Continuous process for preparing aromatic hydrocarbon through supercritical carbon dioxide hydrogenation

Abstract

The invention provides a continuous process for preparing aromatic hydrocarbon by supercritical carbon dioxide hydrogenation, which comprises the following steps: 1) adding a reaction liquid containing a solvent and a catalyst into a reaction kettle, spraying supercritical carbon dioxide and hydrogen into a reactor by adopting a Venturi ejector, and contacting with the catalyst in the reaction kettle to react; (2) after the reaction in the step (1), the gas in the reactor is sequentially subjected to primary condensation and secondary condensation to respectively condense the solvent and the water, the solvent of the primary condensation returns to the reaction kettle, the water of the secondary condensation is discharged, and the gas after the secondary condensation sequentially passes through an absorption tower and a desorption tower to capture non-aromatic hydrocarbon, the residual gas is used as recycle gas and returns to the Venturi ejector to enter the reaction kettle; and (3) pressurizing the reaction liquid discharged from the reaction kettle after the reaction in the step (1) through a circulating pump, feeding the reaction liquid into a heat exchanger to remove reaction heat, dividing the reaction liquid subjected to heat exchange into two streams through a cross-flow filter, feeding one stream of penetrating fluid into a rectification device, and feeding the other stream of concentrated liquid back to the reaction kettle for continuous reaction.

Description

Technical Field

[0001] This invention relates to the field of aromatic hydrocarbon preparation technology, and more specifically to a continuous process for preparing aromatic hydrocarbons by supercritical carbon dioxide hydrogenation. Background Technology

[0002] With the rapid development of modern industry, fossil fuels such as coal, oil, and natural gas have been extensively mined and used, leading to a continuous increase in carbon dioxide (CO2) emissions. Since the Industrial Revolution, global CO2 emissions have been steadily increasing, and its concentration in the atmosphere has also been rising, making the greenhouse effect increasingly prominent and bringing about many serious environmental problems such as ocean acidification and climate change, which has become a focus of global attention.

[0003] Aromatics, as one of the most important basic chemical raw materials, plays a crucial role in the production of fragrances, pharmaceuticals, dyes, and high molecular weight polymers, accounting for nearly one-third of the petrochemical product market. However, the production of most aromatics still heavily relies on petroleum, while petroleum resources are facing increasing depletion. Meanwhile, although my country has abundant coal resources, its petroleum resources are relatively scarce, posing a serious energy security challenge. Therefore, developing new avenues for aromatics production is extremely urgent.

[0004] CO2 is not only a major greenhouse gas, but also the most abundant renewable carbon resource. From a resource perspective, it is the world's cheapest C1 resource. Therefore, converting CO2 into valuable chemicals, especially aromatics, can not only solve the problem of insufficient chemical supply, but also achieve large-scale CO2 emission reduction, which has significant economic and social implications.

[0005] Currently, there are two routes for producing aromatics from CO2: one is that CO2 first generates CO through a reverse water-gas shift reaction, then undergoes Fischer-Tropsch synthesis to generate low-carbon olefins, and then the low-carbon olefins undergo aromatization on a molecular sieve to generate aromatics; the other is that CO2 first forms a methanol intermediate on a bifunctional catalyst, and then methanol-to-aromatics (MTA) occurs at the acid sites of the molecular sieve.

[0006] CN107840778B reports a method for producing aromatics from carbon dioxide via hydrogenation using a composite catalyst. This composite catalyst consists of an iron-based carbon dioxide hydrogenation catalyst for producing low-carbon olefins as the first component, mixed with a metal-modified or unmodified molecular sieve as the second component. CN110496639B discloses a catalyst for the hydrogenation of CO2 to aromatics, its preparation method, and its application. The catalyst is composed of zinc-aluminum spinel oxide and an acidic molecular sieve. CN110743606B provides a method for preparing a catalyst for the direct hydrogenation of carbon dioxide to aromatics and a method for synthesizing aromatics. The catalyst is composed of a physical mixture of a metal oxide catalyst M and a molecular sieve catalyst Z. CN114160183B provides a method for preparing a sheet-like HZSM-5 molecular sieve catalyst and its application in the hydrogenation of carbon dioxide to aromatics. CN114950534B reports a bifunctional catalyst... A process for producing aromatics via methanol intermediates through carbon dioxide hydrogenation using a bifunctional catalyst is described. This bifunctional catalyst uses alumina as a support, transition metals as active components, and alkali metals as co-active components. The molecular sieve catalyst is a composite of SAPO-n and HZSM-5 molecular sieves. CN117654402A discloses a process and separation system for efficiently separating feed gas from light hydrocarbons in carbon dioxide-to-aromatics production. CN119591464A provides a process for synthesizing aromatic chemicals, which improves CO2 conversion and aromatic selectivity by incorporating a certain amount of CO into the reactant gas and selecting a specific composite catalyst of metal oxide and molecular sieve. Current research mainly focuses on bifunctional catalyst systems for CO2 hydrogenation to aromatics, with few reports on carbon dioxide-to-aromatics processes. Research on improving reaction efficiency and aromatic selectivity from a process perspective is extremely scarce.

[0007] Therefore, in view of the above situation, providing a novel and efficient method for producing aromatics by carbon dioxide hydrogenation has become a technical problem that urgently needs to be solved in this field. Summary of the Invention

[0008] In view of the shortcomings of the prior art described above, the problem to be solved by the present invention is to provide a continuous process for the preparation of aromatics by supercritical carbon dioxide hydrogenation, so as to avoid the problems of hot spots and runaway temperature, internal diffusion of catalyst, easy blockage of catalyst, and slow diffusion of gaseous reactants in traditional gas-solid phase reaction processes, while improving reaction efficiency, CO2 conversion rate and aromatic selectivity.

[0009] To address the above problems, the present invention provides the following technical solution:

[0010] A continuous process for producing aromatics by supercritical carbon dioxide hydrogenation includes the following steps:

[0011] 1) Add a reaction liquid containing solvent and catalyst to the reactor. Use a Venturi injector to simultaneously inject supercritical carbon dioxide and hydrogen into the reactor, so that hydrogen dissolves in the supercritical carbon dioxide fluid and disperses into the reaction liquid. Hydrogen and supercritical carbon dioxide react with the catalyst in the reactor.

[0012] 2) After the reaction in step 1), the gas in the reactor passes through a primary condenser and a secondary condenser to condense the solvent and water respectively. The solvent from the primary condenser is returned to the reactor, and the water from the secondary condenser is discharged as wastewater. The gas after the secondary condenser passes through an absorption tower and a desorption tower to capture non-aromatic hydrocarbons. The remaining gas is returned as circulating gas to the Venturi ejector and then enters the reactor.

[0013] 3) After the reaction in step 1), the reaction liquid discharged from the reactor is pressurized by the circulating pump and enters the heat exchanger to remove the heat of reaction. The removed heat of reaction is absorbed by the hot oil on the cold side and can be used to produce medium-pressure steam. After heat exchange, the reaction liquid is divided into two streams through the cross-flow filter. One stream is a permeate rich in aromatics, which is sent to the distillation unit. The other stream is a concentrated liquid containing inert solvent and catalyst particles, which enters the Venturi ejector and returns to the reactor for continuous reaction.

[0014] Compared with the prior art, the beneficial effects of the present invention are as follows:

[0015] 1. This invention uses a fine-particle catalyst suspended in a liquid-phase heat carrier solvent for reaction. The fine-particle catalyst can eliminate the influence of internal diffusion, increase the single-pass conversion rate and aromatic yield. At the same time, the solvent as a heat carrier can significantly improve the heat transfer capacity of the system, avoid the hot spots and runaway temperature problems of traditional gas-solid reactions, and the solvent has good solubility and selectivity for aromatics, which facilitates the separation of aromatics in subsequent distillation units. In addition, since bifunctional catalysts are usually used in this field, using a solvent as a carrier can make the two components of the catalyst mix more evenly, improve the reaction efficiency, and avoid the complex catalyst mixing and forming process.

[0016] 2. This invention selects supercritical carbon dioxide as a reactant for hydrogenation to prepare aromatics, which can eliminate the mass transfer resistance caused by the solubility of hydrogen, accelerate the reaction rate by dissolving a large amount of hydrogen, solve the problem of slow diffusion of gaseous reactants, reduce the solvation effect, significantly increase catalyst activity, and improve reactant conversion rate and selectivity.

[0017] 3. This invention uses a jet-circulating reactor, including a reaction vessel and a Venturi ejector, as the key process equipment for the continuous preparation of aromatics by carbon dioxide hydrogenation. The reaction fluid circulates at a high flow rate in the loop and is injected at high speed through the Venturi ejector. A negative pressure is formed at the working nozzle, which causes the process gas to be drawn into the ejector to form micron-sized bubbles with a large specific surface area. This increases the gas-liquid-solid contact, accelerates the gas-liquid-solid multiphase reaction rate, and improves the reaction efficiency, thereby increasing the space velocity while ensuring the yield.

[0018] 4. The present invention uses a circulating gas separation device to remove water and non-aromatic hydrocarbons from the system in the circulating gas, promotes the reaction equilibrium to shift in a direction that is favorable to product formation, improves reactant conversion rate and aromatic hydrocarbon selectivity, and can also recover solvent and reduce solvent loss. Attached Figure Description

[0019] Figure 1 This is a schematic diagram of a CO2 hydrogenation to aromatics reaction system according to a preferred embodiment of the present invention;

[0020] Figure 2 This is a schematic diagram of the Venturi injector.

[0021] Figure 3 This is a schematic diagram of a traditional CO2 hydrogenation reaction system for producing aromatics.

[0022] Figure 1 middle,

[0023] 1-Reaction vessel; 101-Gas outlet; 102-Condensate reflux inlet; 103-Reaction liquid outlet;

[0024] 2-Circulation pump;

[0025] 3-Reaction heat exchanger; 301-Process side, 302-Hot oil side;

[0026] 4-Cross-flow filter; 401-Permeate outlet, 402-Concentrate outlet;

[0027] 5 - First stage condenser; 501 - Air inlet; 502 - First condensate outlet;

[0028] 6-Secondary condenser; 601-Second condensate outlet;

[0029] 7-Absorber; 701-Absorber gas phase inlet, 702-Absorber gas phase outlet, 703-Absorbent liquid inlet, 704-Absorber liquid phase outlet;

[0030] 8-Desorption tower; 801-Liquid phase inlet of desorption tower, 802-Hydrogen inlet, 803-Gas phase outlet of desorption tower, 804-Liquid phase outlet of desorption tower;

[0031] 9- Circulating gas heater;

[0032] 10-Venturi injector; 1001-Recirculating gas inlet;

[0033] 11-Distillation unit; 1101-Liquid phase inlet for distillation, 1102-Permeate inlet, 1103-Absorbate outlet, 1104-Aromatic product outlet;

[0034] 12-Steam drum; 1201-Hot oil outlet; 1202-Medium pressure steam outlet;

[0035] 13-Supercritical CO2 inlet.

[0036] Figure 2 In the middle section, 1002 is the inlet section, 1003 is the gas chamber, 1004 is the nozzle, 1005 is the mixing section, and 1006 is the diffusion section.

[0037] Figure 3 In the middle, 14-tube fixed bed reactor, 15-preheater, 16-post-condenser, 17-three-phase separator. Detailed Implementation

[0038] The endpoints and any values ​​of the ranges 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 endpoint values ​​of the various ranges, the endpoint values ​​of the various ranges and individual point values, and individual point values ​​can be combined with each other to obtain one or more new numerical ranges, which should be considered as specifically disclosed herein.

[0039] To make the present invention more apparent and understandable, preferred embodiments are described in detail below with reference to the accompanying drawings.

[0040] Continuous process for producing aromatics by supercritical carbon dioxide hydrogenation

[0041] This invention provides a continuous process for the preparation of aromatics by supercritical carbon dioxide hydrogenation, comprising the following steps:

[0042] 1) Add a reaction liquid containing solvent and catalyst to the reactor. Use a Venturi injector to simultaneously inject supercritical carbon dioxide and hydrogen into the reactor, so that the hydrogen dissolves in the supercritical carbon dioxide fluid and disperses into the reaction liquid. The hydrogen and supercritical carbon dioxide fluid react with the catalyst in the reactor.

[0043] 2) After the reaction in step 1), the gas in the reactor passes through a primary condenser and a secondary condenser to condense the solvent and water respectively. The solvent from the primary condenser is returned to the reactor, and the water from the secondary condenser is discharged as wastewater. The gas after the secondary condenser passes through an absorption tower and a desorption tower to capture non-aromatic hydrocarbons. The remaining gas is returned as circulating gas to the Venturi ejector and then enters the reactor.

[0044] 3) After the reaction in step 1), the reaction liquid discharged from the reactor is pressurized by the circulating pump and enters the heat exchanger to remove the heat of reaction. The removed heat of reaction is absorbed by the hot oil on the cold side and can be used to produce medium-pressure steam. After heat exchange, the reaction liquid is divided into two streams through the cross-flow filter. One stream is a permeate rich in aromatics, which is sent to the distillation unit. The other stream is a concentrated liquid containing inert solvent and catalyst particles, which enters the Venturi ejector and returns to the reactor for continuous reaction.

[0045] In the continuous process for producing aromatics by supercritical carbon dioxide hydrogenation provided by this invention, step 1) involves adding a reaction liquid containing solvent and catalyst to the reactor, and simultaneously injecting supercritical carbon dioxide and hydrogen into the reactor using a Venturi injector. This allows the hydrogen to dissolve in the supercritical carbon dioxide fluid and be fully dispersed in the reaction liquid, ensuring that the hydrogen and supercritical carbon dioxide fully contact and react with the catalyst in the reactor. Specifically:

[0046] Step 1) of this invention describes a reaction already in a steady-state state. Typically, initially, a reaction liquid containing solvent and catalyst is added to the reactor. After sealing, CO2 and H2 (hydrogen and carbon monoxide volume ratio 2.0–4.0:1) are introduced into the reactor until the system pressure reaches 0.8–2.5 MPaG. The circulation pump is then turned on, and the air inside the reactor is replaced multiple times. After the last replacement, the reaction system is maintained at a certain pressure, and the temperature is raised to the reaction temperature of 320–360°C via a reaction heat exchanger. Immediately afterwards, CO2 or a mixture of supercritical carbon dioxide and hydrogen (hydrogen and carbon dioxide volume ratio 3:1) is added until the reactor pressure reaches 7.5–15.0 MPaG. This point is recorded as the reaction start time. After reaching steady state, the reaction process is kept constant at the above pressure, and supercritical carbon dioxide and hydrogen are simultaneously injected into the reactor using a Venturi injector.

[0047] In step 1) of the present invention, the solvent has strong selectivity and large dissolving ability for aromatic hydrocarbons, and good thermal stability. It is preferably one or more of macrocyclobutane sulfone, N-methylpyrrolidone, N-formylmorpholine, triethylene glycol, and tetraethylene glycol, and more preferably tetraethylene glycol.

[0048] In step 1) of this invention, the catalyst is a bifunctional catalyst composed of metal oxide and molecular sieve.

[0049] Optionally, the metal element in the metal oxide is selected from one or more of Cr, La, Zr, Fe, Ce, Al, Ti, Zn, In and Ga, and an additive may or may not be added to the metal oxide; the additive is selected from one or more of K, Na, Cu, Mn and V.

[0050] Optionally, the molecular sieve is selected from one or more of H-ZSM-5 molecular sieve, ZSM-11 molecular sieve, Silicalite-1 molecular sieve, and Silicalite-2 molecular sieve. The silica-alumina molar ratio of the molecular sieve is greater than or equal to 20, preferably greater than or equal to 40.

[0051] Optionally, the mass ratio of metal oxide to molecular sieve in the bifunctional catalyst is 1:3 to 3:1; further, it can be 1:2 to 2:1 or 2:3 to 3:2.

[0052] Optionally, the metal oxide and molecular sieve are mixed as follows: the metal oxide and molecular sieve are ground into fine powder, sieved, and then uniformly mixed by a mixer, and the particle size of the catalyst powder is controlled between 100 μm and 300 μm; further options include 100 μm to 150 μm, 150 μm to 300 μm, 150 μm to 250 μm, 250 μm to 300 μm, 100 μm to 200 μm, and 200 μm to 300 μm; preferably 150 μm.

[0053] In step 1) of the present invention, the mass concentration of the catalyst in the solvent is 30-100 g / L; further preferably 30-60 g / L, 60-100 g / L; more preferably 60 g / L.

[0054] In step 1) of the present invention, the liquid linear velocity at the nozzle of the Wenchuli injector in the jet loop reactor is controlled at 30-150 m / s, which can be selected as 30-60 m / s, 60-150 m / s, 60-120 m / s, or 120-150 m / s, and is preferably 60 m / s.

[0055] In step 1) of this invention, for a specific chemical reaction process under certain pressure and temperature conditions, the design and structural dimensions of the Venturi injector 10 greatly affect the dispersion and contact effect between reactants, thereby ultimately affecting chemical production efficiency. Combined with... Figure 2The Venturi injector 10 includes an inlet section 1002, a nozzle 1004, a mixing section 1005, and a diffuser section 1006 connected in sequence. The inlet section 1002 and the nozzle 1004 are tapered tubular. An annular air chamber 1003 is provided around the outer periphery of the inlet section 1002 and the nozzle 1004, and the air chamber 1003 is connected to the mixing section 1005. The diffuser section 1006 is diffusing tubular. The dimensions of the Venturi injector 10 include:

[0056] The ratio of the inner diameter of the inlet section 1002 opening to the inner diameter of the nozzle 1004, the inner diameter of the gas chamber 1003 closing, the length of the mixing section 1005, and the length of the diffuser section 1006 is 50:(1-5):(1.5-7.5):(15-75):(750-1500). Optionally, the ratio of the inner diameter of the inlet section 1002 opening to the inner diameter of the nozzle 1004, the inner diameter of the gas chamber 1003 closing, the length of the mixing section 1005, and the length of the diffuser section 1006 is 50:(1-4):(1.5-6):(15-40):(750-800) or 50:(4-5):(6-7.5):(40-75):(800-1500). Preferably, the ratio of the inner diameter of the inlet section 1002 of the Venturi injector 10 to the inner diameter of the nozzle 1004, the inner diameter of the gas chamber 1003, the length of the mixing section 1005, and the length of the diffusion section 1006 is 50:4:6:40:800. Within this range, the advantage is that it can draw in as much raw material gas as possible while forming finer bubbles, increasing the gas-liquid-solid mass transfer rate and thus increasing the reaction efficiency.

[0057] The opening angle α of the diffuser section 1006 is 10–40°. It can be selected as 10–25° or 25–40°. The opening angle α of the diffuser section 1006 is 25°. Within the above range, the advantage is that it prevents boundary layer separation or vortex formation, while converting kinetic energy into pressure energy, maintaining a high compression ratio (discharge side pressure / intake side pressure), and saving energy consumption in the upstream raw material transportation.

[0058] Furthermore, the bottom of the Venturi injector 10 is inserted at a position of 10-35% of the total height of the reactor 1. Optionally, the bottom of the Venturi injector 10 is inserted at a position of 10-15% or 15-35% of the total height of the reactor 1. Preferably, the bottom of the Venturi injector 10 is inserted at a position of 15% of the total height of the reactor 1. Within the above range, the advantage is that the injector structure ensures that the catalyst particles are uniformly suspended within the reactor, improving reaction selectivity and avoiding localized hot spots.

[0059] The reaction pressure in the reactor of the jet loop reactor is controlled at 7.5–15.0 MPaG, optionally 7.5–8.5 MPaG, 8.5–15.0 MPaG, 8.5–10 MPaG, 10.0–12.0 MPaG, 12.0–15.0 MPaG, or 10.0–15.0 MPaG, preferably 8.5 MPaG. The reaction temperature is controlled at 280–350℃, optionally 280–300℃, 300–320℃, or 320–350℃, preferably 300–320℃. The space velocity (WHV) of the hydrogen and supercritical carbon dioxide reaction gas in the reactor is 1000–10000 mL / (gcat·h), optionally 1000–2500 mL / (gcat·h), 2500–4000 mL / (gcat·h), 2500–3500 mL / (gcat·h), 3500–4000 mL / (gcat·h), or 4000–10000 mL / (gcat·h), preferably 2500–4000 mL / (gcat·h), and more preferably 3500 mL / (gcat·h). The molar ratio of hydrogen to supercritical carbon dioxide in the feed gas is 2.0–4.0:1; optionally 2.0–3.0:1 or 3.0–4.0:1; preferably 3:1.

[0060] In the continuous process for producing aromatics by supercritical carbon dioxide hydrogenation provided by this invention, step 2) involves sequentially passing the gas in the reactor after the reaction in step 1) through a primary condenser and a secondary condenser to condense the solvent and water, respectively. The solvent from the primary condensation is returned to the reactor, while the water from the secondary condensation is discharged as wastewater. The gas after the secondary condensation sequentially passes through an absorption tower and a desorption tower to capture non-aromatic hydrocarbons. The remaining gas is returned as recirculated gas to the Venturi ejector and then back into the reactor. Specifically:

[0061] In step 2) of this invention, the temperature of the primary condensation of the gas phase in the reactor is 180-220°C, which can be selected as 180-200°C or 200-220°C, preferably 200°C. This mainly cools down the high-boiling-point solvent in the gas phase and refluxes it into the reactor. The temperature of the secondary condensation is 30-50°C, which can be selected as 30-40°C or 40-50°C, preferably 40°C. This cools down the water in the gas phase and removes it from the system.

[0062] In step 2) of this invention, the stripping gas used in the desorption tower is fresh hydrogen. Using fresh hydrogen to strip unreacted raw material gas and recycle it back to the reactor fully utilizes both raw materials and intermediate products, saving energy.

[0063] In the continuous process for producing aromatics by supercritical carbon dioxide hydrogenation provided by this invention, step 3) involves pressurizing the reaction liquid discharged from the reactor after the reaction in step 1) via a circulating pump and then introducing it into a heat exchanger to remove the heat of reaction. The removed heat of reaction is absorbed by the hot oil on the cold side and can be used to produce medium-pressure steam. The reaction liquid after heat exchange is divided into two streams through a cross-flow filter. One stream, a permeate rich in aromatics, is sent to a distillation unit, while the other stream, a concentrated liquid containing inert solvent and catalyst particles, enters a Venturi injector and returns to the reactor for continued reaction. Specifically:

[0064] In step 3) of the present invention, the reaction liquid extracted from the jet loop reactor is divided into two streams after passing through the cross-flow filter. The mass flow ratio of the concentrated liquid returning to the jet loop reactor and the aromatic-rich permeate entering the distillation unit is (2-4):1, which can be (2-3):1 or (3-4):1, preferably 3:1.

[0065] In step 3) of this invention, the aromatic-rich permeate is fed into a distillation unit for distillation to obtain aromatic products and aromatic-rich oil. The aromatic-rich oil is used as the absorbent in the absorption tower. The absorbent in the absorption tower is the aromatic-rich oil returned after the distillation unit has removed light hydrocarbons, thereby capturing non-aromatic hydrocarbons in the gas phase inside the reactor.

[0066] The continuous process for producing aromatics by supercritical carbon dioxide hydrogenation of the present invention can, for example, employ... Figure 1 The reaction system proceeds.

[0067] The reaction system provided by this invention includes a jet-loop reactor comprising five parts: a reaction vessel 1, a circulating pump 2, a reaction heat exchanger 3, a Venturi ejector 10, and a cross-flow filter 4. The reaction vessel 1, circulating pump 2, reaction heat exchanger 3, cross-flow filter 4, and Venturi ejector 10 are sequentially connected, specifically via pipelines. At least a portion of the Venturi ejector 10 is located within the reaction vessel 1, forming a circulation loop.

[0068] Further, the reactor 1 includes a gas outlet 101, a condensate reflux inlet 102, and a reaction liquid outlet 103, the reaction liquid outlet 103 being connected to the inlet of the circulating pump 2. In some embodiments, the height-to-diameter ratio of the reactor 1 is (3-8):1, optionally, the height-to-diameter ratio of the reactor 1 is (3-5.5):1 or (5.5-8):1, preferably, the height-to-diameter ratio of the reactor 1 is 5.5:1.

[0069] Furthermore, the reaction heat exchanger 3 includes a process side 301 and a hot oil side 302, wherein the process side 301 of the reaction heat exchanger 3 is connected to the cross-flow filter 4. Optionally, the heat exchanger in this invention can be a shell-and-tube heat exchanger or a plate heat exchanger.

[0070] Furthermore, the cross-flow filter 4 includes a permeate outlet 401 and a concentrate outlet 402. The concentrate outlet 402 is connected to the Venturi injector 10 via a pipeline. A supercritical CO2 inlet 13 is provided on the connecting pipeline between the cross-flow filter 4 and the Venturi injector 10.

[0071] Furthermore, the Venturi injector 10 includes a recirculating air inlet 1001.

[0072] In the reaction system provided by the present invention, the distillation apparatus 11 includes a distillation liquid inlet 1101, a permeate inlet 1102, an absorbent outlet 1103, and an aromatic product outlet 1104.

[0073] In the reaction system provided by this invention, the circulating gas separation device consists of an absorption tower 7, a desorption tower 8, a primary condenser 5, a secondary condenser 6, and a circulating gas heater 9. The primary condenser 5, the secondary condenser 6, the absorption tower 7, and the desorption tower 8 are connected in sequence, specifically through pipelines.

[0074] Furthermore, the primary condenser 5 includes an air inlet 501 and a first condensate outlet 502. The air inlet 501 is connected to the air outlet 101 of the reactor 1. The first condensate outlet 502 is connected to the condensate reflux inlet 102 of the reactor 1. The first condensate is a solvent (e.g., tetraethylene glycol), which is refluxed back to the reactor 1 to continue the reaction.

[0075] Furthermore, the secondary condenser 6 includes a second condensate outlet 601, where the second condensate is water, which is discharged as wastewater.

[0076] Furthermore, the absorption tower 7 includes a gas phase inlet 701, a gas phase outlet 702, an absorbent liquid inlet 703, and a liquid phase outlet 704. The gas phase inlet 701 is connected to the secondary condenser 6. The gas phase outlet 702 is connected to the circulating gas heater 9. The absorbent liquid inlet 703 is connected to the absorbent liquid outlet 1103 of the distillation unit 11. The liquid phase outlet 704 is connected to the desorption tower 8.

[0077] Furthermore, the desorption tower 8 includes a liquid phase inlet 801, a hydrogen inlet 802, a gas phase outlet 803, and a liquid phase outlet 804. The liquid phase inlet 801 is connected to the liquid phase outlet 704 of the absorption tower. The gas phase outlet 803 is connected to the circulating gas heater 9. The liquid phase outlet 804 is connected to the distillation liquid phase inlet 1101.

[0078] Furthermore, the circulating gas heater 9 is connected to the circulating gas inlet 1001.

[0079] In the reaction system provided by this invention, the reaction system further includes a steam drum 12. The hot oil side 302 of the reaction heat exchanger 3 is connected to the steam drum 12 via a pipeline through the hot oil side of the circulating gas heater 9. The hot oil outlet 1201 of the steam drum 12 is connected to the process side 301 of the reaction heat exchanger 3 via a pipeline. The steam drum 12 also includes a medium-pressure steam outlet 1202.

[0080] The working process of the reaction system of the present invention is as follows:

[0081] When the jet-loop reactor is working, the circulating pump 2 starts. The reaction liquid circulates in the loop at a high flow rate, and the Venturi ejector 10 sprays at high speed, creating a negative pressure at the working nozzle 1004, which draws the gas into the Venturi ejector 10. A gas outlet 101 is located on one side of the top of the reactor 1, connected to a circulating gas separator. The reactor gas phase passes through the first-stage condenser 5 to condense the solvent and return it to the reactor 1. It then passes through the second-stage condenser 6 to condense the water, which is collected as wastewater. Afterward, it enters the absorption tower 7 to absorb non-aromatic hydrocarbons from the aromatic oil gas. The absorbed absorbent then enters the desorption tower 8, where hydrogen is used to extract the light component gas. This gas, along with the gas at the top of the absorption tower 7, is then drawn back into the reactor 1 by the Venturi ejector 10, forming a gas loop. Microbubbles with a large specific surface area are formed in the Venturi ejector 10, increasing gas-liquid contact and accelerating the reaction rate. The lower end of the Venturi ejector 10 is below the liquid surface, causing the gas-liquid mixture to impact the liquid in the reactor 1, promoting dispersion and mixing, and further accelerating the reaction. The material enters the reaction heat exchanger 3 from the bottom of the reactor 1 via the circulating pump 2, and then enters the cross-flow filter 4 for filtration. The filtrate is collected as the product and enters the reactor. Fresh supercritical CO2 is injected into the concentrate containing the heterogeneous catalyst and enters the Venturi ejector 10 from the top of the reactor 1. The reaction heat exchanger 3 removes or provides the heat released or absorbed during the reaction, controlling the reaction temperature fluctuation to ±2℃.

[0082] The beneficial effects of the present invention will be further illustrated below with reference to the embodiments.

[0083] To make the inventive objectives, technical solutions, and beneficial effects of this invention clearer, the invention is further described in detail below with reference to embodiments. However, it should be understood that the embodiments of this invention are merely for illustrative purposes and not for limiting the invention, and the embodiments are not limited to those given in the specification. Unless otherwise specified, specific experimental or operational conditions in the embodiments were prepared under conventional conditions or according to the conditions recommended by the material supplier.

[0084] Furthermore, it should be understood that the existence of other method steps before or after the combined steps, or the insertion of other method steps between these explicitly mentioned steps, does not preclude the existence of other method steps before or after the combined steps, or the insertion of other method steps between these explicitly mentioned steps, unless otherwise stated. It should also be understood that the combined connection relationship between one or more devices / apparatus mentioned in this invention does not preclude the existence of other devices / apparatus before or after the combined devices / apparatus, or the insertion of other devices / apparatus between these explicitly mentioned devices / apparatus, unless otherwise stated. Moreover, unless otherwise stated, the numbering of each method step is merely a convenient tool for identifying each method step, and not for limiting the order of the method steps or limiting the scope of the invention. Changes or adjustments to their relative relationships, without substantially altering the technical content, should also be considered within the scope of the invention.

[0085] In the following embodiments, unless otherwise specified, all the raw materials of the present invention are commercially available or prepared according to conventional methods in the art.

[0086] The metal oxides described in this invention are commercially available. Those skilled in the art can also prepare metal oxides of corresponding metal elements by selecting precursors of the metal element and using methods such as co-precipitation, solid-phase mixing, and sol-gel methods. The Cr2O3 used in this invention can be directly obtained commercially, while the ZnFeO used in this invention... x The preparation methods of -4.25Na and ZnCr2O4 are as follows.

[0087] Among them, ZnFeO x The preparation method of -4.25Na is as follows:

[0088] At 60℃, Fe(NO3)3-9H2O and Zn(NO3)2-6H2O were dissolved in a certain amount of deionized water in a Fe:Zn ratio of 3:1. Then, 1.5 mol L... -1 NaOH solution was added dropwise to the mixed metal nitrate solution until the pH of the resulting suspension reached approximately 10.0. The turbid liquid was aged at 60°C for 60 minutes, then the resulting suspension was centrifuged and washed with a certain amount of deionized water. By controlling the washing time during centrifugation, a sample with a Na concentration of 4.25% could be obtained.

[0089] The preparation method of ZnCr2O4 is as follows:

[0090] Zn(NO3)2·6H2O and Cr(NO3)3·9H2O were dissolved in a certain amount of deionized water at a ratio of Cr:Zn = 2:1. 1.5 mol L -1(NH4)2CO3 solution was added dropwise to a mixed metal nitrate solution, and precipitation was carried out at 70°C while the pH was maintained at 7.0-8.0 with stirring. After standing for 3 hours at the same temperature, the precipitate was washed three times with distilled water and then collected by filtration. The product was dried at 120°C for 8 hours and then calcined at 500°C for 3 hours to obtain the sample.

[0091] According to the present invention, the selection of the molecular sieve is not limited as long as the purpose of the present invention can be achieved. In some embodiments, the molecular sieve includes one or more of H-ZSM-5 molecular sieve, ZSM-11 molecular sieve, Silicalite-1 molecular sieve and Silicalite-2 molecular sieve.

[0092] The molecular sieves described in this invention can be obtained commercially or prepared using raw materials for molecular sieve preparation. For example, those skilled in the art can prepare a mother liquor containing tetraethyl orthosilicate, aluminum isopropoxide, and tetrapropylammonium hydroxide according to the required silicon-aluminum molar ratio for the molecular sieve. The resulting mother liquor is then subjected to hydrothermal treatment, separation, and calcination to obtain the ZSM-5 molecular sieve. Therefore, the preparation of various molecular sieves will not be elaborated upon in this invention.

[0093] The test methods used in Examples 1-5 are as follows:

[0094] Testing instrument: Agilent 7890B GC (Agilent J&W DB-HeavyWAX column, 30m×0.25mm×0.25μm, FID detector);

[0095] Test conditions: Carrier gas helium, 1 mL / min, column temperature 70℃ (10.0 min), increased to 120℃ (1.0 min) at 5℃ / min, increased to 280℃ (60 min) at 20℃ / min, detector 280℃.

[0096] Example 1

[0097] Metal oxide Cr2O3 and H-ZSM-5 molecular sieve (Si / Al molar ratio = 50) were ground and mixed at a mass ratio of 1:1 to form fine Cr2O3 / ZSM-5 catalyst powder with a particle size of 150 μm.

[0098] A tetraethylene glycol suspension containing the aforementioned Cr2O3 / ZSM-5 catalyst powder was added to a 50L injection loop reactor (reactor volume 50L) until the reactor level reached 50%. The initial catalyst concentration in the solution was 60 g / L. After sealing, CO2 and H2 (hydrogen and carbon monoxide volume ratio 3:1) were introduced into the reactor until the system pressure reached 1.0 MPaG. The circulation pump was then turned on, and the air in the reactor was replaced six times. After the last replacement, the reaction system pressure was maintained at a certain level, and the temperature was raised to the reaction temperature of 350°C through a reaction heat exchanger. Immediately afterward, a mixture of CO2 and hydrogen (hydrogen and carbon dioxide volume ratio 3:1) was added until the reactor pressure reached 8.0 MPaG. This point was recorded as the reaction start time. During the reaction, the reaction system pressure remained constant at 8.0 MPaG. The volume ratio of hydrogen to supercritical carbon dioxide entering the reaction system was 3:1, and the feed gas space velocity was 3500 mL / (gcat·h). After being heated by the reaction heat exchanger at the bottom of the reactor, the product enters the cross-flow filter. The valve on the filtrate side is opened to begin product collection, and the mass flow ratio of the concentrate returning to the jet loop reactor to the filtrate entering the downstream separation system is controlled at 3:1. The gas phase in the reactor enters the primary condenser and is cooled to 200°C, with the condensed tetraethylene glycol solvent returned to the reactor. It then enters the secondary condenser and is cooled to 40°C, with the condensed water discharged from the system. The gas then enters the absorption tower, where aromatic oil is used as the absorbent to capture non-aromatic hydrocarbons. The captured absorbent then enters the desorption tower, where fresh hydrogen is used for stripping to obtain recycle gas. This recycle gas is then drawn back to the reactor under negative pressure by a Venturi ejector to continue participating in the reaction.

[0099] Venturi ejectors in the reactor (e.g.) Figure 2 The liquid linear velocity at the nozzle (as shown) is 60 m / s. The detailed design dimensions are as follows: the ratio of the inlet opening inner diameter, nozzle inner diameter, gas chamber closing inner diameter, mixing section length, and diffuser section length is 50:4:6:40:800, and the diffuser opening angle is 25°. The lowest insertion point of the Venturi injector is located at 15% of the total reactor height. In steady state, the reactor liquid level is stable at 60 ± 5% of the total reactor height. The reactor height-to-diameter ratio is 5.5:1.

[0100] During the reaction, the average consumption rate of total CO2 and H2 gas was measured to be 2835 L(STP) / h, the circulation ratio was 1.4:1, the single-pass conversion rate of carbon dioxide was 45.5%, the total conversion rate of feed gas was 98.2%, and the selectivity of aromatics was 83.3%.

[0101] Example 2

[0102] The difference between this embodiment and Example 1 is that the gas hourly space velocity of the feed gas is 1200 mL / (gcat·h), while the other conditions are the same as in Example 1.

[0103] During the reaction, the average consumption rate of total CO2 and H2 gas was measured to be 1036 L(STP) / h, the circulation ratio was 1.2:1, the single-pass conversion rate of carbon dioxide was 48.2%, the total conversion rate of feed gas was 98.5%, and the selectivity of aromatics was 85.5%.

[0104] Example 3

[0105] The difference between this embodiment and Example 1 is that ZnCr2O4 / H-ZSM-5 molecular sieve (mass ratio 2:1) is used, while the other conditions are the same as in Example 1.

[0106] During the reaction, the average consumption rate of total CO2 and H2 gas was measured to be 2205 L(STP) / h, the circulation ratio was 2:1, the single-pass conversion rate of carbon dioxide was 35.3%, the total conversion rate of feed gas was 98.3%, and the selectivity of aromatics was 89.5%.

[0107] Example 4

[0108] The difference between this embodiment and Embodiment 1 is that ZnFeO is used. x -4.25Na / S-HZSM-5 molecular sieve (1:2), with the remaining conditions being the same as those in Example 1.

[0109] During the reaction, the average consumption rate of total CO2 and H2 gas was measured to be 2583 L(STP) / h, the circulation ratio was 1.6:1, the single-pass conversion rate of carbon dioxide was 41%, the total conversion rate of feed gas was 98.1%, and the selectivity of aromatics was 72.5%.

[0110] Example 5

[0111] The difference between this embodiment and Embodiment 1 lies in the Venturi injector (such as...) Figure 2 The liquid linear velocity at the nozzle (as shown) is 38 m / s. The detailed design dimensions are as follows: the ratio of the inner diameter of the inlet section opening, the inner diameter of the nozzle, the inner diameter of the gas chamber closing, the length of the mixing section, and the length of the diffusion section is 50:5:7.5:60:1000. The opening angle of the diffusion section is 10°. The mass flow ratio of the concentrate returning to the jet loop reactor to the filtrate entering the downstream separation system is 2:1. The rest are the same as in Example 1.

[0112] During the reaction, the average consumption rate of total CO2 and H2 gas was 1653 L(STP) / h, the circulation ratio was 1.05:1, the single-pass conversion rate of carbon dioxide was 45.2%, the total conversion rate of feed gas was 92.5%, and the selectivity of aromatics was 72.5%.

[0113] Comparative Example

[0114] In the aforementioned reaction apparatus, the reactor and Venturi ejector in the jet loop reactor are replaced with a 50L tubular fixed-bed reactor 14, and the reaction heat exchanger, circulating pump, and cross-flow filter are eliminated. The apparatus of Comparative Example 1 is as follows: Figure 3 The reactor comprises a tubular fixed-bed reactor 14, a preheater 15, a post-condenser 16, a three-phase separator 17, and a distillation unit 11 connected in sequence. Hot oil is directly introduced into the shell side of the tubular fixed-bed reactor 14 for heating and heat transfer. See the detailed process diagram below. Figure 3 .

[0115] Metal oxide Cr2O3 and HZSM-5 (Si / Al molar ratio = 50) molecular sieve particles were uniformly mixed at a mass ratio of 1:1 and then packed into the above-mentioned tubular fixed-bed reactor. After the feed gas was introduced, the reaction was carried out under the conditions of H2 / CO2 = 3.0, temperature of 350℃, pressure of 3.0 MPaG, and feed gas space velocity of 1200 mL / (gcat·h). During the process, the circulation ratio was 2:1, the single-pass conversion rate of carbon dioxide was 34.5%, the total conversion rate of feed gas was 98.2%, and the selectivity of aromatics was 71.9%.

[0116] In summary, this invention effectively overcomes the various shortcomings of the prior art and has high industrial application value.

[0117] The above description is merely a preferred embodiment of the present invention and is not intended to limit the present invention in any form or substance. It should be noted that those skilled in the art can make various improvements and additions without departing from the method of the present invention, and these improvements and additions should also be considered within the scope of protection of the present invention. Any modifications, alterations, and equivalent changes made by those skilled in the art based on the above-disclosed technical content without departing from the spirit and scope of the present invention are equivalent embodiments of the present invention. Furthermore, any modifications, alterations, and evolutions made to the above embodiments based on the essential technology of the present invention still fall within the scope of the technical solution of the present invention.

Claims

1. A continuous process for preparing aromatics by supercritical carbon dioxide hydrogenation, characterized in that, Includes the following steps: 1) Add a reaction liquid containing solvent and catalyst to the reactor. Use a Venturi injector to simultaneously inject supercritical carbon dioxide and hydrogen into the reactor, so that hydrogen dissolves in the supercritical carbon dioxide fluid and disperses into the reaction liquid. Hydrogen and supercritical carbon dioxide react with the catalyst in the reactor. 2) After the reaction in step 1), the gas in the reactor passes through a primary condenser and a secondary condenser to condense the solvent and water respectively. The solvent from the primary condenser is returned to the reactor, and the water from the secondary condenser is discharged as wastewater. The gas after the secondary condenser passes through an absorption tower and a desorption tower to capture non-aromatic hydrocarbons. The remaining gas is returned as circulating gas to the Venturi ejector and then enters the reactor. 3) After the reaction in step 1), the reaction liquid discharged from the reactor is pressurized by the circulating pump and enters the heat exchanger to remove the heat of reaction. The removed heat of reaction is absorbed by the hot oil on the cold side and can be used to produce medium-pressure steam. After heat exchange, the reaction liquid is divided into two streams through the cross-flow filter. One stream is a permeate rich in aromatics, which is sent to the distillation unit. The other stream is a concentrated liquid containing inert solvent and catalyst particles, which enters the Venturi ejector and returns to the reactor for continuous reaction.

2. The continuous process for producing aromatics by supercritical carbon dioxide hydrogenation as described in claim 1, characterized in that, In step 1), the solvent is selected from one or more of sulfolane, N-methylpyrrolidone, N-formylmorpholine, triethylene glycol, and tetraethylene glycol, preferably tetraethylene glycol.

3. The continuous process for producing aromatics by supercritical carbon dioxide hydrogenation as described in claim 1, characterized in that, The catalyst is a bifunctional catalyst composed of metal oxides and molecular sieves.

4. The continuous process for producing aromatics by supercritical carbon dioxide hydrogenation as described in claim 3, characterized in that, It also includes one or more of the following conditions: A1) The metal element in the metal oxide is selected from one or more of Cr, La, Zr, Fe, Ce, Al, Ti, Zn, In and Ga; A2) The molecular sieve is selected from one or more of H-ZSM-5 molecular sieve, ZSM-11 molecular sieve, Silicalite-1 molecular sieve and Silicalite-2 molecular sieve; A3) The silica-alumina molar ratio of the molecular sieve is greater than or equal to 20; A4) The mass ratio of metal oxide to molecular sieve in the bifunctional catalyst is 1:3 to 3:1; A5) The mixing method of the metal oxide and molecular sieve is to grind the metal oxide and molecular sieve into fine powder, then sieve them and then mix them evenly through a mixer. A6) The particle size of the catalyst powder is controlled between 100 μm and 300 μm.

5. The continuous process for producing aromatics by supercritical carbon dioxide hydrogenation as described in claim 4, characterized in that, It also includes one or more of the following conditions: A11) In feature A1), the metal oxide may or may not contain an additive; the additive is selected from one or more of the metal elements K, Na, Cu, Mn, and V; A31) In feature A3), the silicon-to-aluminum molar ratio of the molecular sieve is greater than or equal to 40; In feature A6), the particle size of the catalyst powder is controlled at 150 μm.

6. The continuous process for producing aromatics by supercritical carbon dioxide hydrogenation as described in claim 1, characterized in that, Step 1) also includes one or more of the following conditions: B1) The mass concentration of the catalyst in the solvent is 30–100 g / L; B2) The linear velocity of the liquid at the nozzle of the Venturi injector is controlled to be between 30 and 100 m / s; B3) The reaction pressure in the reactor is controlled at 7.5–15.0 MPaG; B4) The reaction temperature of the reactor is controlled at 280–350°C; B5) The space velocity of hydrogen and supercritical carbon dioxide reaction gas in the reactor is 1000-10000 mL / (gcat·h); B6) The molar ratio of hydrogen to supercritical carbon dioxide in the feed gas is 2.0 to 4.

0.

7. The continuous process for producing aromatics by supercritical carbon dioxide hydrogenation as described in claim 6, characterized in that, Step 1) also includes one or more of the following conditions: B11) In feature B1), the mass concentration of the catalyst in the solvent is 60 g / L; B31) In feature B3), the reaction pressure in the reactor is controlled at 8.0 MPaG; B41) In feature B4), the reaction temperature of the reactor is controlled at 300–320°C; B51) In feature B5), the space velocity of the hydrogen and supercritical carbon dioxide reaction gas in the reactor is 3500 mL / (gcat·h). In feature B6), the molar ratio of hydrogen to supercritical carbon dioxide in the feed gas is 3:

1.

8. The continuous process for producing aromatics by supercritical carbon dioxide hydrogenation as described in claim 1, characterized in that, Step 2) also includes one or more of the following conditions: C1) The temperature of the first-stage condenser is 180–220℃; C2) The secondary condensation temperature is 30-50℃; C3) The stripping gas used in the desorption tower is fresh hydrogen.

9. The continuous process for producing aromatics by supercritical carbon dioxide hydrogenation as described in claim 1, characterized in that, In step 3), the reaction liquid extracted from the reactor is divided into two streams after passing through a cross-flow filter. The mass flow ratio of the concentrated liquid returning to the reactor and the aromatic-rich permeate entering the distillation unit is (2-4):

1.

10. The continuous process for producing aromatics by supercritical carbon dioxide hydrogenation as described in claim 1, characterized in that, In step 3), the aromatic-rich permeate is fed into a distillation unit for distillation to obtain aromatic products and aromatic-rich oil, which is used as the absorbent in the absorption tower.

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