraction, isobutane-lean C4 hydrocarbon fraction are shown in table 3, unit is mass percent) and the weight ratio of light gasoline (properties of light gasoline are shown in table 5) to heavy hydrocarbon oil introduced into the riser 1 are 0.026, 0.074, 0.08, respectively, and the reaction conditions and results are shown in table 6.
Example 3
The procedure of this example is the same as example 2 except that the weight ratios of the isobutane-rich C4 hydrocarbon fraction, the isobutane-lean C4 hydrocarbon fraction and the light gasoline introduced into the riser reactor 2 to the heavy hydrocarbon oil introduced into the riser reactor 1 were 0.052, 0.148 and 0.08, respectively, and the reaction conditions and the reaction results are shown in table 6.
Example 4
The flow of this example is the same as that of example 3, except that the isobutane concentration in the isobutane-rich C4 hydrocarbon fraction introduced into the riser 2 is higher (the composition of the C4 hydrocarbon fraction, the isobutane-rich C4 hydrocarbon fraction and the isobutane-depleted C4 hydrocarbon fraction is shown in table 4, the units are mass percentages), and the reaction conditions and the reaction results are shown in table 6. The theoretical plate number of the C4 hydrocarbon fraction separation is 190, the reflux ratio is 0.25, the tower pressure is 0.7MPa, the tower top temperature is 68 ℃, and the tower kettle temperature is 80 ℃.
Comparative example 1
Heavy hydrocarbon oil (the properties of which are shown in table 1) is introduced into the riser reactor 1, and after the heavy hydrocarbon oil is contacted with the hot catalyst from the regenerator 6 for reaction, the reaction oil gas and the catalyst do not need to be separated, and the heavy hydrocarbon oil is introduced into the fluidized bed reactor 3 for continuous reaction. C4 hydrocarbon fraction and light gasoline fraction (the composition of the C4 hydrocarbon fraction is shown in table 2, the properties of the light gasoline are shown in table 5) are respectively introduced from the bottom and the middle upper part of the riser reactor 2 and are in contact reaction with a hot catalyst, and reaction oil gas and the catalyst are introduced into the fluidized bed reactor 3 for continuous reaction. The oil gas and the catalyst pass through an outlet of the fluidized bed reactor 3, most of the catalyst in the settler returns to the fluidized bed reactor 3 through sedimentation, the oil gas carrying a small amount of catalyst enters a cyclone separator, and the separated catalyst returns to the fluidized bed. The carbon-deposited catalyst in the fluidized bed enters a stripper 5, and after oil gas carried by the catalyst is stripped, the catalyst is conveyed to a regenerator for regeneration. The oil gas obtained by separation enters a fractionating device. The weight ratios of the C4 hydrocarbon fraction and the light gasoline introduced into riser reactor 2 to the heavy hydrocarbon oil introduced into riser 1 were 0.10 and 0.08, respectively, and the reaction conditions and the reaction results are shown in table 7.
Comparative example 2
The flow of this comparative example was the same as that of comparative example 1 except that no second reactor was provided. The heavy hydrocarbon oil (the properties of which are shown in table 1) is introduced into the riser reactor 1, and after the heavy hydrocarbon oil is contacted and reacted with the hot catalyst from the regenerator, the reaction oil gas and the catalyst do not need to be separated, and the heavy hydrocarbon oil is introduced into a third reactor for continuous reaction. The oil gas and the catalyst pass through an outlet of the fluidized bed reactor 3, most of the catalyst returns to the fluidized bed through sedimentation in the settler, the oil gas carrying a small amount of catalyst enters the cyclone separator, and the separated catalyst returns to the fluidized bed. The catalyst to be generated in the fluidized bed enters a stripper 5, and is conveyed to a regenerator for regeneration after oil gas carried by the catalyst is stripped. And the oil gas obtained by separation enters a fractionating device. The reaction conditions and the reaction results are shown in Table 7.
TABLE 1
TABLE 2
TABLE 3
TABLE 4
TABLE 5
| Example 1 | Examples 2, 3 and 4 | |
| Before hydrogenation | After hydrogenation | |
| The composition of light gasoline family, weight% | ||
| Alkane(s) | 30 | 30 |
| Olefins | 64 | 66 |
| |
3 | 3 |
| Aromatic hydrocarbons | 1 | 1 |
| Diene and alkyne content | 2 | 0 |
TABLE 6
TABLE 7
| Comparative example 1 | Comparative example 2 | |
| Reaction conditions of heavy hydrocarbon oil in riser | ||
| Reaction temperature of | 560 | 580 |
| Ratio of agent to oil | 8.0 | 8.5 |
| Residence time in seconds | 2.5 | 2.3 |
| Water injection (in weight percent based on the heavy hydrocarbon oil) | 20 | 25 |
| C4 Hydrocarbon fraction reaction conditions | ||
| Reaction temperature of | 655 | / |
| Ratio of solvent to oil (to C4 hydrocarbon fraction) | 32 | / |
| Residence time in seconds | 0.65 | / |
| Reaction conditions of light gasoline | ||
| Reaction temperature of | 610 | / |
| Solvent to oil ratio (for C4 hydrocarbon fraction and light gasoline) | 17.8 | / |
| Residence time in seconds | 0.82 | / |
| Injecting water (accounting for heavy hydrocarbon oil) into the second reactor in percentage by weight | 5 | / |
| Fluidized bed reaction conditions | ||
| Reaction temperature of | 570 | 565 |
| Space velocity, h -1 | 4.0 | 4.2 |
| Reaction pressure (absolute pressure), MPa | 0.18 | 0.18 |
| Weight ratio of C4 hydrocarbon fraction to heavy hydrocarbon oil | 0.1 | / |
| Weight ratio of light gasoline to heavy hydrocarbon oil | 0.08 | / |
| Distribution of the product, weight% | ||
| H 2 S+H 2 ～C 2 | 11.56 | 10.64 |
| Liquefied gas | 37.63 | 40.64 |
| Pyrolysis gasoline | 28.59 | 26.90 |
| Cracking light oil | 10.1 | 9.91 |
| Heavy oil | 3.42 | 3.50 |
| Coke | 8.7 | 8.41 |
| Total of | 100 | 100 |
| Yield of propylene,% by weight | 20.49 | 19.50 |
* The catalyst to oil ratio in tables 6 and 7 refers to the weight ratio of catalyst to reactor feed. The material balance is calculated based on the heavy hydrocarbon oil. The propylene yield is obtained by dividing the weight of propylene in the product by the weight of the heavy hydrocarbon oil.
* The reaction pressure in tables 6 and 7 refers to the settler outlet pressure.
As can be seen from tables 6 and 7, the process of the present invention has a higher propylene yield.
Claims (20)
1. A catalytic conversion process for producing lower olefins comprising:
(1) The heavy hydrocarbon oil contacts with a cracking catalyst in a first reactor to carry out catalytic cracking reaction;
(2) The light hydrocarbon and the cracking catalyst are contacted in a second reactor to carry out catalytic cracking reaction; the light hydrocarbon is an isobutane-rich C4 hydrocarbon fraction, an isobutane-lean C4 hydrocarbon fraction and an optional light gasoline fraction;
(3) Introducing the catalyst and the oil-gas mixture reacted in the first reactor in the step (1) and the catalyst and the oil-gas mixture reacted in the second reactor in the step (2) into a third reactor for continuous reaction, then separating to obtain a spent catalyst and a reaction product, introducing the reaction product into a fractionating device for fractionation, and returning the spent catalyst to the first reactor and the second reactor for recycling after stripping and regeneration.
2. The process of claim 1 wherein the first reactor has a reaction temperature of 550 ℃ to 580 ℃, a catalyst to oil ratio of 6 to 15, and an oil gas residence time of 0.1 seconds to 5 seconds.
3. The process of claim 2 wherein the first reactor has a reaction temperature of 555 to 575 ℃, a catalyst to oil ratio of 8 to 12, and a gas oil residence time of 1 to 4 seconds.
4. The method according to claim 1, wherein in the second reactor, the reaction temperature of the isobutane-rich C4 hydrocarbon fraction is 660 ℃ to 690 ℃, the agent-oil ratio is 60 to 150, and the oil-gas residence time is 0.1 second to 0.5 second; the reaction temperature of the isobutane-depleted C4 hydrocarbon fraction is 630-660 ℃, the agent-oil ratio is 20-40, and the oil-gas retention time is 0.5-1.0 second; the reaction temperature of the light gasoline is 600-630 ℃, the agent-oil ratio is 12-30, and the oil gas retention time is 0.5-1.5 seconds.
5. The process of claim 1, wherein the third reactor has a reaction temperature of 550 ℃ to 590 ℃ and a weight hourly space velocity of 1h -1 -20h -1 。
6. The method according to claim 5, wherein the reaction temperature of the third reactor is 555-585 ℃ and the weight hourly space velocity is 2h -1 -15h -1 。
7. The process of claim 1, wherein the heavy hydrocarbon oil is selected from one or more of vacuum gas oil, atmospheric residue, vacuum residue, hydrogenated vacuum gas oil, hydrogenated atmospheric residue, hydrocracked tail oil, deasphalted oil.
8. The process according to claim 1, characterized in that the weight ratio of isobutane C4 hydrocarbon-rich fraction to heavy hydrocarbon oil is from 0.01 to 0.07.
9. The process according to claim 1, characterized in that the weight ratio of isobutane-depleted C4 hydrocarbon fraction to heavy hydrocarbon oil is from 0.05 to 0.18.
10. The process according to claim 1, characterized in that the weight ratio of said light gasoline fraction to the heavy hydrocarbon oil is from 0 to 0.15.
11. The method according to claim 1, wherein said isobutane-rich C4 hydrocarbon fraction and isobutane-lean C4 hydrocarbon fraction are derived from the C4 hydrocarbon fraction of the fractionation plant according to the present invention.
12. The process according to claim 11, characterized in that said C4 hydrocarbon fraction separation conditions are: the theoretical plate number is 100-200, the reflux ratio is 0.10-0.30, the tower pressure is 0.4-0.8MPa, the tower top temperature is 60-70 ℃, and the tower kettle temperature is 75-85 ℃.
13. The method according to claim 1, wherein one or more of the isobutane-rich C4 hydrocarbon fraction, the isobutane-lean C4 hydrocarbon fraction and the light gasoline fraction enters the second reactor after being in contact with hydrogen for reaction at a reaction temperature of 20 ℃ to 90 ℃, a reaction pressure of 0.2MPa to 1MPa and a molar ratio of hydrogen to olefin of 1 to 5.
14. The process according to claims 1 and 13, wherein the isobutane content in said isobutane-rich C4 hydrocarbon fraction is not less than 50 wt.%, and said isobutane-depleted C4 hydrocarbon fraction has an olefin content of not less than 60 wt.%.
15. The process according to claims 1 and 4, wherein the light gasoline fraction has an olefin content of 30 to 90% by weight and a final boiling point of not more than 75 ℃, preferably 60 ℃ to 65 ℃.
16. The process of claim 1 wherein said first reactor is a riser reactor, said second reactor is a riser reactor, and said third reactor is a fluidized bed reactor.
17. The process according to claim 1, characterized in that a diluent is introduced into the first reactor, in a weight ratio of diluent to said heavy hydrocarbon oil ranging from 0.01 to 0.5, preferably from 0.05 to 0.2; introducing a diluent into the second reactor, wherein the weight ratio of the diluent to the light hydrocarbon is 0.01-0.5, preferably 0.05-0.4; the diluent is selected from one or more of water vapor, low-carbon alkane and nitrogen.
18. The process of claim 1 wherein the cracking catalyst comprises 1 to 60 wt% zeolite, 5 to 99 wt% refractory inorganic oxide, and 0 to 70 wt% clay, based on the total weight of the catalyst; the zeolite contains 50-100 wt% of zeolite with MFI structure and 0-50 wt% of large pore zeolite, based on the total weight of the zeolite mixture.
19. A catalytic hydrocarbon conversion apparatus for producing lower olefins according to claim 1, comprising a riser reactor (1), a riser reactor (2), a fluidized bed reactor (3), a settler (4) and a stripper (5), wherein the stripper (5) is located below the fluidized bed reactor (3), and the outlets of the riser reactor (1) and the riser reactor (2) are communicated with any position of the fluidized bed reactor (3).
20. Catalytic converter unit according to claim 19, characterized in that said settler (4) is coaxial with and directly above said fluidized bed reactor (3); the stripper (5) is coaxial with the fluidized bed reactor (3) and is positioned right below the fluidized bed reactor; the riser reactor (1) extends through the interior of the stripper (5) into the bottom of the fluidized bed reactor, and the riser reactor (2) is located outside the stripper and extends through the interior of the settler (4) into the upper part of the fluidized bed reactor.
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