Method and system for producing aromatics by countercurrent hydrogenation of CO

Abstract

The present invention relates to the field of aromatic hydrocarbon production and discloses a method and system for producing aromatic hydrocarbons by countercurrent hydrogenation of CO. The method comprises: under hydrogenation conditions, a feed gas and an alkylating agent are introduced into a reactor in countercurrent flow, where they contact and react in at least two catalyst beds to produce a stream containing aromatic hydrocarbons; the catalyst in the catalyst bed comprises a metal oxide and a molecular sieve, with the weight ratio of the metal oxide to the molecular sieve decreasing stepwise along the flow of the feed gas. The method of the present invention significantly improves the selectivity for C9 aromatic hydrocarbons.

Description

Technical Field

[0001] The present invention relates to the technical field of preparation of aromatic hydrocarbons, and in particular to a method and system for preparing aromatic hydrocarbons by countercurrent hydrogenation of CO. Background Art

[0002] Aromatic hydrocarbons are important chemical raw materials and are widely used in the manufacture of synthetic resins, man-made fibers, synthetic rubber, and other products. Trimethylbenzene, among others, is primarily used in the production of pharmaceuticals, dyes, and synthetic resins. It is also used in the synthesis of trimellitic anhydride, pyromellitic dianhydride, water-soluble alkyd resins, unsaturated polyester resins, as well as plasticizers, epoxy resin curing agents, and surfactants. With petroleum resources dwindling, new non-petroleum-based technologies for producing aromatic hydrocarbons are being researched and developed. The production of aromatic hydrocarbons using H₂ / CO as feedstock is a typical technology. These routes fall into two main categories: methanol synthesis and Fischer-Tropsch synthesis. Fischer-Tropsch synthesis follows the Anderson-Schulz-Flory distribution, resulting in a broad product distribution and very low aromatics selectivity. In contrast, the methanol synthesis-methanol conversion combined reaction process, developed based on existing mature technologies, requires a longer production route and higher equipment investment.

[0003] In recent years, researchers have developed a novel direct CO hydrogenation to aromatics process based on bifunctional oxide and zeolite catalysts. This technology allows for the coupling of multiple cascade reactions and facilitates reaction equilibrium shifts. Compared to indirect, combined reaction processes, it also offers fixed cost advantages. Chem 2017, 3, 1-14; Chemical Communications 2017, 53, 11146-11149; ChemCatChem 2018, 10, 4519-4524; and ChemCatChem 2019, 11, 1-9, respectively, reported the application of Zn-Zr, Zn-Cr, Ce-Zr, and Mo-Zr oxide-ZSM-5 zeolite coupled catalysts in CO hydrogenation to aromatics, achieving aromatics selectivity as high as 83%. In a similar CO2 hydrogenation system, zinc-aluminum spinel coupled with an acidic zeolite catalyst also achieved aromatics selectivity approaching 80%. CN110496639A, Nature Communications 2018, 9, 3457 and ACS Catalysis 2019, 9, 895 achieved 34% CO2 conversion and 76% aromatics selectivity in the chromium oxide / ZSM-5 system. In the above CO / CO2 hydrogenation to aromatics system, the aromatics distribution is often wide, with a typical composition of 30% C6-C8, 50% C9 and 20% C 10+ Since the separation of different aromatic components is energy-intensive and difficult, achieving a narrower aromatic distribution has always been a challenge for the system.

[0004] Existing reports focus on the optimization and improvement of the selectivity of light aromatics. CN106540740A discloses the application of zirconium-containing composite oxide-modified zeolite molecular sieve in CO hydrogenation reaction, achieving high selectivity for light aromatics; CN107486234A uses modified cerium-zirconium solid solution-multi-level porous silica-alumina solid acid material, and also achieves high selectivity for light aromatics in CO hydrogenation to aromatics reaction. In addition to the screening and optimization of catalyst components, some studies have also focused on process optimization. For example, ACS Catalysis 2020, 10, 7389-7397 loaded USY or Beta molecular sieve into the lower layer of MnCr / ZSM-5 catalyst, and in the CO hydrogenation to aromatics system, increased the proportion of BTX from 30% to 80%; CN112295597A discloses a double-bed catalyst in CO conversion to BTX-rich aromatics and its application. About oxide-molecular sieve catalytic system C 9+ There is currently little research on improving the selectivity of aromatics.

[0005] ACS Catalysis 2019, 9, 2203-2212 High C was achieved using short b-axis ZSM-5 zeolite 10 There are few reports on the improvement of aromatics (durene) selectivity, but few reports on the improvement of C9 selectivity. In addition, a large number of studies have shown that the regulation of aromatics distribution is often accompanied by a decrease in aromatics selectivity.

[0006] In general, CO hydrogenation to aromatics can be achieved based on multifunctional catalyst systems of oxides and molecular sieves. However, how to achieve aromatic distribution regulation on the basis of higher aromatic selectivity, especially improving the selectivity of C9 aromatics, remains a challenge. Summary of the Invention

[0007] The purpose of the present invention is to overcome the problems of low aromatic selectivity, difficult to control aromatic distribution and low C9 aromatic selectivity in the preparation of aromatics, and to provide a method and system for preparing aromatics by countercurrent hydrogenation of CO. The method not only has high catalytic activity and aromatic selectivity, but also can realize the control of aromatic distribution. The aromatic selectivity can reach more than 70%, C6-C8 aromatics: C9 aromatics: C 10+ The ratio of aromatics can be optimized to a range of approximately 11:78:11, and the selectivity of C9 aromatics can reach a level higher than 70%.

[0008] To achieve the above objectives, the present invention provides, in a first aspect, a method for producing aromatics by countercurrent hydrogenation of CO, comprising: under hydrogenation conditions, allowing a feed gas and an alkylating agent to enter a reactor in a countercurrent manner, and contacting and reacting in at least two catalyst beds to produce a stream containing aromatics; the catalyst in the catalyst bed comprises a metal oxide and a molecular sieve, and the weight ratio of the metal oxide to the molecular sieve decreases step by step along the flow direction of the feed gas.

[0009] A second aspect of the present invention provides a system for producing aromatics by countercurrent hydrogenation of CO. The system comprises a reaction unit comprising at least two catalyst beds loaded in a reactor, for contacting a feed gas with an alkylating agent to react and produce a stream containing aromatics; a material inlet and outlet unit comprising a feed gas inlet, an alkylating agent feed pipe, and a product outlet, for respectively introducing the feed gas and the alkylating agent and withdrawing the stream containing aromatics; the feed gas inlet and the alkylating agent feed pipe being arranged so as to achieve opposite flow directions for the introduction of the feed gas and the introduction of the alkylating agent.

[0010] The technical solution of the present invention is that the weight ratio of metal oxide and molecular sieve in each catalyst bed decreases step by step, and the feed gas contacts the alkylating agent in countercurrent, triggering alkylation, isomerization and disproportionation of aromatic products, thereby optimizing the distribution of aromatic products in the CO hydrogenation reaction. It not only achieves high catalytic activity and aromatic selectivity, but also can flexibly control the distribution of aromatics. When the catalyst system is used in the CO hydrogenation reaction, the aromatic selectivity can reach more than 70%, and the C9 aromatic selectivity can reach a level higher than 70%, C6-C8 aromatics: C9 aromatics: C 10+ The ratio of aromatics can be optimized to a range of approximately 11:78:11, achieving good technical results. BRIEF DESCRIPTION OF THE DRAWINGS

[0011] Figure 1 The present invention relates to a system for producing aromatics by countercurrent hydrogenation of CO according to a preferred embodiment of the present invention.

[0012] Description of Reference Numerals

[0013] 1. Raw gas inlet; 2. Catalyst upper bed; 3. Catalyst lower bed; 4. Alkylating agent feed pipe; 5. Product outlet. DETAILED DESCRIPTION

[0014] The endpoints of the ranges and any values ​​disclosed herein are not limited to the precise ranges or values, and these ranges or values ​​should be understood to include values ​​close to these ranges or values. For numerical ranges, the endpoints of each range, the endpoints of each range and individual point values, and the individual point values ​​can be combined with each other to obtain one or more new numerical ranges, which should be considered to be specifically disclosed herein.

[0015] A first aspect of the present invention provides a method for producing aromatics by countercurrent hydrogenation of CO, comprising: under hydrogenation conditions, allowing a feed gas and an alkylating agent to enter a reactor in a countercurrent manner, and contacting and reacting in at least two catalyst beds to obtain a stream containing aromatics; the catalyst in the catalyst bed comprises a metal oxide and a molecular sieve, and the weight ratio of the metal oxide to the molecular sieve decreases step by step along the flow direction of the feed gas.

[0016] According to the present invention, those skilled in the art will understand that when the catalyst bed contains both metal oxides and molecular sieves, the metal oxides and molecular sieves can be mixed and formed and loaded into a reactor. In some embodiments, the mixed forming method of the metal oxides and molecular sieves is: first, the oxide and molecular sieve powders are uniformly mixed by grinding or the like, and then a binder is selectively added, and a molded catalyst is obtained by tableting or extrusion or the like. When the catalyst bed contains only molecular sieves, they can be molded alone and loaded into a reactor. In some embodiments, the molecular sieve is molded alone by: selectively adding a binder, and then a molded molecular sieve is obtained by tableting or extrusion or the like.

[0017] In the present invention, when a binder is used for molding, the binder content in the catalyst in each bed layer can be selected within a wide range, preferably 10-30% by weight.

[0018] According to the present invention, those skilled in the art can select a suitable molding method according to their needs, and no further details will be given in the present invention.

[0019] According to the present invention, the catalyst bed can be loaded into one or more reactors, and the specific configuration needs to be determined based on the actual production equipment.

[0020] In the present invention, by loading the catalyst bed in stages and reacting the feed gas with the alkylating agent in countercurrent, the hydrogenation reaction and the alkylation, isomerization and disproportionation of the aromatic products are triggered under the action of the catalyst, thereby optimizing the distribution of aromatic products in the CO hydrogenation reaction. Not only does it have high catalytic activity and aromatic selectivity, but it can also achieve regulation of aromatic distribution. When the catalyst system is used in the CO hydrogenation reaction, the aromatic selectivity can reach above 70%, and the C9 aromatic selectivity can reach a level higher than 70%, C6-C8 aromatics: C9 aromatics: C 10 The ratio of aromatics can be optimized to a range of approximately 11:78:11, achieving good technical results.

[0021] According to the present invention, as long as the purpose of the present invention can be achieved, the temperature of the catalyst bed is not limited. In some embodiments, the temperature of the catalyst bed at each stage is 300-580°C (for example, 300°C, 320°C, 350°C, 400°C, 450°C, 500°C, 560, 580°C, and any combination of values ​​within the above ranges).

[0022] According to the present invention, in some preferred embodiments, the temperature of each catalyst bed increases stepwise along the feed gas flow direction. The aforementioned embodiment can better modulate the aromatic hydrocarbon products of the CO hydrogenation reaction and increase the selectivity of aromatic hydrocarbons.

[0023] According to the present invention, in some preferred embodiments, the temperature difference between two adjacent catalyst beds is 20-200°C, more preferably 50-150°C (e.g., 50°C, 60°C, 80°C, 100°C, 110°C, 120°C, 130°C, 140°C, 150°C, and any combination thereof). The aforementioned embodiments can better trigger transalkylation reactions, such as alkylation and disproportionation, of the aromatic products, thereby increasing the selectivity for C9 aromatics.

[0024] According to the present invention, in some embodiments, the weight ratio difference between the metal oxide and the molecular sieve in two adjacent catalyst beds is 0.2-5.0. The aforementioned embodiment can better trigger transalkylation reactions such as alkylation and disproportionation of aromatic products, thereby increasing the selectivity of C9 aromatics.

[0025] According to the present invention, in some preferred embodiments, along the feed gas flow direction, the weight ratio of metal oxide to molecular sieve in the first-stage catalyst bed is (0.2-5.0):1, and the metal oxide content in the final-stage catalyst bed is zero. The aforementioned embodiments can better re-adjust the distribution of carbon monoxide products, thereby increasing CO conversion while also enhancing C9 aromatics selectivity.

[0026] According to the present invention, as long as the purpose of the present invention can be achieved, the specific number of stages of the catalyst bed is not limited. In some preferred embodiments, the catalyst bed is 2-5 stages, for example, 2, 3, 4 or 5 stages.

[0027] According to the present invention, as long as the objectives of the present invention can be achieved, the volume of each catalyst bed is not limited. In some embodiments, along the flow direction of the feed gas, the volume ratio of the catalyst beds of two adjacent stages is (0.5-10.0):1, preferably (1.0-5.0):1. The aforementioned embodiments can better increase the CO conversion rate while increasing the selectivity of aromatics, especially C9 aromatics.

[0028] According to the present invention, as long as the purpose of the present invention can be achieved, there is no limitation on the positions where the feed gas and the alkylating agent enter the reactor. In some embodiments, the feed gas enters the reactor from the upper part or top of the reactor, and the alkylating agent enters the reactor from the bottom or lower part of the reactor; in other embodiments, the feed gas enters the reactor from the bottom or lower part of the reactor, and the alkylating agent enters the reactor from the upper part or top of the reactor.

[0029] According to the present invention, in some preferred embodiments, the feed gas enters the reactor from the top of the reactor, and the alkylating agent enters the reactor from the bottom of the reactor.

[0030] According to the present invention, in some other preferred embodiments, the feed gas enters the reactor from the bottom of the reactor, and the alkylating agent enters the reactor from the top of the reactor.

[0031] According to the present invention, as long as the purpose of the present invention can be achieved, the type of the metal oxide is not limited, and any metal oxide capable of undergoing hydrogenation reaction is suitable for the system of the present invention, for example, a single metal oxide or a composite metal oxide. In some embodiments, the metal element in the metal oxide includes one or more of Cr, Mn, Zr, La, Ce, Al, Ti, Zn, In, and Ga. By adopting the above-mentioned embodiment, the conversion rate of CO can be increased, and transalkylation reactions such as alkylation and disproportionation of aromatic products can be triggered.

[0032] According to the present invention, the metal oxides can be obtained commercially. Those skilled in the art can also select precursors of metal elements as needed to prepare metal oxides of corresponding metal elements through co-precipitation, solid-phase mixing, sol-gel method and other methods. Therefore, the preparation of various metal oxides will not be elaborated in the present invention.

[0033] According to the present invention, as long as the purpose of the present invention can be achieved, the selection of the molecular sieve is not limited. In some embodiments, the molecular sieve includes one or more of ZSM-5 molecular sieve, ZSM-11 molecular sieve, Silicalite-1 molecular sieve and Silicalite-2 molecular sieve.

[0034] According to the present invention, in some preferred embodiments, the silicon to aluminum molar ratio (Si / Al) of the molecular sieve is greater than or equal to 10, for example, greater than or equal to 20, greater than or equal to 50, greater than or equal to 100, greater than or equal to 300 or greater than or equal to 400.

[0035] According to the present invention, in some more preferred embodiments, the silicon-aluminum molar ratio of the molecular sieve is greater than or equal to 50. The aforementioned embodiment can trigger transalkylation reactions such as alkylation and disproportionation of aromatic products, optimize the distribution of aromatic products in the CO hydrogenation reaction, and increase the selectivity of C9 aromatics.

[0036] In the present invention, it can be understood that the silicon-aluminum molar ratio can be infinite, that is, when the silicon-aluminum molar ratio of a certain molecular sieve is infinite, it means that the molecular sieve is an all-silicon molecular sieve.

[0037] According to the present invention, the molecular sieve of the present invention can be obtained from commercial sources or prepared from raw materials for preparing molecular sieves. For example, when preparing ZSM-5 molecular sieve, a person skilled in the art can prepare a mother liquor containing tetraethyl orthosilicate, aluminum isopropoxide, and tetrapropylammonium hydroxide according to the silicon-aluminum molar ratio required for the molecular sieve, and the resulting mother liquor is subjected to hydrothermal treatment, separation, and calcination to obtain the ZSM-5 molecular sieve. Therefore, the preparation of various molecular sieves is not described in detail in the present invention.

[0038] According to the present invention, those skilled in the art will appreciate that the feed gas includes H2 and CO. As long as the objectives of the present invention are achieved, the amounts of H2 and CO are not limited. In some embodiments, the molar ratio of H2 to CO is 0.5-6.0; in some preferred embodiments, the molar ratio of H2 to CO is 0.5-3.0 (e.g., 0.5:1, 1:1, 1.5:1, 2:1, 3:1, and any combination thereof). The aforementioned embodiments can improve CO selectivity and optimize the distribution of aromatic hydrocarbon products in the CO hydrogenation reaction.

[0039] According to the present invention, as long as the purpose of the present invention can be achieved, the selection of the alkylating agent is not limited. In some embodiments, the alkylating agent is selected from one or more of methanol, dimethyl ether, halogenated alkanes, and low-carbon olefins.

[0040] According to the present invention, in some preferred embodiments, the number of carbon atoms in the halogenated alkane is 1-4, for example, monochloromethane, monochloroethane or dichlorobutane; in some preferred embodiments, the number of carbon atoms in the low-carbon olefin is 2-4, for example, ethylene, propylene or isobutylene.

[0041] According to the present invention, the feed amounts of the alkylating agent and the feed gas are not limited as long as the objectives of the present invention can be achieved. In some embodiments, the feed gas is calculated as CO, and the alkylating agent is calculated as the provided alkyl group, and the molar ratio of the feed gas to the alkylating agent is (30-300):1 (e.g., 30:1, 50:1, 100:1, 150:1, 200:1, 300:1, and any combination of values ​​within the foregoing range). The aforementioned embodiments can effectively improve the CO conversion rate.

[0042] According to the present invention, those skilled in the art may, as needed, pretreat the catalyst before carrying out the reaction. In some embodiments, the pretreatment comprises pretreating with H2 at 350-420°C for 1-5 hours.

[0043] According to the present invention, as long as the purpose of the present invention can be achieved, the hydrogenation conditions are not limited. In some embodiments, the hydrogenation conditions include: the volume space velocity of the feed gas is 600-10000 mL g -1 h-1 (e.g. 600 mL g -1 h -1 、1000mL g -1 h -1 、1500mL g -1 h -1 , 2000mL g -1 h -1 、3000mL g -1 h -1 、4000mL g -1 h -1 、8000mL g -1 h -1 、10000mL g -1 h -1 , and any combination of numerical values ​​within the above range).

[0044] According to the present invention, in some embodiments, the hydrogenation conditions include: a reaction pressure of 1.0-8.0 MPa, preferably 3.0-6.0 MPa (e.g., 3.0 MPa, 4.0 MPa, 5 MPa, 6.0 MPa, and any combination of values ​​within the above range).

[0045] The second aspect of the present invention provides a system for producing aromatics by countercurrent hydrogenation of CO, the system comprising:

[0046] The reaction unit includes at least two catalyst beds loaded in a reactor, and is used for reacting the feed gas with the alkylating agent to produce a stream containing aromatic hydrocarbons;

[0047] The material inlet and outlet unit includes a raw gas inlet, an alkylating agent feed pipe, and a product outlet, which are used to introduce raw gas, alkylating agent, and lead out the logistics containing aromatics respectively;

[0048] The arrangement of the raw gas feed port and the alkylating agent feed pipe can achieve that the flow directions of the introduced raw gas and the introduced alkylating agent are opposite.

[0049] According to the system of the present invention, in some preferred embodiments, the alkylating agent feed pipe is provided with a one-way valve to prevent the raw gas and the alkylating agent from leaking through the alkylating agent feed pipe.

[0050] According to the present invention, the feed gas inlet is located at the upper part or top of the reactor, preferably at the top; the alkylating agent feed pipe inlet is located at the lower part or bottom of the reactor, preferably at the bottom.

[0051] According to the present invention, the product outlet is arranged at the bottom of the reactor.

[0052] According to the present invention, in some other preferred embodiments, the alkylating agent feed pipe is provided with a one-way valve to prevent the raw gas and the alkylating agent from leaking through the alkylating agent feed pipe.

[0053] According to the present invention, the feed gas inlet is located at the lower part or bottom of the reactor, preferably at the bottom; the alkylating agent feed pipe inlet is located at the upper part or top of the reactor, preferably at the top; and / or the product outlet is located at the bottom of the reactor.

[0054] The following combination Figure 1 Preferred embodiments of the method and system of the present invention are described.

[0055] like Figure 1 As shown, under hydrogenation conditions, the feed gas enters the reactor from the feed gas inlet 1 at the top of the reactor, and the alkylating agent enters the reactor from the alkylating agent feed pipe 4 provided with a one-way valve at the bottom of the reactor. The alkylating agent reacts in the loaded two-stage catalyst bed, the upper catalyst bed 2 and the lower catalyst bed 3 to produce a stream containing aromatic hydrocarbons, which flows out through the product outlet 5 at the bottom of the reactor.

[0056] The present invention will be described in detail below by way of examples. In the following examples:

[0057] The aromatics-containing stream includes unconverted CO and H2, and hydrocarbon products containing aromatics and / or C1-C 5+ Hydrocarbon composition. Aromatic hydrocarbons include C6-C 10+ Aromatics, C 5+ Hydrocarbons refer to aliphatic hydrocarbon compounds with a carbon number of 5 or more. The composition of the feed gas and products was obtained by online gas chromatography analysis. A hydrogen flame detector was used to detect hydrocarbon products, and the amount of each hydrocarbon product was calculated based on its response factor.

[0058] The selectivity of each product is defined as the percentage (mol%) of each product in the total carbon number of the organic product. The specific calculation method is as follows:

[0059] Total carbon number of organic product = Σ(amount of substance of organic product i × number of carbon atoms in organic product i molecule);

[0060] Selectivity of organic product j = amount of substance of organic product j × number of carbon atoms in organic product j molecule / total carbon number of organic products × 100%

[0061] Aromatic selectivity = C6 aromatic selectivity + C7 aromatic selectivity + C8 aromatic selectivity + C9 aromatic selectivity + C 10+ Aromatic selectivity;

[0062] C n Aromatic selectivity (denoted as A n )=Cn Aromatic selectivity / aromatic selectivity×100% (n=6, 7, 8, 9, 10...).

[0063] A 6-8 :A9:A 10 C6-C8 aromatics: C9 aromatics: C 10+ The molar ratio of aromatic hydrocarbons.

[0064] Example 1

[0065] like Figure 1 As shown, a metal oxide (Cr2O3) and a ZSM-5 molecular sieve (Si / Al ratio = 50) are mixed in a mass ratio of 1:1 by grinding, mixing, and then pressing into tablets, which are then loaded into the upper layer of the reactor to obtain a catalyst upper bed 2; a ZSM-5 molecular sieve (Si / Al molar ratio = 50) and a silica sol binder are uniformly mixed and extruded into tablets, which are then loaded into the lower layer of the reactor to obtain a catalyst lower bed 3 (the silica sol binder content in the formed catalyst is 20 wt %), and the volume ratio of the catalyst upper bed 2 to the catalyst lower bed 3 is 5:1; the catalyst is pretreated with H2 at 380°C for 2 h;

[0066] Under hydrogenation conditions, a feed gas (a mixture of H2 and CO, with a molar ratio of H2 to CO of 1:1) enters the reactor from a feed gas inlet 1 at the top of the reactor, and an alkylating agent (dimethyl ether) enters the reactor from an alkylating agent feed pipe 4 equipped with a one-way valve at the bottom of the reactor (i.e., the feed gas and the alkylating agent enter the reactor in a countercurrent manner). The feed gas and the alkylating agent react in the loaded two-stage catalyst bed, i.e., the upper catalyst bed 2 and the lower catalyst bed 3, to produce an aromatic hydrocarbon-containing stream, which then flows out through a product outlet 5 at the bottom of the reactor.

[0067] The hydrogenation conditions include: feed gas volume space velocity 2000mLg -1 h -1 , the reaction pressure is 4.0MPa;

[0068] The temperatures of the catalyst upper bed 2 and the catalyst lower bed 3 are 400° C. and 500° C. respectively; the feed gas is calculated as CO, the alkylating agent is calculated as the provided alkyl group, and the molar ratio of the feed gas to the alkylating agent is 300:1.

[0069] The results of CO conversion, aromatics selectivity, and aromatics distribution are shown in Table 1.

[0070] Example 2

[0071] like Figure 1As shown, a metal oxide (Cr2O3) and a ZSM-5 molecular sieve (Si / Al molar ratio = 50) are mixed by grinding and mixing in a mass ratio of 1:1, and then a silica sol binder is added and extruded to form a mixture (the silica sol binder content in the formed catalyst is 20 wt %), and loaded into the upper layer of the reactor to obtain a catalyst upper bed 2; a ZSM-5 molecular sieve (Si / Al ratio = 50) is separately formed by tableting and loaded into the lower layer of the reactor to obtain a catalyst lower bed 3, and the volume ratio of the catalyst upper bed 2 to the catalyst lower bed 3 is 5:1; the catalyst is pretreated with H2 at a temperature of 380°C for 2 h;

[0072] Under hydrogenation conditions, a feed gas (a mixture of H2 and CO, with a molar ratio of H2 to CO of 1:1) enters the reactor from a feed gas inlet 1 at the top of the reactor, and an alkylating agent (methanol) enters the reactor from an alkylating agent feed pipe 4 equipped with a one-way valve at the bottom of the reactor (i.e., the feed gas and the alkylating agent enter the reactor in a countercurrent manner). The feed gas and the alkylating agent react in the loaded two-stage catalyst bed, i.e., the upper catalyst bed 2 and the lower catalyst bed 3, to produce an aromatic hydrocarbon-containing stream, which then flows out through a product outlet 5 at the bottom of the reactor.

[0073] The hydrogenation conditions include: feed gas volume space velocity 2000mLg -1 h -1 , the reaction pressure is 4.0MPa;

[0074] The temperatures of the catalyst upper bed 2 and the catalyst lower bed 3 are 400° C. and 500° C. respectively; the feed gas is calculated as CO, the alkylating agent is calculated as the provided alkyl group, and the molar ratio of the feed gas to the alkylating agent is 300:1.

[0075] The results of CO conversion, aromatics selectivity, and aromatics distribution are shown in Table 1.

[0076] Example 3

[0077] A metal oxide (ZrO2) and a ZSM-5 molecular sieve (Si / Al ratio = 100) were mixed and molded in a weight ratio of 3:1 and loaded into the upper layer of a reactor to obtain a catalyst upper bed; ZrO2 and a ZSM-5 molecular sieve (Si / Al ratio = 200) were mixed and molded in a weight ratio of 1:1 and loaded into the middle layer of a reactor to obtain a catalyst middle bed; ZSM-11 molecular sieve (Si / Al ratio = 400) was separately molded and loaded into the lower layer of the reactor to obtain a catalyst lower bed; the volume ratio of the catalyst upper bed, the catalyst middle bed, and the catalyst lower bed was 3:2:1; the catalyst was pretreated with H2 at 380°C for 2 hours;

[0078] Under hydrogenation conditions, a feed gas (a mixture of H2 and CO with a molar ratio of H2 to CO of 2:1) enters the reactor from a feed gas inlet at the top of the reactor, and an alkylating agent (ethylene) enters the reactor from an alkylating agent feed pipe equipped with a one-way valve at the bottom of the reactor (i.e., the feed gas and the alkylating agent enter the reactor in a countercurrent manner). The alkylating agent reacts in the three-stage catalyst bed, i.e., the upper catalyst bed, the middle catalyst bed, and the lower catalyst bed, to produce an aromatics-containing stream, which then flows out through a product outlet at the bottom of the reactor.

[0079] The hydrogenation conditions include: feed gas volume space velocity 1500mLg -1 h -1 , the reaction pressure is 5.0MPa;

[0080] The temperatures of the catalyst upper bed, catalyst middle bed and catalyst lower bed are 300° C., 400° C. and 520° C. respectively; the feed gas is calculated as CO, the alkylating agent is calculated as the provided alkyl group, and the molar ratio of the feed gas to the alkylating agent is 30:1.

[0081] The results of CO conversion, aromatics selectivity, and aromatics distribution are shown in Table 1.

[0082] Example 4

[0083] A metal oxide (La2O3) and a ZSM-5 molecular sieve (Si / Al ratio = 200) were mixed and molded in a weight ratio of 5:1 and loaded into the upper layer of a reactor to obtain a catalyst upper bed; La2O3 and a ZSM-5 molecular sieve (Si / Al ratio = 200) were mixed and molded in a weight ratio of 3:1 and loaded into the middle layer of a reactor to obtain a catalyst middle bed; ZSM-11 molecular sieve (Si / Al ratio = 300) was separately molded and loaded into the lower layer of the reactor to obtain a catalyst lower bed; the volume ratio of the catalyst upper bed, the catalyst middle bed, and the catalyst lower bed was 5:4:1; the catalyst was pretreated with H2 at 400°C for 2 hours;

[0084] Under hydrogenation conditions, a feed gas (a mixture of H2 and CO with a molar ratio of H2 to CO of 3:1) enters the reactor from a feed gas inlet at the top of the reactor, and an alkylating agent (isobutylene) enters the reactor from an alkylating agent feed pipe equipped with a one-way valve at the bottom of the reactor (i.e., the feed gas and the alkylating agent enter the reactor in a countercurrent manner). The reaction proceeds in the three-stage catalyst bed, i.e., the upper catalyst bed, the middle catalyst bed, and the lower catalyst bed, to produce an aromatics-containing stream, which then flows out through a product outlet at the bottom of the reactor.

[0085] The hydrogenation conditions include: feed gas volume space velocity 1000mLg -1 h -1 , the reaction pressure is 6.0MPa;

[0086] The temperatures of the catalyst upper bed, catalyst middle bed and catalyst lower bed are 350° C., 400° C. and 550° C. respectively; the feed gas is calculated as CO, the alkylating agent is calculated as the provided alkyl group, and the molar ratio of the feed gas to the alkylating agent is 180:1.

[0087] The results of CO conversion, aromatics selectivity, and aromatics distribution are shown in Table 1.

[0088] Example 5

[0089] The metal oxide (ZnCr3O x , Zn / Cr molar ratio 1:3) and ZSM-5 molecular sieve (Si / Al ratio = 100) were mixed in a weight ratio of 2:1 and loaded into the upper layer of the reactor to obtain the catalyst upper bed; Cr2O3 and ZSM-5 molecular sieve (Si / Al ratio = 100) were mixed in a weight ratio of 1:1 and loaded into the upper middle layer of the reactor to obtain the catalyst upper bed; ZSM-11 molecular sieve (Si / Al ratio = 300) was separately molded and loaded into the lower middle layer of the reactor to obtain the catalyst lower bed; Silicalite-1 molecular sieve (all-silicon molecular sieve) was separately molded and loaded into the lower layer of the reactor to obtain the catalyst lower bed. The volume ratio of the catalyst upper bed, the catalyst upper middle bed, the catalyst lower middle bed, and the catalyst lower bed was 5:5:1:1; the catalyst was pretreated with H2 at 420°C for 2 hours;

[0090] Under hydrogenation conditions, a feed gas (a mixture of H2 and CO, with a molar ratio of H2 to CO of 1:1) enters the reactor from a feed gas inlet at the top of the reactor, and an alkylating agent (methylene chloride) enters the reactor from an alkylating agent feed pipe equipped with a one-way valve at the bottom of the reactor (i.e., the feed gas and the alkylating agent enter the reactor in a countercurrent manner). The catalyst bed layers, i.e., the upper catalyst bed layer, the middle upper catalyst bed layer, the middle lower catalyst bed layer, and the lower catalyst bed layer, are contacted and reacted to produce an aromatic hydrocarbon-containing stream, which then flows out through a product outlet at the bottom of the reactor.

[0091] The hydrogenation conditions include: feed gas volume space velocity 6000mLg -1 h -1 , the reaction pressure is 8.0MPa;

[0092] The temperatures of the catalyst upper bed, catalyst middle bed and catalyst lower bed are 300°C, 350°C, 400°C and 500°C respectively; the feed gas is calculated as CO, the alkylating agent is calculated as the provided alkyl group, and the molar ratio of the feed gas to the alkylating agent is 100:1.

[0093] The results of CO conversion, aromatics selectivity, and aromatics distribution are shown in Table 1.

[0094] Example 6

[0095] The method of Example 2 is different in that:

[0096] The temperatures of the catalyst upper bed 2 and the catalyst lower bed 3 are 300° C. and 500° C., respectively.

[0097] The results of CO conversion, aromatics selectivity, and aromatics distribution are shown in Table 1.

[0098] Example 7

[0099] The method of Example 2 is different in that:

[0100] Metal oxide (Cr2O3) and ZSM-5 (Si / Al ratio = 50) were mixed and molded in a mass ratio of 0.8:1 and loaded into the lower layer of the reactor to obtain a catalyst lower bed 3. The volume ratio of the catalyst upper bed 2 to the catalyst lower bed 3 was 0.5:1.

[0101] The results of CO conversion, aromatics selectivity, and aromatics distribution are shown in Table 1.

[0102] Comparative Example 1

[0103] Metal oxide (Cr2O3) and ZSM-5 (Si / Al ratio = 50) were mixed in a mass ratio of 1:1 and loaded into a reactor to obtain a catalyst bed. The catalyst was pretreated with H2 at 380°C for 2 hours.

[0104] Under hydrogenation conditions, the raw gas (a mixture of H2 and CO, with a molar ratio of H2 to CO of 1:1) enters the reactor from the raw gas inlet at the top of the reactor to react to obtain a stream containing aromatic hydrocarbons, which then flows out through the product outlet at the bottom of the reactor;

[0105] The hydrogenation conditions include a feed gas space velocity of 2000 mL / min. -1 h -1 ; The reaction pressure is 4.0MPa; The reaction temperature is 400℃.

[0106] The results of CO conversion, aromatics selectivity, and aromatics distribution are shown in Table 1.

[0107] Comparative Example 2

[0108] ZSM-5 (Si / Al ratio = 50) is separately formed and loaded into a reactor to form a catalyst bed. A raw material (xylene, trimethylbenzene, and tetramethylbenzene mixed in a molar ratio of 3:5:2 to obtain an aromatic hydrocarbon raw material) enters the reactor through a raw gas inlet at the top of the reactor, and an alkylating agent (methanol) enters the reactor through an alkylating agent feed pipe equipped with a one-way valve at the bottom of the reactor. The alkylating agent reacts in the loaded catalyst bed to obtain an aromatic hydrocarbon-containing stream, which then flows out through a product outlet at the bottom of the reactor.

[0109] Among them, the space velocity of aromatic hydrocarbon raw material is 0.3 mol g -1 h -1 (in terms of carbon atoms), the reaction pressure is 4.0 MPa; the temperature of the catalyst bed is 400° C.; and the molar ratio of the aromatic hydrocarbon feedstock to the alkylating agent is 300:1.

[0110] The aromatic selectivity and aromatic distribution results are shown in Table 1.

[0111] Comparative Example 3

[0112] ZSM-5 (Si / Al ratio = 50) is separately formed and loaded into a reactor to form a catalyst bed; hydrogen and a raw material (xylene, trimethylbenzene, and tetramethylbenzene mixed in a molar ratio of 3:5:2 to obtain an aromatic hydrocarbon raw material) are introduced into the reactor through a raw gas inlet at the top of the reactor; an alkylating agent (methanol) is introduced into the reactor through an alkylating agent feed pipe equipped with a one-way valve at the bottom of the reactor to react with each other in the loaded catalyst bed to obtain an aromatic hydrocarbon-containing stream, which then flows out through a product outlet at the bottom of the reactor;

[0113] Among them, the hydrogen space velocity is 1000mLg -1 h -1 , aromatic space velocity 0.3 mol g -1 h -1 (in terms of carbon atoms), the reaction pressure is 4.0 MPa; the temperature of the catalyst bed is 400° C.; and the molar ratio of the aromatic hydrocarbon feedstock to the alkylating agent is 300:1.

[0114] The aromatic selectivity and aromatic distribution results are shown in Table 1.

[0115] Comparative Example 4

[0116] The method of Example 2 is as follows, except that

[0117] A metal oxide (Cr2O3) and ZSM-5 (Si / Al ratio = 50) are mixed in a mass ratio of 1:1, formed, and loaded into a reactor to form a catalyst bed (i.e., only a first-stage catalyst bed); the feed gas and the alkylating agent contact and react in the catalyst bed to produce an aromatic hydrocarbon-containing stream, which flows out through a product outlet at the bottom of the reactor;

[0118] The temperature of the catalyst bed is 400°C.

[0119] The results of CO conversion, aromatics selectivity, and aromatics distribution are shown in Table 1.

[0120] Comparative Example 5

[0121] The method of Example 2 is as follows, except that

[0122] Metal oxide (Cr2O3) and ZSM-5 (Si / Al ratio = 50) are mixed and molded in a mass ratio of 1:1, and loaded into the lower layer of the reactor to obtain a catalyst lower bed layer 3; ZSM-5 (Si / Al ratio = 50) is separately molded and loaded into the upper layer of the reactor to obtain a catalyst upper bed layer 2. The volume ratio of the catalyst upper bed layer 2 to the catalyst lower bed layer 3 is 1:5.

[0123] The results of CO conversion, aromatics selectivity, and aromatics distribution are shown in Table 1.

[0124] Comparative Example 6

[0125] The method of Example 2 is different in that:

[0126] The feed gas enters the reactor from a feed gas inlet 1 at the bottom of the reactor, and the alkylating agent enters the reactor from an alkylating agent feed pipe 4 at the bottom of the reactor equipped with a one-way valve (i.e., the feed gas and the alkylating agent enter the reactor in parallel). The feed gas and the alkylating agent are contacted in the loaded two-stage catalyst bed, i.e., the upper catalyst bed 2 and the lower catalyst bed 3, to undergo a hydrogenation reaction to produce an aromatic hydrocarbon-containing stream, which then flows out through a product outlet 5 at the bottom of the reactor.

[0127] The results of CO conversion, aromatics selectivity, and aromatics distribution are shown in Table 1.

[0128] Table 1

[0129]

[0130] The results in Table 1 show that the aromatics-containing streams prepared in Examples 1-7 of the present invention, in which the feed gas and the alkylating agent are introduced into the reactor in a countercurrent manner and the weight ratio of the metal oxide to the molecular sieve is gradually decreased, have a high CO conversion rate and a high selectivity for aromatics, especially C9 aromatics. Examples 1-7 of the present invention and Comparative Examples 2-3 illustrate that, compared with the reaction of aromatics with an alkylating agent, the embodiment of the present invention, in which the feed gas and the alkylating agent are introduced into the reactor in a countercurrent manner and the weight ratio of the metal oxide to the molecular sieve is gradually decreased, has a better effect on regulating the distribution of aromatics.

Claims

1. A method for preparing aromatics by countercurrent hydrogenation of CO, characterized in that: The method comprises: under hydrogenation conditions, feed gas and alkylating agent enter a reactor in a countercurrent manner, contact and react in at least two catalyst beds filled with the feed gas to obtain a stream containing aromatic hydrocarbons; the catalyst in the catalyst bed contains metal oxide and molecular sieve, and the weight ratio of the metal oxide to the molecular sieve decreases step by step along the feed gas flow direction; The temperature of the catalyst bed at each stage is 300-580°C; Along the direction of the raw material gas flow, the temperature of the catalyst bed at each level increases step by step; The raw gas includes H2 and CO, and the molar ratio of H2 to CO is (0.5-6.0):1; The raw gas is calculated based on CO, and the alkylating agent is calculated based on the provided alkyl group, and the molar ratio of the raw gas to the alkylating agent is (30-300):1; The hydrogenation conditions include: The volume space velocity of the raw gas is 600-10000 mL g -1 h -1 ; The reaction pressure is 1.0-8.0 MPa.

2. The method according to claim 1, wherein The temperature difference between the two adjacent catalyst beds is 20-200°C.

3. The method according to claim 2, wherein: The temperature difference between the two adjacent catalyst beds is 50-150°C.

4. The method according to claim 1, wherein The difference in weight ratio of metal oxide to molecular sieve between two adjacent catalyst beds is 0.2-5.0; The catalyst bed is of level 2-5; and / or Along the direction of raw material gas flow, the volume ratio of the catalyst beds of two adjacent stages is (0.5-10.0):

1.

5. The method according to claim 4, wherein Along the flow direction of the raw material gas, the weight ratio of metal oxide to molecular sieve in the first stage catalyst bed is (0.2-5.0):1, and the metal oxide content in the last stage catalyst bed is 0; and / or Along the direction of raw material gas flow, the volume ratio of the catalyst beds of two adjacent stages is (1.0-5.0):

1.

6. The method according to claim 1, wherein The raw gas enters the reactor from the upper part or top of the reactor; The alkylating agent enters the reactor from the bottom or lower part of the reactor; or, The raw gas enters the reactor from the bottom or lower part of the reactor; The alkylating agent enters the reactor from the upper part or top of the reactor.

7. The method according to claim 6, wherein: The raw gas enters the reactor from the top of the reactor; The alkylating agent enters the reactor from the bottom of the reactor; or, The raw gas enters the reactor from the bottom of the reactor; The alkylating agent enters the reactor from the top of the reactor.

8. The method according to claim 1, wherein The metal elements in the metal oxide include one or more of Cr, Mn, Zr, La, Ce, Al, Ti, Zn, In and Ga; and / or The molecular sieve comprises one or more of ZSM-5 molecular sieve, ZSM-11 molecular sieve, Silicalite-1 molecular sieve and Silicalite-2 molecular sieve; and / or The silicon-aluminum molar ratio of the molecular sieve is greater than or equal to 10.

9. The method according to claim 8, wherein The silicon-aluminum molar ratio of the molecular sieve is greater than or equal to 50.

10. The method according to claim 1, wherein The molar ratio of H2 to CO is (0.5-3.0):1; and / or The alkylating agent is selected from one or more of methanol, dimethyl ether, halogenated alkanes and low-carbon olefins.

11. The method according to claim 1, wherein The method further comprises pretreating the catalyst before carrying out the reaction; and / or The reaction pressure is 3.0-6.0MPa.

12. The method according to claim 11, wherein The pretreatment includes pretreatment with H2 at 350-420°C for 1-5 hours.

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