A device and method for producing light aromatics by catalytic conversion of petroleum hydrocarbons

Abstract

The present application relates to a device and method for catalytic conversion of petroleum hydrocarbons to produce light aromatics, the device comprising: a reactor system, wherein at least one of the main reactor and the auxiliary reactor in the reactor system is a variable diameter catalytic conversion reactor, the variable diameter catalytic conversion reactor comprising: a reaction zone, a pre-lifting section, a first transition section, an outlet zone, and a second transition section, the inner cavity wall of the variable diameter reaction section is an arc-shaped wall; the cross-sectional area along the radial direction of the variable diameter reaction section gradually increases from the two ends of the variable diameter reaction section to the middle of the variable diameter reaction section. The method and device provided by the present invention can improve the conversion rate of the catalytic conversion reaction, and improve the yield and selectivity of hydrocarbon oil conversion to produce light aromatics.

Description

Technical Field

[0001] The present invention relates to the field of petroleum catalytic cracking, and in particular to a device and method for catalytically converting petroleum hydrocarbons to produce light aromatic hydrocarbons in maximum quantity. Background Art

[0002] With the trend toward heavier crude oil and the rapid growth in market demand for lighter oil products, catalytic cracking technology, a method for converting heavy oil to lighter products, has rapidly developed in my country. Traditional catalytic cracking technology primarily produces gasoline and diesel. However, with the expansion of refineries and capacity, as well as the widespread adoption of electric vehicles, the refined oil industry, primarily focused on gasoline and diesel production, is experiencing severe overcapacity. According to statistics, China's refining industry had an excess capacity of 100 million tons in 2016. Therefore, "oil conversion" has become an inevitable trend.

[0003] Light aromatics (benzene, toluene, and xylene, or BTX), represented by paraxylene, are important raw materials for the production of chemical products, and demand for them is increasing both domestically and internationally. Catalytic cracking gasoline and diesel contain a certain amount of aromatics. Conventional catalytic cracking of heavy oil, such as FCC and MIP, typically produces gasoline with an aromatic content of 20% to 35%, while high-temperature catalytic cracking (DCC) technology typically produces gasoline with an aromatic content of 20% to 50%. Diesel oil also has a higher aromatic content, with single-ring aromatics ranging from 20% to 40%. However, current technology makes it difficult to directly produce large quantities of light aromatics and other chemical feedstocks from inferior heavy oil.

[0004] Patents CN200710120106.5, CN200810101852.4, CN200810227651.9, CN200810227652.3, and CN200410006189.1 all disclose catalytic conversion methods for catalytic cracking heavy feedstocks to increase propylene production and simultaneously produce aromatics. By recycling difficult-to-crack feedstocks, re-cracked feedstocks, or recycled cracked feedstocks (180-250°C cycle oil fractions), propylene yields are increased while producing an aromatics-rich fraction. Patent CN201510493527.7 discloses a heavy oil catalytic cracking process for increasing the production of light and heavy aromatics. By enriching light and heavy aromatics through fractionation and gasoline recycling, the aromatics content in the light and heavy aromatic products can reach over 85%. The above patented technology takes into account the recycling of by-product aromatic hydrocarbon fractions (cycle oil) to produce aromatic hydrocarbons, but it mainly realizes the saturated cracking and decomposition reaction in the cycle oil fraction to produce light olefins, and has no obvious effect on increasing the yield of light aromatic hydrocarbons.

[0005] Patent CN201310517666.X discloses a catalytic conversion method for producing aromatics-rich gasoline. This involves splitting catalytically cracked light cycle oil (LCO) into light and heavy fractions. The heavy fraction is then hydrogenated and fed into a catalytic cracking unit riser reactor with the light fraction for catalytic cracking, resulting in aromatics-rich gasoline. CN201310010219.5 discloses a catalytic conversion method for producing aromatic compounds from low-quality heavy cycle oil. Hydrogenated diesel, obtained by hydrogenating the low-quality heavy cycle oil, is fed into a catalytic cracking unit for catalytic cracking, resulting in catalytic gasoline rich in benzene, toluene, and xylene.

[0006] From the above-disclosed literature, it can be found that most technologies for producing more light aromatics use conventional riser reactors or simple variable-diameter riser reactors. There are problems of uneven solid distribution and wall effect with sparse middle and dense side walls in the reactors. Although they can reduce the residence time of oil and gas, they cannot effectively take into account the gas-solid contact efficiency, resulting in reduced conversion rate and low yield of ideal components. Summary of the Invention

[0007] The present application provides a dual-reactor catalytic cracking method and apparatus using a novel multi-stage variable diameter reactor as one of the reactors. The main reactor can be used to process heavy petroleum hydrocarbons, and the secondary reactor is used to process or recycle light petroleum hydrocarbons, thereby enabling the maximum production of light aromatics.

[0008] The present application provides a device for catalytic conversion of petroleum hydrocarbons to produce light aromatics, comprising:

[0009] A reactor system, comprising:

[0010] Gas-solid separation equipment,

[0011] Stripping equipment,

[0012] main reactor, and

[0013] Auxiliary reactor;

[0014] The outlet of the main reactor and the outlet of the auxiliary reactor are connected to the gas-solid separation device, so as to separate the logistics from the main reactor and the auxiliary reactor into the spent catalyst and oil gas in the gas-solid separation device;

[0015] The stripping device is connected to the gas-solid separation device and is used to strip the spent catalyst separated by the gas-solid separation device in the stripping device; and

[0016] The main reactor is provided with a heavy crude oil feed port, and the auxiliary reactor is provided with a light crude oil feed port;

[0017] A regenerator, the regenerator is connected to the stripping equipment and is used to allow the stripped catalyst to enter the regenerator for regeneration; the regenerator is also connected to the main reactor and the auxiliary reactor respectively, and is used to allow the regenerated catalyst regenerated by the regenerator to circulate back to the main reactor and the auxiliary reactor;

[0018] a product separation system, the product separation system being in communication with the gas-solid separation device and configured to separate the oil and gas from the gas-solid separation device in the product separation system;

[0019] Wherein, at least one of the main reactor and the auxiliary reactor is a variable diameter catalytic conversion reactor,

[0020] The variable diameter catalytic conversion reactor comprises:

[0021] A reaction zone, the reaction zone comprising at least one variable diameter reaction section and an inlet and an outlet communicated with the variable diameter reaction section;

[0022] a pre-lift section, the pre-lift section being in communication with the inlet;

[0023] a first transition section, the first transition section being disposed between the pre-lift section and the inlet and being in communication with the pre-lift section and the inlet, respectively, the inner diameter of an end of the first transition section connected to the pre-lift section being smaller than the inner diameter of an end of the first transition section connected to the inlet;

[0024] an outlet region communicating with the outlet; and

[0025] a second transition section, the second transition section being disposed between the outlet area and the outlet and being in communication with the outlet area and the outlet, respectively, the inner diameter of an end of the second transition section connected to the outlet area being smaller than the inner diameter of an end of the second transition section connected to the outlet;

[0026] The inner cavity wall of the variable diameter reaction section is an arc-shaped wall; the cross-sectional area along the radial direction of the variable diameter reaction section gradually increases from the two ends of the variable diameter reaction section to the middle of the variable diameter reaction section.

[0027] In one embodiment, the inner diameter D of the variable diameter reaction section is determined by the following formula:

[0028] D=Ah 6 +Bh 5 +Ch 4 +Dh 3 +Eh 2 +Fh+G;

[0029] Wherein, D is the inner diameter of the variable diameter reaction section at the height h, in m;

[0030] h is the height of the variable diameter reaction section, in meters;

[0031] A is -50000 to 50000, preferably -40000 to 40000;

[0032] B is -10000 to 10000, preferably -8000 to 8000;

[0033] C is -6000 to 6000, preferably -3000 to 3000;

[0034] D is -5000 to 5000, preferably -1000 to 1000;

[0035] E is -1000 to 0, preferably -500 to 0, more preferably -100 to 0, and E is not equal to 0;

[0036] F is 0-2000, preferably 0-1000, more preferably 0-100, and F is not equal to 0;

[0037] G is 0-1000, preferably 0-100, more preferably 0-10, and G is not equal to 0;

[0038] In one embodiment, in the same variable diameter reaction section, the maximum inner diameter of the variable diameter reaction section is 1 to 5 times the minimum inner diameter of the variable diameter reaction section, the maximum inner diameter of the variable diameter reaction section is 0.2m to 10m, and the height of a certain section of the variable diameter reaction section is 1 to 5 times the maximum inner diameter of the variable diameter reaction section.

[0039] In one embodiment, the reaction zone includes multiple sections of the variable diameter reaction section, wherein, along the direction from the inlet to the outlet, the maximum inner diameter of the variable diameter reaction section arranged adjacent to the outlet is not less than the maximum inner diameter of the variable diameter reaction section arranged adjacent to the inlet.

[0040] In one embodiment, in the axial direction of the pre-lifting section, the ratio of the height of the pre-lifting section to the overall height of the catalytic conversion reactor is 0.01:1 to 0.1:1, and the inner diameter of the pre-lifting section shaft is 0.05m to 3m.

[0041] In one embodiment, the inner diameter of the end of the first transition section connected to the inlet is 1.1 to 5 times the inner diameter of the end of the first transition section connected to the pre-lifting section; in the axial direction of the pre-lifting section, the ratio of the height of the first transition section to the height of the pre-lifting section is 0.1:1 to 3:1.

[0042] In one embodiment, in the axial direction of the variable diameter reaction section, the ratio of the height of the reaction zone to the total height of the catalytic conversion reactor is 0.1:1 to 0.9:1; the ratio of the maximum inner diameter of the largest variable diameter reaction section among the multiple variable diameter reaction sections to the total height of the catalytic conversion reactor is 0.001:1 to 0.1:1.

[0043] In one embodiment, in the axial direction of the second transition section, the ratio of the second transition section to the total height of the catalytic conversion reactor is 0.01:1 to 0.1:1.

[0044] In one embodiment, the inner diameter of the outlet zone is 0.05 m to 3 m; in the axial direction of the outlet zone, the ratio of the height of the outlet zone to the total height of the catalytic conversion reactor is 0.01:1 to 0.2:1.

[0045] In one embodiment, one of the main reactor and the secondary reactor is the variable diameter catalytic converter reactor, and the other of the main reactor and the secondary reactor is a riser reactor.

[0046] In one embodiment, at least one fraction pipeline of the product separation system is connected to the feed oil inlet of the secondary reactor, so as to circulate at least one fraction passing through the product separation system back to the secondary reactor.

[0047] The present application also provides a method for catalytic conversion of petroleum hydrocarbons to produce light aromatics, using the above-mentioned device of the present application, wherein:

[0048] Feeding heavy oil feedstock to the primary reactor, and

[0049] feeding light crude oil into the secondary reactor;

[0050] Light aromatics are obtained through the product separation system.

[0051] In one embodiment, the reaction conditions in the main reactor include: reaction temperature of 550-700° C., agent / oil weight ratio of 5-20, reaction pressure of 0.1-1.0 MPa, and oil-gas residence time of 1-15 seconds.

[0052] In one embodiment, the reaction conditions of the secondary reactor include: reaction temperature of 500-650° C., agent / oil mass ratio of 1-20, oil-gas residence time of 1-10 seconds, and reaction pressure of 0.15-0.4 MPa.

[0053] In one embodiment, at least one of the main reactor and the auxiliary reactor is the variable diameter catalytic conversion reactor, and the average apparent linear velocity of oil and gas in the variable diameter catalytic conversion reactor is 0.5-5.0 m / s.

[0054] In one embodiment, the heavy feedstock oil is selected from one or more of straight-run wax oil, coker wax oil, deasphalted oil, hydrorefined oil, hydrocracking tail oil, vacuum residue oil, and atmospheric residue oil; the light feedstock oil is selected from one or more of heavy gasoline fraction, FCC light cycle oil light fraction, and hydrogenated diesel.

[0055] In one embodiment, at least one fraction line of the product separation system is connected to the feedstock oil inlet of the secondary reactor, so that at least a portion of the light feedstock oil fed to the secondary reactor comes from the product separation system.

[0056] The device and method of the present invention uses a variable diameter reactor as at least one of the main and auxiliary reactors to carry out a catalytic conversion method, which has the following advantages:

[0057] 1. The method and apparatus of the present invention use a variable diameter reactor of the present invention. Compared with conventional catalytic cracking reactors, the variable diameter reactor adopts a multi-section arc-shaped variable diameter design to change the linear velocity of oil and gas in the reaction zone, strengthen gas-solid disturbance and contact, enhance heat and mass transfer efficiency, and improve reaction conversion rate.

[0058] 2. The method and device of the present invention use the variable diameter reactor of the present invention. The variable diameter reactor adopts a multi-section arc-shaped variable diameter design, which can effectively reduce the reactor side wall effect, make the catalyst particles in the reactor more evenly distributed, effectively reduce the agglomeration of catalyst particles, etc., make the reaction more stable, and improve product selectivity.

[0059] 3. The method and apparatus of the present invention use a variable diameter reactor of the present invention, the volume of which can be smaller than that of a conventional riser reactor, thereby reducing the oil and gas residence time while maintaining high gas-solid contact efficiency, conforming to the molecular reaction laws of catalytic conversion of hydrocarbons to produce ethylene, propylene, and light aromatics, and improving yield and selectivity. BRIEF DESCRIPTION OF THE DRAWINGS

[0060] The following figures are intended only to illustrate and explain the present application and are not intended to limit the scope of the present application.

[0061] Figure 1 This is a structural diagram of the first embodiment of the variable diameter reactor described in this application;

[0062] Figure 2 1 is a structural schematic diagram of a second embodiment of the variable diameter reactor described in this application;

[0063] Figure 3 is a schematic cross-sectional view of the reaction zone of the variable diameter reactor of the present application, showing three variable diameter reaction sections connected in series;

[0064] Figure 41 is a partial cross-sectional structural schematic diagram of the variable diameter reaction section of the variable diameter reactor of the present application, showing the height h and the inner diameter D;

[0065] Figure 5 1 is a schematic structural diagram of a third embodiment of a variable diameter reactor of the present application;

[0066] Figure 6 1 is a schematic structural diagram of a fourth embodiment of a variable diameter reactor of the present application;

[0067] Figure 7 It is a schematic diagram of the device of the present application. DETAILED DESCRIPTION

[0068] The present application will be further described in detail below through the accompanying drawings and examples, through which the features and advantages of the present application will become more clear and distinct.

[0069] The word "exemplary" is used exclusively herein to mean "serving as an example, example, or illustration." Any embodiment described herein as "exemplary" is not necessarily to be construed as preferred or advantageous over other embodiments. Although various aspects of the embodiments are shown in the drawings, the drawings are not necessarily drawn to scale unless otherwise indicated.

[0070] In addition, the technical features involved in different embodiments of the present application described below can be combined with each other as long as they do not conflict with each other.

[0071] like Figure 7 As shown, the present application provides a device for catalytic conversion of petroleum hydrocarbons to produce light aromatics, comprising:

[0072] A reactor system, comprising:

[0073] Gas-solid separation equipment 100,

[0074] Stripping equipment 200,

[0075] Main reactor I, and

[0076] Secondary reactor II;

[0077] The outlet of the main reactor I and the outlet of the auxiliary reactor II are connected to the gas-solid separation device 100, so as to separate the logistics from the main reactor I and the auxiliary reactor II into the spent catalyst and oil gas in the gas-solid separation device 100;

[0078] The stripping device 200 is connected to the gas-solid separation device 100 and is used to strip the spent catalyst separated by the gas-solid separation device 100 in the stripping device 200; and

[0079] The main reactor I is provided with a heavy crude oil feed port, and the secondary reactor II is provided with a light crude oil feed port;

[0080] The regenerator 300 is connected to the stripping device 200 and is used to allow the stripped catalyst to enter the regenerator 300 for regeneration; the regenerator 300 is also connected to the main reactor I and the auxiliary reactor II respectively, and is used to circulate the regenerated catalyst regenerated by the regenerator back to the main reactor I and the auxiliary reactor II;

[0081] The product separation system 1000 is connected to the gas-solid separation device 100 and is used to separate the oil and gas from the gas-solid separation device 100 in the product separation system 1000.

[0082] In the present application, at least one of the main reactor I and the secondary reactor II is a variable diameter catalytic conversion reactor. The variable diameter catalytic conversion reactor used in the present application is described in detail below.

[0083] like Figure 1 and Figure 2 The variable diameter catalytic conversion reactor used in the present application includes: a reaction zone 10 , a pre-lift section 20 , a first transition section 30 , an outlet section 40 and a second transition section 50 .

[0084] like Figure 3 As shown, the reaction zone 10 includes at least one variable diameter reaction section 11 and an inlet 12 and an outlet 13 connected to the variable diameter reaction section 11. In a specific embodiment of the present application, the reaction zone 10 includes 3 to 5 variable diameter reaction sections 11. The reaction zone may include a plurality of sub-areas, each sub-area includes a variable diameter reaction section, and a plurality of sub-areas are connected in series to form a reaction zone. Alternatively, the reaction zone may be an integral structure formed by integral casting. A gas distributor may be provided between adjacent variable diameter reaction sections, and a gas distributor may be provided between the variable diameter reaction section and the first transition section. As shown Figure 5 As shown, in one embodiment of the present application, one or more supplementary catalyst inlets 60 can be independently set at any position of each variable diameter reaction section 11; preferably, the supplementary catalyst inlet 60 is located in the middle position of the variable diameter reaction section 11; the operating conditions of each variable diameter reaction section 11 are controlled in sections.

[0085] like Figure 1 and Figure 2 As shown, the pre-lift section 20 is in communication with the inlet 12 .

[0086] like Figure 1 and Figure 2As shown, the first transition section 30 is disposed between the pre-elevation section 20 and the inlet 12, and is in communication with the pre-elevation section 20 and the inlet 12, respectively. The inner diameter of the end of the first transition section 30 connected to the pre-elevation section 20 is smaller than the inner diameter of the end of the first transition section 30 connected to the inlet 12. The first transition section 30 is generally truncated cone-shaped and is located at the bottom end of the reaction zone 10. The raw material feed port is located in the middle of the first transition section.

[0087] like Figure 1 and Figure 2 As shown, the outlet area 40 is connected to the outlet 13. Figure 6 As shown, the outlet section of the outlet area 40 may be open or directly connected to the cyclone separator 70 .

[0088] like Figure 1 and Figure 2 As shown, the second transition section 50 is disposed between the outlet region 40 and the outlet 13 and is in communication with both the outlet region 40 and the outlet 13. The inner diameter of the end of the second transition section 50 connected to the outlet region 40 is smaller than the inner diameter of the end of the second transition section 50 connected to the outlet 13. The second transition section 50 is generally truncated cone-shaped and is located at the top of the reaction zone 10.

[0089] like Figure 3 As shown, the inner wall of the variable diameter reaction section 11 is an arc-shaped wall. In the radial direction of the variable diameter reaction section 11, the cross-sectional area of ​​the variable diameter reaction section 11 gradually increases from the two ends of the variable diameter reaction section 11 to the middle of the variable diameter reaction section 11. The shape of the variable diameter reaction section 11 is similar to an ellipse.

[0090] For the variable diameter catalytic conversion reactor of the present application, when the oil and gas enter the variable diameter reaction section 11 from the inlet 12 of the variable diameter reaction section 11, due to the variable diameter design of the variable diameter reaction section 11, the inner diameter of the variable diameter reaction section 11 gradually increases, and the speed of the oil and gas gradually decreases. The oil and gas will diffuse along the cavity wall of the variable diameter reaction section 11 and slowly fill the variable diameter reaction section 11, so that the oil and gas are fully in contact with the solid catalyst, the gas-solid disturbance is strengthened, and the heat and mass transfer efficiency is enhanced, thereby improving the reaction conversion rate of the oil and gas. When the oil and gas move from the middle of the variable diameter reaction section 11 to the outlet 13 of the variable diameter reaction section 11, due to the variable diameter design of the variable diameter reaction section 11, the inner diameter of the variable diameter reaction section 11 gradually becomes smaller, and the speed of the oil and gas gradually increases, and the oil and gas will gather toward the axis of the variable diameter reaction section 11 along the cavity wall of the variable diameter reaction section 11; that is, when the gas passes through the variable diameter reaction section 11 with a variable diameter setting, the arc-shaped wall surface of the variable diameter reaction section 11 guides the oil and gas, causing the oil and gas to diffuse into the variable diameter reaction section 11 first, and then gather in the middle of the variable diameter reaction section 11, thereby effectively reducing the side wall effect of the reaction zone 10, making the catalyst particles in the variable diameter reaction section 11 more evenly distributed, effectively reducing the agglomeration of catalyst particles, making the reaction more stable, and improving the selectivity of the product.

[0091] Therefore, compared with a variable diameter reactor with a locally constant inner diameter and a variable diameter reactor with a linear diameter change, the reactor of the present application with an arc-shaped diameter change can more obviously avoid the wall effect, strengthen the gas-solid disturbance, and thus enhance the heat and mass transfer efficiency.

[0092] like Figure 4 As shown, in a specific embodiment of the present application, in a specific variable diameter reaction section, the inner diameter D at the height h of the variable diameter reaction section can be determined by the following formula: D=Ah 6 +Bh 5 +Ch 4 +Dh 3 +Eh 2 +Fh+G,

[0093] Where D is the inner diameter of a certain point in the variable diameter reaction section, in m.

[0094] h is the height of the inner diameter D of the variable diameter reaction section (the distance from the lower end of the variable diameter reaction section), in meters.

[0095] A is -50000 to 50000, preferably -40000 to 40000;

[0096] B is -10000 to 10000, preferably -8000 to 8000;

[0097] C is -6000 to 6000, preferably -3000 to 3000;

[0098] D is -5000 to 5000, preferably -1000 to 1000;

[0099] E is -1000 to 0, preferably -500 to 0, more preferably -100 to 0, and E is not equal to 0;

[0100] F is 0-2000, preferably 0-1000, more preferably 0-100, and F is not equal to 0;

[0101] G is 0-1000, preferably 0-100, more preferably 0-10, and G is not equal to 0.

[0102] In the same variable diameter reaction section, the height of a certain section of the variable diameter reaction section is 1 to 5 times the maximum inner diameter of the variable diameter reaction section; preferably, the height of a certain section of the variable diameter reaction section is 1.5 to 3 times the maximum inner diameter of the variable diameter reaction section.

[0103] In one embodiment of the present application, in the same variable diameter reaction section, the maximum inner diameter of the variable diameter reaction section is 1 to 5 times the minimum inner diameter of the variable diameter reaction section; preferably, the maximum inner diameter of the variable diameter reaction section is 1.2 to 2 times the minimum inner diameter of the variable diameter reaction section.

[0104] In the same variable diameter reaction section, the maximum inner diameter of the variable diameter reaction section is 0.2m to 10m; preferably, the maximum inner diameter of the variable diameter reaction section is 2m to 5m.

[0105] like Figure 3 As shown, in one embodiment of the present application, along the direction from the inlet 12 to the outlet 13, the maximum inner diameter of the variable diameter reaction section 11 disposed adjacent to the outlet 13 is not less than the maximum inner diameter of the variable diameter reaction section 11 disposed adjacent to the inlet 12. That is, when there are multiple variable diameter reaction sections, along the direction from the inlet 12 to the outlet 13, the maximum cross-sectional inner diameter of the subsequent variable diameter reaction section is greater than or equal to the maximum cross-sectional inner diameter of the preceding variable diameter reaction section; in other words, the maximum cross-sectional inner diameter of each variable diameter reaction section gradually increases, or the maximum cross-sectional inner diameter of each variable diameter reaction section is equal.

[0106] By limiting the size of the variable diameter reaction section, the oil and gas entering the variable diameter reaction section can further enhance the disturbance and contact rate between the oil and gas and the catalyst, thereby increasing heat and mass transfer efficiency and improving reaction conversion rate. Furthermore, it can further effectively reduce the side wall effect of the reaction zone, making the catalyst particles more evenly distributed within the reactor and effectively reducing catalyst particle agglomeration, making the reaction more stable and improving product selectivity.

[0107] In one embodiment of the present application, in the axial direction of the variable diameter reaction section, the ratio of the total height of the reaction zone to the total height of the catalytic conversion reactor is 0.1:1 to 0.9:1; preferably, the ratio of the total height of the reaction zone to the total height of the catalytic conversion reactor is 0.2:1 to 0.8:1.

[0108] The ratio of the maximum inner diameter of the largest variable diameter reaction section among the multiple variable diameter reaction sections to the total height of the catalytic conversion reactor is 0.001:1 to 0.1:1; preferably, the ratio of the maximum inner diameter of the largest variable diameter reaction section among the multiple variable diameter reaction sections to the total height of the catalytic conversion reactor is 0.01:1 to 0.05:1.

[0109] In one embodiment of the present application, in the axial direction of the pre-lifting section, the ratio of the height of the pre-lifting section to the overall height of the catalytic conversion reactor is 0.01:1 to 0.1:1; preferably, the ratio of the height of the pre-lifting section to the overall height of the catalytic conversion reactor is 0.02:1 to 0.08:1.

[0110] The inner diameter of the pre-lifting section is 0.05m to 3m; preferably, the inner diameter of the pre-lifting section is 0.1m to 2m.

[0111] In one embodiment of the present application, in the axial direction of the pre-lifting section, the ratio of the height of the first transition section to the height of the pre-lifting section is 0.1:1 to 3:1; preferably, the ratio of the height of the first transition section to the height of the pre-lifting section is 0.5:1 to 1.5:1.

[0112] In one embodiment of the present application, in the axial direction of the second transition section, the ratio of the second transition section to the total height of the catalytic conversion reactor is 0.01:1 to 0.1:1; preferably, the ratio of the second transition section to the total height of the catalytic conversion reactor is 0.02:1 to 0.05:1.

[0113] In one embodiment of the present application, the inner diameter of the outlet zone is 0.05 m to 3 m; preferably, the inner diameter of the outlet zone is 0.1 m to 2 m.

[0114] In the axial direction of the outlet zone, the ratio of the height of the outlet zone to the total height of the catalytic converter reactor is 0.01:1 to 0.2:1; preferably, the ratio of the height of the outlet zone to the total height of the catalytic converter reactor is 0.05:1 to 0.18:1.

[0115] By setting the above parameters, the radial dimensions, axial dimensions and proportional relationships between the various components of the pre-lifting section, the first transition section, the variable diameter reaction section, the second transition section and the outlet zone are respectively defined, thereby limiting the overall size of the catalytic conversion reactor. Compared with existing reactors, the catalytic conversion reactor has a smaller volume, which can effectively reduce the residence time of oil and gas in the catalytic conversion reactor, while ensuring sufficient contact between oil and gas and the catalyst, complying with the molecular reaction laws of catalytic conversion of hydrocarbons to produce ethylene, propylene and light aromatics, and improving yield and selectivity.

[0116] In one embodiment, both the main reactor I and the secondary reactor II are the variable diameter catalytic conversion reactors.

[0117] In one embodiment, one of the main reactor I and the secondary reactor II is the variable diameter catalytic conversion reactor, and the other of the main reactor I and the secondary reactor II is a riser reactor. The riser reactor can be selected from one of a constant diameter riser, a constant linear velocity riser, a variable diameter riser, a variable linear velocity riser, or a fluidized bed, or can be a composite reactor consisting of a constant diameter riser, a variable diameter riser, and a fluidized bed, preferably a constant diameter riser reactor.

[0118] The spent catalysts from both reactors can be charred and regenerated in the same regenerator 300 under an oxygen-containing atmosphere, with some of the hot regenerated catalyst being directly returned to the primary reactor I and the secondary reactor II for recycling. Alternatively, multiple regenerators 300 can be used, with one regenerator 300 used to regenerate the spent catalyst in the primary reactor I and another regenerator 300 used to regenerate the spent catalyst in the secondary reactor II.

[0119] The gas-solid separation equipment 100, the stripping equipment 200 and the regenerator 300 used in this application can use various equipment known in the art, and their connection relationship with the main reactor I and the sub-reactor II can also be connected according to methods known in the art, which will not be repeated here.

[0120] The product separation system 1000 of the present application can separate oil and gas to obtain various fractions. Figure 7 As shown, in one embodiment, the product separation system includes a fractionation tower 400. Oil and gas are transported to the fractionation tower 400 to obtain C1-C4 gas 408, a light naphtha fraction 409, a light aromatics-rich gasoline fraction 410, a heavy gasoline fraction 411, a light cycle oil fraction 412, a light cycle oil fraction 413, and a slurry oil 414 according to the distillation range. In one embodiment, the product separation system 1000 also includes an aromatics extraction unit 500. The light aromatics-rich gasoline fraction 410 can be transported to the aromatics extraction unit 500 for extraction, and light aromatics such as BTX 518 and aromatic raffinate 517 can be output as products. In one embodiment, the separated light aromatics-rich gasoline fraction 410 is subjected to adsorption and / or selective hydrogenation to remove impurities such as nitrogen compounds, sulfides, colloids, and / or heavy metals, and then transported to the aromatics extraction unit 500 for extraction. In one embodiment, the product separation system further includes a hydrogenation unit 600, which can transport the light cycle oil heavy fraction 413 to the hydrogenation unit 600 for hydrogenation. The hydrogenated product can then enter the fractionation tower 400 for separation. The separation in the fractionation tower 400, the hydrogenation in the hydrogenation unit 600, and the aromatics extraction in the aromatics extraction unit 500 can be performed using equipment, methods, and conditions known in the art and will not be further described herein.

[0121] In one embodiment, at least one fraction pipeline of the product separation system 1000 is connected to the feed oil inlet of the secondary reactor II, so as to allow at least one fraction processed by the product separation system 1000 to be recycled back to the secondary reactor II.

[0122] In one embodiment, the heavy gasoline fraction 411 pipeline and / or the light cycle oil light fraction 412 pipeline of the distillation tower 400 are connected to the raw oil inlet of the secondary reactor II, so that the heavy gasoline fraction 411 and / or the light cycle oil light fraction 412 separated by the distillation tower 400 enter the secondary reactor II for recycling.

[0123] In one embodiment, the material outlet pipeline 619 of the hydrogenation device 600 is connected to the feed oil inlet of the secondary reactor II, so that the hydrogenated light cycle oil heavy fraction 413 enters the secondary reactor II for recycling.

[0124] Therefore, the present application provides a method for catalytic conversion of petroleum hydrocarbons to produce light aromatics, using the device of the present application, wherein:

[0125] Feeding heavy crude oil into the main reactor I, and

[0126] Feeding light crude oil into the secondary reactor II;

[0127] Light aromatics are obtained through the product separation system 1000.

[0128] In one embodiment, the reaction conditions in the main reactor include: reaction temperature of 550-700° C., agent / oil weight ratio of 5-20, reaction pressure of 0.1-1.0 MPa, and oil-gas residence time of 1-15 seconds.

[0129] In one embodiment, the reaction conditions of the secondary reactor include: reaction temperature of 500-650° C., agent / oil mass ratio of 1-20, oil-gas residence time of 1-10 seconds, and reaction pressure of 0.15-0.4 MPa.

[0130] In one embodiment, when at least one of the main reactor and the auxiliary reactor is the variable diameter catalytic conversion reactor, the average apparent linear velocity of oil and gas in the variable diameter catalytic conversion reactor is 0.5-5.0 m / s.

[0131] In one embodiment, the heavy feedstock oil can be selected from one or more of straight-run wax oil, coker wax oil, deasphalted oil, hydrorefined oil, hydrocracking tail oil, vacuum residue oil, and atmospheric residue oil. The light feedstock oil can be selected from one or more of heavy gasoline fraction, FCC light cycle oil light fraction, and hydrotreated diesel oil.

[0132] As described above, in one embodiment, at least one fraction line of the product separation system is connected to the feedstock oil inlet of the secondary reactor, so that at least a portion of the light feedstock oil fed to the secondary reactor comes from the product separation system of the device of the present application. Of course, the light feedstock oil can also be catalytically cracked light cycle oil (HLCO) after hydrogenation from outside the device of the present application.

[0133] As described above, the pipeline for the heavy gasoline fraction 411 and / or the pipeline for the light cycle oil fraction 412 of the fractionation tower 400 can be connected to the feedstock oil inlet of the secondary reactor II, so that the heavy gasoline fraction 411 and / or the light cycle oil fraction 412 with a distillation range of 170-250° C. obtained by the fractionation tower 400 can enter the secondary reactor II for recycling. In one embodiment, the material outlet pipeline 619 of the hydrogenation equipment 600 can be connected to the feedstock oil inlet of the secondary reactor II, so that the hydrogenated light cycle oil heavy fraction 413 can enter the secondary reactor II for recycling.

[0134] The catalytic cracking catalyst used in the present invention may be any catalyst known in the art, for example, it may include zeolite, inorganic oxide, and optionally clay. The content of each component may be: 10-50 wt% zeolite, 5-90 wt% inorganic oxide, 0-70 wt% clay. The active component is one, two, or more selected from Y, HY, USY, or Beta zeolites, which may or may not contain rare earth elements.

[0135] like Figure 7 As shown, heavy oil (e.g., hydrogenated residue) I1 enters the main reactor I under the action of steam I2, and undergoes catalytic cracking reaction under the reaction atmosphere of fluidized medium I3, and light oil (e.g., hydrogenated LCO) II26 enters the secondary reactor II, and undergoes catalytic cracking reaction under the conditions of fluidized medium II4.

[0136] The regenerated catalyst after the reaction is subjected to gas-solid separation. The separated regenerated catalyst is input into the stripping medium 205 through the stripping equipment 200 for stripping and then transported to the regenerator 300 through pipeline 223. The regenerated catalysts from the two reactors can be charred and regenerated under the atmosphere of oxygen-containing gas 321 in the same regenerator. Part of the hot regenerated catalyst is directly returned to the main reactor I and the auxiliary reactor II through pipelines 324 and 325 for recycling, and the flue gas is discharged through pipeline 322.

[0137] After reacting in the main riser I and secondary reactor II, the oil and gas are separated in gas-solid separation equipment 100 and then fed through pipeline 106 to a precision separation tower 400 for separation, yielding C1-C4 gas 408, a light naphtha fraction 409, a gasoline fraction rich in light aromatics 410, a heavy gasoline fraction 411, a light cycle oil fraction 412, a heavy fraction 413, and slurry oil 414. The separated light cycle oil heavy fraction 413 is hydrogenated in hydrogenation equipment 600, after which a portion is returned to the fractionation system 400 for precision separation. The separated light aromatics-rich gasoline fraction 410 is subjected to adsorption and / or selective hydrogenation to remove impurities such as nitrogen compounds, sulfides, gums, and / or heavy metals. The fraction is then fed to the aromatics extraction unit 500, where aromatics extraction yields light aromatics such as BTX 518 and aromatic raffinate 517 as output products.

[0138] The following examples will further illustrate the present invention, but are not intended to limit the present invention.

[0139] The properties of the raw oil used in Example 1 and Comparative Example 1 are listed in Table 1.

[0140] Example 1:

[0141] In this example, secondary reactor II employed the variable-diameter reactor provided by the present invention, while primary reactor I employed a conventional constant-diameter riser reactor. The feedstock for primary reactor I consisted of hydrogenated residual oil, preheated to 280°C, entering primary reactor I at a feed rate of 1.92 kg / h. The feedstock for secondary reactor II consisted of hydrogenated LCO, preheated to 380°C, entering secondary reactor II at a feed rate of 1.86 kg / h. The main operating parameters for both the primary and secondary reactors are shown in Table 2.

[0142] The specific structure of the variable diameter reactor used in this embodiment is as follows:

[0143] The variable diameter reactor includes: a reaction zone 10, a pre-lift section 20, a first transition section 30, an outlet zone 40, and a second transition section 50. The reaction zone includes three variable diameter sections, all of which have the same height of 0.12m. The maximum inner diameter of the first variable diameter section is 0.05m, and the maximum inner diameters of the second and third variable diameter sections are 0.062m. The heights of the first and second transition sections are 0.05m and 0.08m, respectively. The heights of the pre-lift section and the outlet section are both 0.2m. The diameters of the connection points from bottom to top are 0.032m, 0.043m, 0.043m, and 0.043m, respectively. The diameters of the pre-lift section and the outlet section are both 0.016m.

[0144] In the first diameter-changing section, the inner diameter D (in m) at a certain point is determined according to its height h (in m) as follows:

[0145] D (unit: m) = Ah 6 +Bh 5 +Ch 4 +Dh 3 +Eh 2 +Fh+G,

[0146] A is 2676.7; B is -946.15; C is 117.81; D is -5.42;

[0147] E is -3.25; F is 0.48; G is 0.032.

[0148] In the second diameter-changing section, the inner diameter D (unit: m) at a certain point is determined according to its height h (unit: m) as follows:

[0149] D (unit: m) = Ah 6 +Bh 5 +Ch 4 +Dh 3 +Eh 2 +Fh+G,

[0150] A is -3355.1; B is 1285; C is -220.06; D is 20.47;

[0151] E is -6.33; F is 0.66; G is 0.043.

[0152] In the third diameter-reducing section, the inner diameter D (in m) at a certain point along with its height h (in m) is the same as that of the second section and is determined as follows:

[0153] D (unit: m) = Ah 6 +Bh 5 +Ch 4 +Dh 3 +Eh 2 +Fh+G,

[0154] A is -3355.1; B is 1285; C is -220.06; D is 20.47;

[0155] E is -6.33; F is 0.66; G is 0.043.

[0156] Comparative Example 1

[0157] The catalytic conversion method of this comparative example 1 is the same as that of Example 1, except that the secondary reactor II is a conventional catalytic cracking riser reactor. The operating parameters of the apparatus of comparative example 1 are shown in Table 2.

[0158] Comparing the product distributions of Example 1 and Comparative Example 1 (Table 3), it can be seen that the catalytic conversion using the apparatus provided by the present invention comprising a variable diameter reactor can increase the single-pass conversion rate by 1.6% while reducing the oil and gas residence time. The BTX yield is increased to 14.11%, a 1.0% increase compared to the comparative example. The BTX content of the product is also increased to 31.36%, a 1.3% increase compared to the comparative example. For the catalytic cracking field, an increase of approximately 1% in conversion rate and BTX yield can significantly increase economic benefits.

[0159] In the description of this application, it should be noted that, unless otherwise specified or limited, the terms "installed", "connected", and "connected" should be understood in a broad sense. For those skilled in the art, the specific meanings of the above terms in this application can be understood according to specific circumstances.

[0160] The present application has been described above in conjunction with preferred embodiments, but these embodiments are merely exemplary and serve only as an illustrative example. On this basis, various replacements and improvements can be made to the present application, all of which fall within the scope of protection of the present application.

[0161] Table 1 Properties of crude oil

[0162]

[0163] Table 2 Operating parameters

[0164]

[0165] Table 3 Product distribution

[0166] Example 1 Comparative Example 1 Product distribution, weight % dry gas 4.92 5.15 Liquefied gas 18.80 18.41 gasoline 44.92 43.69 diesel fuel 19.79 21.41 Slurry 2.96 3.07 coke 8.63 8.28 Conversion rate, weight % 78.14 76.52 BTX content, weight % 31.36 30.07 benzene 2.25 2.28 Toluene 12.20 11.76 Xylene 16.91 16.03 BTX yield, weight % 14.11 13.14 benzene 1.01 1.00 Toluene 5.50 5.14 Xylene 7.60 7.00 gasoline octane number RON 100.5 100.2 MON 90.0 89.7

Claims

1. A device for catalytic conversion of petroleum hydrocarbons to produce light aromatics, comprising: A reactor system, the reactor system comprising: Gas-solid separation equipment, Stripping equipment, The main reactor, and Auxiliary reactor; The outlet end of the main reactor and the outlet end of the auxiliary reactor are connected to the gas-solid separation device, so as to separate the logistics from the main reactor and the auxiliary reactor into the catalyst to be produced and the oil and gas in the gas-solid separation device; The stripping device is connected to the gas-solid separation device, and is used to strip the spent catalyst separated by the gas-solid separation device in the stripping device; and The main reactor is provided with a heavy crude oil feed port, and the auxiliary reactor is provided with a light crude oil feed port; A regenerator, the regenerator is connected to the stripping device, and is used to allow the stripped catalyst to enter the regenerator for regeneration; the regenerator is also connected to the main reactor and the auxiliary reactor, respectively, and is used to allow the regenerated catalyst regenerated by the regenerator to circulate back to the main reactor and the auxiliary reactor; a product separation system, the product separation system being in communication with the gas-solid separation device and being used for separating the oil and gas from the gas-solid separation device in the product separation system; Wherein, at least one of the main reactor and the secondary reactor is a variable diameter catalytic conversion reactor, The variable diameter catalytic conversion reactor comprises: A reaction zone, the reaction zone comprising at least one variable diameter reaction section and an inlet and an outlet connected to the variable diameter reaction section; a pre-lift section, the pre-lift section being in communication with the inlet; A first transition section, the first transition section is arranged between the pre-lifting section and the inlet, and is communicated with the pre-lifting section and the inlet respectively, and the inner diameter of one end of the first transition section connected to the pre-lifting section is smaller than the inner diameter of one end of the first transition section connected to the inlet; an outlet region communicating with the outlet; and a second transition section, the second transition section being disposed between the outlet area and the outlet and being in communication with the outlet area and the outlet, respectively, the inner diameter of an end of the second transition section connected to the outlet area being smaller than the inner diameter of an end of the second transition section connected to the outlet; The inner cavity wall of the variable diameter reaction section is an arc-shaped wall; the cross-sectional area along the radial direction of the variable diameter reaction section gradually increases from the two ends of the variable diameter reaction section to the middle of the variable diameter reaction section; The inner diameter D of the variable diameter reaction section is determined by the following formula: D=Ah 6 +Bh 5 +Ch 4 +Dh 3 +Eh 2 +Fh+G; Wherein, D is the inner diameter at the height h of the variable diameter reaction section, in m; h is the height of the variable diameter reaction section, in m; A is -50000 to 50000; B is -10000~10000; C is -6000~6000; D is -5000~5000; E is -1000~0, and E is not equal to 0; F is 0-2000, and F is not equal to 0; G is 0-1000, and G is not equal to 0; In the same variable diameter reaction section, the maximum inner diameter of the variable diameter reaction section is 1 to 5 times the minimum inner diameter of the variable diameter reaction section, the maximum inner diameter of the variable diameter reaction section is 0.2m to 10m, and the height of a certain section of the variable diameter reaction section is 1 to 5 times the maximum inner diameter of the variable diameter reaction section; The reaction zone includes a plurality of variable diameter reaction sections, wherein, along the direction from the inlet to the outlet, the maximum inner diameter of the variable diameter reaction section disposed adjacent to the outlet is not less than the maximum inner diameter of the variable diameter reaction section disposed adjacent to the inlet.

2. The device according to claim 1, characterized in that In the axial direction of the pre-lifting section, the ratio of the height of the pre-lifting section to the overall height of the catalytic conversion reactor is 0.01:1 to 0.1:1, and the inner diameter of the pre-lifting section shaft is 0.05m to 3m.

3. The device according to claim 1, characterized in that The inner diameter of the end of the first transition section connected to the inlet is 1.1 to 5 times the inner diameter of the end of the first transition section connected to the pre-lifting section; in the axial direction of the pre-lifting section, the ratio of the height of the first transition section to the height of the pre-lifting section is 0.1:1 to 3:

1.

4. The device according to claim 1, characterized in that In the axial direction of the variable diameter reaction section, the ratio of the height of the reaction zone to the total height of the catalytic conversion reactor is 0.1:1 to 0.9:1; the ratio of the maximum inner diameter of the largest variable diameter reaction section among the multiple variable diameter reaction sections to the total height of the catalytic conversion reactor is 0.001:1 to 0.1:

1.

5. The device according to claim 1, characterized in that In the axial direction of the second transition section, the ratio of the second transition section to the total height of the catalytic conversion reactor is 0.01:1 to 0.1:

1.

6. The device according to claim 1, characterized in that The inner diameter of the outlet zone is 0.05m to 3m; in the axial direction of the outlet zone, the ratio of the height of the outlet zone to the total height of the catalytic conversion reactor is 0.01:1 to 0.2:

1.

7. The device according to any one of claims 1 to 6, wherein: One of the main reactor and the auxiliary reactor is the variable diameter catalytic conversion reactor, and the other of the main reactor and the auxiliary reactor is a riser reactor.

8. The device according to any one of claims 1 to 6, wherein: At least one fraction pipeline of the product separation system is connected to the feedstock oil inlet of the secondary reactor, so as to allow at least one fraction passing through the product separation system to circulate back to the secondary reactor.

9. The device according to claim 1, characterized in that A is -40000-40000.

10. The device according to claim 1, characterized in that B is -8000-8000.

11. The device according to claim 1, characterized in that C is -3000-3000.

12. The device according to claim 1, characterized in that D is -1000-1000.

13. The device according to claim 1, characterized in that E is -500-0, and E is not equal to 0.

14. The device according to claim 1, characterized in that E is -100-0, and E is not equal to 0.

15. The device according to claim 1, characterized in that F is 0-1000, and F is not equal to 0.

16. The device according to claim 1, characterized in that F is 0-100, and F is not equal to 0.

17. The device according to claim 1, characterized in that G is 0-100, and G is not equal to 0.

18. The device according to claim 1, characterized in that G is 0-10, and G is not equal to 0.

19. A method for producing light aromatics by catalytic conversion of petroleum hydrocarbons, carried out in the apparatus of any one of claims 1 to 18, wherein: Feeding heavy crude oil to the main reactor, and Feeding light crude oil into the secondary reactor; Light aromatics are obtained through the product separation system.

20. The method according to claim 19, wherein: The reaction conditions in the main reactor include: reaction temperature of 550-700°C, agent / oil weight ratio of 5-20, reaction pressure of 0.1-1.0 MPa, and oil-gas residence time of 1-15 seconds.

21. The method according to claim 19, wherein: The reaction conditions of the secondary reactor include: reaction temperature of 500-650° C., agent / oil mass ratio of 1-20, oil-gas residence time of 1-10 seconds, and reaction pressure of 0.15-0.4 MPa.

22. The method according to any one of claims 19 to 21, wherein: At least one of the main reactor and the auxiliary reactor is the variable diameter catalytic conversion reactor, and the average apparent linear velocity of oil and gas in the variable diameter catalytic conversion reactor is 0.5-5.0 m / s.

23. The method according to any one of claims 19 to 21, wherein: The heavy feedstock oil is selected from one or more of straight-run wax oil, coker wax oil, deasphalted oil, hydrorefined oil, hydrocracking tail oil, vacuum residue oil, and atmospheric residue oil; the light feedstock oil is selected from one or more of heavy gasoline fractions and hydrogenated diesel oil.

24. The method according to claim 23, wherein: At least one fraction pipeline of the product separation system is connected to the feedstock oil inlet of the secondary reactor, so that at least a portion of the light feedstock oil fed to the secondary reactor comes from the product separation system.

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