An industrial activation system and activation method for forming a pure-phase chi-carbide catalyst

Abstract

The application discloses an industrial activation system and method for forming pure-phase chi-carbide catalyst, which comprises a catalyst feeding tank, a first gas-solid fluidized bed reactor, a second gas-solid fluidized bed reactor, a catalyst conveying pipe and a catalyst storage tank. The catalyst reduction reaction and carbonization process are realized in one reactor, the oxidation passivation process is completed in another separate gas-solid fluidized bed, the surface purification of the catalyst is completed in the flowing and conveying process, and the catalyst returns to the first gas-solid fluidized bed reactor for carbonization after the passivation is completed, so that the pure-phase chi-carbide is formed and the industrial production is suitable.

Description

Technical Field

[0001] The present invention relates to the field of Fischer-Tropsch synthesis catalyst preparation, and in particular to an industrial activation system and activation method for forming a pure phase χ iron carbide catalyst. Background Art

[0002] Converting coal into high-quality liquid fuels and high-value-added chemicals has always attracted significant attention. Indirect coal liquefaction, a process that uses coal as a raw material, converts it into clean syngas through gasification and purification processes, and then converts it into large-molecule hydrocarbons and alcohols through the Fischer-Tropsch synthesis reaction, is an important way to achieve the rational, efficient, and clean utilization of coal.

[0003] The Fischer-Tropsch synthesis (FTS) reaction is the core of the indirect coal liquefaction process. It is the process of synthesizing hydrocarbon products from synthesis gas (H2 + CO) in the presence of metal catalysts such as iron, cobalt, and ruthenium. The primary active phase of iron-based FTS catalysts is iron carbide. Typically, prepared FTS iron-based catalysts are prepared in an oxidized state. Prior to the FTS reaction, they must be activated with reducing gases such as H2, carbon monoxide, and light olefins to form a stable iron carbide active phase. Therefore, the activation process significantly impacts the performance of FTS iron-based catalysts.

[0004] However, due to the wide variety of iron carbide types and structural forms, their activity and selectivity for hydrocarbon target products are also different. Traditional precipitated iron catalysts can be used in fixed bed and slurry bed reactors under certain temperature, pressure, inlet H2 / CO and other conditions. After activation, there are multiple iron carbides in the catalyst, and it is not a pure phase structure. CN103551207A and CN100404137C propose using a gas-solid fluidized bed reactor to activate a precipitated iron catalyst, with the activated catalyst also being a mixed phase. CN107149948A discloses a gas-solid fluidized bed Fischer-Tropsch synthesis catalyst activation method. To control the water content in the reactor, a dehydration tank is placed before the reactor. The activation temperature rise process is divided into different stages, primarily maintaining a constant temperature at 120°C for 2-8 hours and, to ensure reaction stability, maintaining a constant temperature at 220-230°C for 3-10 hours. The reaction pressure is 0.5-2 MPa, the superficial gas velocity is 0.04-0.12 m / s, and the final constant temperature is 260-280°C. The reducing gas is H2 and CO in a molar ratio of 40-200:1.

[0005] With the advancement of technology, the work "Synthesis of stable and low-CO2selective ε-iron carbide Fischer-Tropsch catalysts" published in Science Advances in 2018 also introduced the first synthesis of pure phase ε(ε′)-Fe2C, whose CO2 selectivity is close to 0. ’ The performance of the catalyst with pure phase iron carbide as the main component is significantly higher than that of existing catalysts formed by multiple components such as elemental iron, ferroferric oxide, and multiple types of iron carbide. At the same time, pure phase iron carbide can significantly reduce the selectivity of CO2, a by-product of Fischer-Tropsch synthesis, and is very suitable for the needs of modern coal chemical Fischer-Tropsch synthesis to efficiently produce more oil wax and other hydrocarbon products.

[0006] CN112569992A and CN112569975A respectively disclose methods for preparing precipitated χ iron carbide and a mixture of χ iron carbide and other types of existing iron carbide to form a pure phase iron carbide, wherein the specific surface area of ​​the precipitated pure phase iron carbide ranges from 20 to 280 m 2 / g, indicating that it has a certain pore structure. CN112569977A and CN112569976A announced the preparation method of supported χ iron carbide. CN112569975A announced a method of first forming pure phase χ iron carbide, pure phase θ iron carbide and pure phase ε / ε ’ Iron carbide is then mixed with the heterogeneous iron carbide composition catalyst.

[0007] CN112569992A, CN112569975A, CN112569983A, and CN112569985A, while disclosing the precursor preparation method, also embed the reduction process into the catalyst preparation process. Since the raw materials, processes, and techniques for preparing the precursors of different pure phase iron carbides are different, the corresponding treatment processes and parameter conditions for forming the pure phase are also different. However, the overall process is complex. The main conditions are that the prepared iron oxide or the precursor containing iron oxide is first reduced and surface purified under H2 conditions, then passivated with oxygen, and finally pure phase χ iron carbide is prepared under synthesis gas conditions.

[0008] The process for forming precipitated pure-phase x-iron carbide in patents CN112569992A and CN112569975A is embedded in the overall catalyst preparation process and is not suitable for large-scale continuous production. Large-scale catalysts are shipped in an oxidized state and are typically transported to the reactor site. A separate activation device activates the iron oxide into an active iron phase capable of Fischer-Tropsch synthesis reactions before transferring it to the Fischer-Tropsch reactor for the reaction.

[0009] Although the methods for preparing pure-phase iron carbide precursors in the above-mentioned patents are different, the precursors are basically in an oxidized state, and the subsequent activation process for forming pure-phase iron carbide is complicated. The main conditions are that the prepared iron oxide or iron oxide-containing precursor is first reduced and surface-purified under H2 conditions, followed by oxygen passivation, and finally the pure-phase iron carbide is prepared under synthesis gas conditions. During the treatment process, after the catalyst H2 reduction is completed, the temperature is lowered to near room temperature and low-concentration air is introduced for oxygen passivation. Not only is the operation complicated, but if it is carried out in the same reactor, the operation of introducing air or a low-oxygen mixed gas is not suitable for completion in an environment containing dangerous gases such as H2 and CO. Therefore, before and after oxygen enters the activation system, the system must undergo N2 replacement operation to ensure that oxygen does not come into contact with H2 and CO to avoid danger. This has problems in terms of operability and safety. The complex operation process will also affect the energy consumption level and flexibility of the activation operation.

[0010] At the same time, the optimized operating pressure of the catalyst activation process in these patents is basically lower than 10atm, and the specific equipment for the activation treatment is not clear. Therefore, large-scale industrial application cannot be achieved. It is necessary to design an activation reactor and process flow that matches the pure phase iron carbide.

[0011] CN103551207A discloses a Fischer-Tropsch catalyst reduction and activation system and process using a fixed fluidized bed or gas-solid bubbling bed. The system includes a cyclone separator to recover catalyst entrained in the gas, a catalyst scale to control the mass of catalyst entering the Fischer-Tropsch reactor, and a gas mixer to regulate the reducing gas atmosphere. This technology is suitable for general catalyst activation, specifically the formation of mixed-phase active components, but does not address the key points and technical methods for forming pure-phase iron carbide.

[0012] CN100404137C discloses an industrial reduction method using a fluidized bed granular iron-based Fischer-Tropsch synthesis catalyst. The reduction process is divided into a reduction stage and an adjustment stage according to different reaction heat stages. The reduction stage corresponds to a heating stage, and the adjustment stage corresponds to a reaction exothermic stage. The catalyst reduction process is carried out at a temperature of 260-450°C, a pressure of 2.0-5.0 MPa, and an inlet gas velocity of 0.15-0.7 m / s in an atmosphere of H2, CO, or a high hydrogen-to-carbon ratio (H2+CO). This technology is also not directed to the formation of pure-phase iron carbide, nor does it identify the key points and technical methods for forming pure-phase iron carbide.

[0013] CN104549559B discloses a method for activating an iron-based catalyst for directly preparing low-carbon olefins from a fluidized bed synthesis gas. The method includes a first activation step and a second activation step: the first activation step is to convert at least part of the iron oxide into an iron phase with Fischer-Tropsch activity; the second activation step is to perform surface treatment on the catalyst activated in the first activation step, so as to make it more suitable for the generation and timely desorption of low-carbon olefins. The catalyst is activated in situ, and after activation, it can be easily switched to a reaction state, and the operation is simple. It can be used for catalyst activation in large-scale production of low-carbon olefins from a fluidized bed synthesis gas. This technology mainly aims at the activation process selection for the purpose of low-carbon olefins as the reaction product, and is not aimed at the formation of pure-phase cementite, nor does it point out the key points and technical methods for forming pure-phase cementite, and also does not give the apparent gas velocity parameters of the fluidized bed.

[0014] CN 107149948A discloses a method for activating a gas-solid fluidized bed Fischer-Tropsch synthesis catalyst, and proposes a method for controlling the water content in the reactor and ensuring the stability of the reaction. It is also not aimed at the formation of pure-phase cementite, nor does it point out the key points and technical methods for forming pure-phase cementite.

[0015] The Fischer-Tropsch synthesis iron-based catalyst for large-scale production is generally in an oxidized state when it is delivered, and does not have catalytic activity, so it needs to be activated separately to have the state of Fischer-Tropsch synthesis catalytic performance.

[0016] At present, the activation process of the iron-based catalyst is a method of converting iron oxide into various active phases such as magnetite, elemental iron, and multi-phase cementite, and there is no activation process suitable for large-scale operation for forming pure-phase cementite as the main active component of the catalyst. SUMMARY

[0017] Therefore, the purpose of the present application is to provide an industrial activation system and method for forming a pure-phase χ cementite catalyst, to solve the problem of industrial large-scale preparation of a pure-phase χ cementite catalyst.

[0018] To achieve the above-mentioned purpose of the application, the present application adopts the following technical solutions:

[0019] An industrial activation system for forming a pure-phase χ cementite catalyst, the industrial activation system comprising a catalyst feeding tank, a first gas-solid fluidized bed reactor, a second gas-solid fluidized bed reactor, a catalyst conveying pipe, and a catalyst storage tank; wherein,

[0020] The catalyst feeding tank is connected to the first gas-solid fluidized bed reactor, and is used to supply the oxidized catalyst to be activated to the first fluidized bed reactor;

[0021] The bottom of the first gas-solid fluidized bed reactor is equipped with a first air inlet pipe for supplying the gas required in the first gas-solid fluidized bed reactor, and the top is equipped with a first tail gas pipe for discharging the gas in the first gas-solid fluidized bed reactor;

[0022] The second gas-solid fluidized bed reactor is equipped with a second gas inlet pipe at the bottom to supply the gas required in the second gas-solid fluidized bed reactor, and a second tail gas pipe at the top to discharge the gas in the second gas-solid fluidized bed reactor;

[0023] One end of the catalyst delivery pipe is connected to the lower part of the first gas-solid fluidized bed, and the other end is connected to the lower part of the second gas-solid fluidized bed, and is used to deliver the catalyst after the reduction treatment in the first gas-solid fluidized bed reactor to the second gas-solid fluidized bed reactor, and return the catalyst after the oxygen-containing atmosphere treatment in the second gas-solid fluidized bed reactor to the first gas-solid fluidized bed reactor for treatment to prepare a pure phase x-iron carbide catalyst;

[0024] The catalyst storage tank is connected to the lower part of the first gas-solid fluidized bed reactor and is used to receive the pure phase χ iron carbide catalyst product prepared from the first gas-solid fluidized bed reactor.

[0025] According to the industrial activation system of the present invention, preferably, the industrial activation system also includes a tail gas separation unit and a tail gas circulation pipe, the tail gas separation unit is used to separate the gaseous reaction products produced by the reaction in the first gas-solid fluidized bed reactor contained in the tail gas discharged from the first tail gas pipe, and the tail gas circulation pipe is connected to the tail gas separation unit, and is used to send at least part of the tail gas separated by the tail gas separation unit into the first gas-solid fluidized bed reactor as circulating gas.

[0026] According to the industrial activation system of the present invention, preferably, the tail gas separation unit comprises:

[0027] a heat exchanger for performing heat exchange and cooling on the tail gas leaving the first gas-solid fluidized bed reactor and transported through the first tail gas pipe to recover heat;

[0028] The first gas-liquid separation tank is used to separate the gas and liquid of the tail gas after being cooled by the first heat exchanger to remove the liquid phase;

[0029] A cooler for cooling the remaining water and organic matter in the tail gas from the first gas-liquid separation tank for condensation;

[0030] a second gas-liquid separation tank, used for performing gas-liquid separation on the tail gas from the cooler to remove the liquid phase to obtain circulating gas; and

[0031] The dehydration adsorption tank is used to dry and dehydrate at least a portion of the circulating gas from the second gas-liquid separation tank to obtain dried circulating gas.

[0032] According to the industrial activation system of the present application, preferably, the heat exchanger is used to exchange heat and cool down the tail gas delivered through the first tail gas pipe and the feed gas delivered through the first feed gas pipe to be introduced into the first gas-solid fluidized bed reactor to recover heat.

[0033] According to the industrial activation system of the present application, preferably, the industrial activation system further comprises:

[0034] a first heater arranged on the first feed gas pipe and used to heat the feed gas from the heat exchanger;

[0035] a second heater arranged on the second feed gas pipe and used to heat the feed gas to be introduced into the second gas-solid fluidized bed reactor; and

[0036] a circulating compressor arranged on the tail gas circulating pipe and used to compress the tail gas from the tail gas circulating pipe and send the pressurized tail gas into the first feed gas pipe.

[0037] To achieve the above-mentioned object, the present application further provides an activation method for forming pure-phase χ-carbide iron catalyst using the above-mentioned industrial activation system, wherein the activation method comprises:

[0038] (1) reducing the catalyst in an oxidized state using hydrogen in the first gas-solid fluidized bed reactor;

[0039] (2) transferring the catalyst reduced by step (1) into the second gas-solid fluidized bed reactor and performing oxidation passivation treatment on the catalyst using an oxygen-containing atmosphere;

[0040] (3) transferring the catalyst oxidized and passivated by step (2) back to the first gas-solid fluidized bed reactor and performing carbonization treatment on the catalyst using synthesis gas to obtain the pure-phase χ-carbide iron catalyst.

[0041] According to the activation method of the present application, preferably, the reduction treatment in step (1) is performed under the following conditions: the reaction atmosphere is H2 atmosphere, the reaction temperature is 240-350℃, the reaction pressure is 1.0-4.0 Mpa, and the reaction time is 1-24 h.

[0042] According to the activation method of the present application, preferably, the oxidation passivation treatment in step (2) is performed under the following conditions: the reaction atmosphere is oxygen-containing nitrogen gas with an oxygen volume content of 0.5-2%, the reaction temperature is 150-300℃, the reaction pressure is 1.0-4.0 Mpa, and the reaction time is 1-24 h.

[0043] According to the activation method of the present invention, preferably, the carbonization treatment conditions in step (3) are: the reaction atmosphere is synthesis gas with a hydrogen-to-carbon ratio of 10-100:1, the reactor pressure is 1.0-4.0 MPa, the reaction temperature is 300-500°C, and the reaction time is 1-36 hours.

[0044] According to the activation method of the present invention, preferably, the H2 space velocity during the reduction treatment in step (1) is 600-6000Nm 3 / h / t, preferably 1000-4000Nm 3 / h / t, during the reduction process, the fluidized bed reactor adopts a bubbling state, and the superficial gas velocity of the gas at the inlet is 0.1-0.5m / s, preferably 0.15-0.3m / s.

[0045] According to the activation method of the present invention, preferably, during the oxidative passivation treatment in step (2), the oxygen space velocity is 100-800 Nm 3 / h / t, preferably 200-500Nm 3 / h / t, the oxidation reactor adopts bubbling reaction, and the superficial gas velocity is 0.1-0.4m / s, preferably 0.15-0.3m / s.

[0046] According to the activation method of the present invention, preferably, during the carbonization treatment in step (3), the synthesis gas space velocity is 500-5000Nm 3 / h / t, preferably 1000-3000Nm 3 / h / t, during the carbonization process, the bed of the fluidized bed reactor adopts a turbulent state, and the superficial velocity of the gas at the inlet is 0.2-0.7 m / s, preferably 0.3-0.6 m / s.

[0047] According to the activation method of the present invention, preferably, when heating in step (1), the temperature is raised to 100-150°C at a rate of 15-25°C / h and maintained for 1-2 hours, and then the temperature is continued to be raised at a rate of 10-20°C / h to reach the reaction temperature of the reduction treatment.

[0048] According to the activation method of the present invention, preferably, in steps (2) and (3), when heating, the heating rate is 10-20°C / h to reach the reaction temperature.

[0049] According to the activation method of the present invention, preferably, when the catalyst is transferred between different devices, the catalyst is delivered to the target device by nitrogen delivery using the pressure difference.

[0050] According to the activation method of the present invention, preferably, the catalyst in the catalyst feeding tank is maintained at 40-60°C under N2 conditions, and then the catalyst is added to the first gas-solid fluidized bed reactor by N2 gas transportation using the pressure difference; the catalyst storage tank is an N2 atmosphere storage tank.

[0051] Compared with the prior art, the present invention has the following advantages:

[0052] 1) The present invention clearly proposes an activation reaction device and process flow for forming precipitated pure phase χ iron carbide that can be scaled up, and enables the advanced pure phase iron carbide catalyst to be used in a Fischer-Tropsch synthesis device; based on the main parameter characteristics of the formation of precipitated χ iron carbide, an intermittent operation mode is adopted to form pure phase χ iron carbide under optimized conditions, the catalyst reduction reaction and carbonization process are realized in one reactor, the oxidation passivation process is completed in another separate gas-solid fluidized bed, the surface purification of the catalyst is completed during the flow and transportation process, and after the passivation is completed, the catalyst is returned to the first gas-solid fluidized bed reactor of the fluidized bed for carbonization to form precipitated pure phase χ iron carbide, which can be transferred to the Fischer-Tropsch synthesis reactor by gas transportation.

[0053] (2) The present invention adopts the conditions of increasing pressure and high temperature oxidation during oxidation passivation, which not only saves energy but also avoids catalyst crushing due to excessive temperature difference and improves yield.

[0054] (3) Catalyst storage tank with N2 atmosphere to achieve offline storage of pure phase iron carbide. BRIEF DESCRIPTION OF THE DRAWINGS

[0055] Figure 1 A schematic diagram of an embodiment of the industrial activation system provided by the present invention;

[0056] Description of reference numerals:

[0057] 1-heat exchanger; 2-first heater; 3-first gas-solid fluidized bed reactor; 4-catalyst feeding tank; 5-first gas-liquid separation tank; 6-cooler; 7-second gas-liquid separation tank; 8-dehydration adsorption tank; 9-circulating compressor; 10-second gas-solid fluidized bed reactor; 11-second heater; 12-catalyst storage tank; 13-catalyst delivery pipe; 14-exhaust circulation pipe; 15-second air inlet pipe; 16-second exhaust pipe; 31-first air inlet pipe; 32-first exhaust pipe. DETAILED DESCRIPTION

[0058] The present invention is further described below with reference to the embodiments and drawings. However, the present invention is not limited to the listed embodiments but also includes equivalent improvements and variations of the technical solutions defined in the claims attached to the present application.

[0059] The catalyst raw material of the activation process of the present application is an iron-based Fischer-Tropsch synthesis catalyst in which iron is in an oxidation state, such as the catalyst precursor prepared in CN112569982A, CN112569993A or CN112569987A, in which the Fe component is in an oxidation state; wherein the average particle size of the finished oxidation state catalyst particles as raw material is preferably 70-80 microns, which belongs to the Geldart A particle range in the field of gas-solid fluidization; preferably, the catalyst attrition as raw material is less than 4g / h.

[0060] In one embodiment, as shown in Figure 1 The industrial activation device of the present application includes a catalyst feeding tank 4, a first gas-solid fluidized bed reactor 3, a second gas-solid fluidized bed reactor 10, a catalyst conveying pipe 13 and a catalyst storage tank 12; wherein,

[0061] The catalyst feeding tank 4 is connected to the first gas-solid fluidized bed reactor 3 for supplying the oxidation state catalyst to be activated to the first fluidized bed reactor as a catalyst raw material; the specific conveying method is well known in the art, such as spiral conveying or pneumatic conveying can be used.

[0062] The bottom of the first gas-solid fluidized bed reactor 3 is equipped with a first gas inlet pipe 31 to supply the required gas in the first gas-solid fluidized bed reactor to fluidize the material in the reactor, and the top is equipped with a first tail gas pipe 32 to discharge the gas in the first gas-solid fluidized bed reactor as the tail gas after reaction.

[0063] The bottom of the second gas-solid fluidized bed reactor 10 is equipped with a second gas inlet pipe 15 to supply the required gas in the second gas-solid fluidized bed reactor to fluidize the material in the reactor, and the top is equipped with a second tail gas pipe 16 to discharge the gas in the second gas-solid fluidized bed reactor as the tail gas after reaction.

[0064] One end of the catalyst delivery pipe 13 is connected to the lower part of the first gas-solid fluidized bed 3 and the other end is connected to the lower part of the second gas-solid fluidized bed 10, and is used to send the catalyst after the reduction treatment in the first gas-solid fluidized bed reactor into the second gas-solid fluidized bed reactor and send the catalyst after the oxygen-containing atmosphere treatment in the second gas-solid fluidized bed reactor back to the first gas-solid fluidized bed reactor for treatment to prepare a pure phase χ iron carbide catalyst. Specifically, the raw catalyst from the catalyst feeding tank is first reduced in the first gas-solid fluidized bed reactor under a hydrogen atmosphere, and then enters the second gas-solid fluidized bed reactor through the catalyst delivery pipe for oxidation passivation treatment in an oxidizing atmosphere, and finally returns to the first gas-solid fluidized bed reactor through the catalyst delivery pipe for carbonization treatment in a synthesis gas atmosphere to prepare a pure phase χ iron carbide catalyst. It can be understood that the smooth delivery of the catalyst can be achieved by regulating the pressure difference between the first and second gas-solid fluidized bed reactors.

[0065] The catalyst storage tank 12 is connected to the lower part of the first gas-solid fluidized bed reactor 3 and is used to receive the pure phase χ iron carbide catalyst product prepared from the first gas-solid fluidized bed reactor.

[0066] In some embodiments, the industrial activation system further includes a tail gas separation unit and a tail gas circulation pipe 14, wherein the tail gas separation unit is used to separate the gaseous reaction products produced by the reaction in the first gas-solid fluidized bed reactor contained in the tail gas discharged from the first tail gas pipe, such as water and organic matter (such as oil phase and light hydrocarbon products), and the tail gas circulation pipe is connected to the tail gas separation unit, and is used to send at least part of the tail gas separated by the tail gas separation unit as circulating gas into the first gas-solid fluidized bed reactor to recycle and utilize the effective gas therein.

[0067] The tail gas separation unit comprises:

[0068] The heat exchanger 1 is used to heat and cool the tail gas leaving the first gas-solid fluidized bed reactor and transported through the first tail gas pipe 32 to recover heat;

[0069] The first gas-liquid separation tank 5 is used to separate the tail gas after being cooled by the first heat exchanger into gas and liquid to remove the liquid phase;

[0070] Cooler 6, used to cool the remaining water and organic matter in the tail gas from the first gas-liquid separation tank for condensation;

[0071] a second gas-liquid separation tank 7, used for performing gas-liquid separation on the tail gas from the cooler to remove the liquid phase to obtain circulating gas; and

[0072] The dehydration adsorption tank 8 is used to dry and dehydrate at least a portion of the circulating gas from the second gas-liquid separation tank to obtain dried circulating gas.

[0073] Of course, it is understandable that when the water content in the circulating gas from the second gas-liquid separation tank 7 meets the requirements, it can also be directly sent out without being processed by the dehydration adsorption tank 8.

[0074] In some embodiments, the heat exchanger 1 is used to exchange heat and cool the tail gas transported through the first tail gas pipe 32 with the intake gas transported through the first intake pipe 31 to enter the first gas-solid fluidized bed reactor 3 to recover heat;

[0075] The industrial activation system also includes:

[0076] a first heater 2, provided on the first air inlet pipe 13, for heating the air inlet from the heat exchanger 1, so as to regulate the temperature in the first gas-solid fluidized bed;

[0077] A second heater 11 is provided on the second air inlet pipe and is used to heat the air to be fed into the second gas-solid fluidized bed reactor so as to regulate the temperature in the second gas-solid fluidized bed; and

[0078] The circulating compressor 9 is provided on the exhaust gas circulation pipe 14 and is used for compressing the exhaust gas from the exhaust gas circulation pipe and sending the pressurized exhaust gas into the first air intake pipe.

[0079] During operation, the industrial activation system of the present invention first completes the stable operation of the gas circulation under N2 conditions. After the N2 is introduced into the system and replaced, the recycle compressor 9 is started. Fresh N2 is mixed with the recycle gas from the recycle compressor 9 to form the inlet gas. After passing through the heat exchanger 1 and the first heater 2, it enters the first gas-solid fluidized bed reactor 3. The gas exiting the reactor is used as the recycle gas. It passes through the heat exchanger 1, the first gas-liquid separator 5, the cooler 6, the second gas-liquid separator 7, and the dehydration adsorption tank 8, and then enters the recycle compressor 9. It is mixed with fresh N2 to form the inlet gas, and the cycle continues. When the system pressure reaches the set value, the gas is recirculated to maintain the system pressure stable. In this cycle, the pressure range of the first gas-solid fluidized bed reactor is 1.0-4.0 MPa, such as 2, 2.5, 3, or 3.5 MPa.

[0080] Then, H2 is introduced into the system to switch to a mixed gas circulation of H2 and N2, or completely replaced with H2 circulation. The system pressure is kept stable during the gas circulation process. When H2 is introduced, part of it is discharged as exhaust gas to adjust the ratio of H2 and N2 in the system.

[0081] The catalyst in the catalyst feeding tank is maintained at 40-60°C under N2 conditions, and under certain pressure difference conditions, the catalyst is added to the first gas-solid fluidized bed activation reactor by N2 gas transportation.

[0082] The temperature of the first gas-solid fluidized bed reactor is raised under a hydrogen atmosphere of H2or a mixture of H2and N2, for example, at a rate of 15-25°C 20°C / h to 100-150°C, such as 110, 120, 130, or 140°C, for 1-2 hours, and then the temperature is continuously raised at a rate of 10-20°C / h, and after the temperature reaches the set temperature, the reactor is treated for 1-24 hours, such as 4, 8, 12, 16, or 20 hours, at a temperature in the range of 240-350°C, such as 260, 280, 300, or 330°C.

[0083] After the reduction is complete, the catalyst is transferred to a separate second gas-solid fluidized bed reactor for oxidative passivation treatment by nitrogen delivery through a catalyst delivery pipe. Nitrogen is first introduced to displace the H2to ensure that the H2content is less than 1%, and then nitrogen containing 0.5-2%, such as 1% or 1.5%, oxygen is introduced, the temperature is maintained in the range of 150-300°C, such as 160, 180, 200, 250, 280, 300, or 330°C, and the pressure is maintained in the range of 1.0-4.0 MPa, such as 2, 2.5, 3, or 3.5 MPa, for 1-24 hours, such as 4, 8, 12, 16, or 20 hours, to complete the oxidative passivation treatment.

[0084] After the oxidative passivation is complete, the second gas-solid fluidized bed reactor is replaced with nitrogen and returned to the first gas-solid fluidized bed reactor by nitrogen delivery to perform carbonization treatment.

[0085] The fresh gas for the carbonization treatment is H2and carbon monoxide, and the synthesis gas obtained by mixing the fresh gas with the recycled gas is used as the feed gas, which is sequentially passed through a first heat exchanger and a heater, and then introduced into the fluidized bed reactor. The gas exiting the reactor is used as tail gas, which is sequentially passed through a heat exchanger, a first gas-liquid separator, a water cooler, a second gas-liquid separator, and a dehydration adsorption tank, and then introduced into a recycling compressor as recycled gas, which is mixed with fresh gas (H2+ CO) to form the feed gas, and the cycle is repeated. The pressure in the reactor during carbonization is in the range of 1.0-4.0 MPa, such as 2, 2.5, 3, or 3.5 MPa, and the temperature is raised at a rate of 10-20°C / h to 300-500°C, such as 320, 350, 380, 400, 450, or 480°C, and maintained for 1-36 hours, such as 4, 8, 12, 16, 20, 24, 30, or 32 hours, to complete the carbonization treatment.

[0086] During the reduction process, the H2space velocity as fresh gas feed is controlled according to the amount of catalyst in the range of 600-6000 Nm 3 / h / t, preferably 1000-4000 Nm 3 / h / t, such as 2000 or 3000 Nm 3 / h / t; preferably, during the reduction process, the first gas-solid fluidized bed reactor adopts a bubbling state, and the superficial gas velocity of the gas at the inlet is 0.1-0.5 m / s, preferably 0.15-0.3 m / s, such as 0.2 or 0.25 m / s.

[0087] During the oxidation passivation process, N2 is fed as fresh gas and mixed with oxygen, where the oxygen content is 0.5-2%, such as 1% or 1.5%. The oxygen space velocity is controlled at 100-800 Nm according to the amount of catalyst. 3 / h / t, preferably 200-500Nm 3 / h / t For example 300 or 400Nm 3 / h / t; preferably, the second gas-solid fluidized bed reactor adopts a bubbling reaction, with an apparent gas velocity of 0.1-0.4 m / s, preferably 0.15-0.3 m / s, such as 0.2 or 0.25 m / s.

[0088] During the carbonization process, the hydrogen-carbon ratio of the synthesis gas (H2+CO) used for carbonization is 10-100:1, preferably 30-80:1, such as 50:1 or 70:1. The synthesis gas space velocity is controlled at 500-5000Nm according to the amount of catalyst. 3 / h / t, preferably 1000-3000Nm 3 / h / t For example 1500, 2000 or 2500Nm 3 / h / t; preferably, during the carbonization process, the bed of the first gas-solid fluidized bed reactor adopts a turbulent state, and the superficial velocity of the gas at the inlet is 0.2-0.7 m / s, preferably 0.3-0.6 m / s, such as 0.4 or 0.5 m / s.

[0089] It is understood in the art that, as needed, the catalyst can be further fed from the catalyst feeding tank to the first gas-solid fluidized bed reactor to perform catalyst reduction, passivation and carbonization operations.

[0090] The present invention is further described below with reference to Examples / Comparative Examples, wherein a catalyst having an average particle size of 70-80 μm of catalyst screening particles prepared according to step (1) of Example 1 of CN112569982A is used as a raw catalyst.

[0091] Example 1

[0092] The first gas-solid fluidized bed reactor used for reduction and carbonization treatment has a diameter of 0.5m and a height of 8m, with a catalyst processing capacity of 80-150kg / batch. The second gas-solid fluidized bed reactor used for oxidation passivation treatment has a diameter of 0.9m and a height of 5m.

[0093] After the replacement of H2 atmosphere, 100 kg of catalyst was added into the first gas-solid fluidized bed reactor through the catalyst feeding tank of the first gas-solid fluidized bed, the temperature was increased to 120°C at a heating rate of 20°C / h and kept constant for 1 h, and then the temperature was continuously increased at a rate of 10°C / h while H2 was introduced, the hydrogen flow rate was 200 Nm 3 / h, and after the temperature reached 300°C, it was kept constant for 12 hours to complete the catalyst reduction. During the reduction process, the reactor pressure was maintained at 2.0 MPa, and the inlet superficial gas velocity was 0.2 m / s.

[0094] After the reduction treatment was completed, the catalyst in the first gas-solid fluidized bed reactor was transferred to the second gas-solid fluidized bed reactor using pressure difference and pneumatic conveying method. The first gas-solid fluidized bed reactor stopped feeding H2, maintained hydrogen atmosphere circulation, and the reactor temperature was reduced to the range of 150-200°C.

[0095] After the transfer was completed, nitrogen was introduced to replace the reactor until the H2 content was less than 1%. After the replacement was completed, nitrogen gas containing 0.6% oxygen was introduced, and the oxygen flow rate was 30 Nm 3 / h, the reactor pressure was 2.0-2.2 MPa, the reaction temperature was 180-200°C, and the oxidation reaction was carried out for 12 hours. After the oxidation reaction, nitrogen was introduced to make the oxygen content of the oxidation passivation system less than 0.1%, and the passivation operation was completed.

[0096] Before the catalyst was transferred again, the pressure of the first gas-solid fluidized bed reactor used for carbonization treatment was maintained at 1.6-1.8 MP, and after the passivation reaction was completed, the catalyst was returned to the first gas-solid fluidized bed reactor through nitrogen conveying. Hydrogen was introduced to increase the pressure to 2.5 MPa, and CO+H2 feeding was switched, the hydrogen to carbon ratio was 40, and the total flow rate of CO+H2 was 200 Nm 3 / h; at the same time, the reactor temperature was increased to 340°C at a rate of 10°C / h and maintained for 10 hours to complete the catalyst carbonization and form pure phase precipitated χ-iron carbide. During the heating and constant temperature process, the flow rate of the circulating compressor was adjusted to maintain the inlet gas superficial gas velocity of the fluidized bed reactor within the range of 0.4-0.45 m / s during the carbonization process. After the carbonization reaction was completed, the temperature was reduced to 120°C, the CO+H2 feeding was stopped, and hydrogen circulation was maintained. The carbonization process circulating gas can not pass through the dehydration adsorption tank.

[0097] During the cooling process, the catalyst storage tank was replaced with N2 atmosphere, and the pressure was controlled at 1.5 MPa. After the fluidized bed reactor was cooled to 120°C, the catalyst was transferred to the catalyst storage tank for replacement under nitrogen condition for standby use.

[0098] According to the above steps, after a batch of catalyst is completed, the next batch of catalyst can be added to form an intermittent activation process.

[0099] The catalyst in the catalyst storage tank was taken for X-ray diffractometry testing, and the obtained product was a monoclinic χ-iron carbide phase. The generated target product χ-iron carbide had good crystallinity, completely corresponding to all characteristic peaks of χ-iron carbide, and was free of impurities. The obtained iron carbide was a pure-phase χ-iron carbide catalyst. At the same time, the catalyst loss during the reduction, passivation and carbonization processes was very small, and the catalyst yield calculated based on the iron element was 99.5%.

[0100] Example 2

[0101] Similar to Example 1, the relevant parameters of the reduction, oxidation and carbonization treatments were changed.

[0102] First, the atmosphere of the reduction reaction system was replaced with H2 atmosphere, and the pressure was controlled at 1.5 MPa. 100 kg of catalyst was added to the reactor through the catalyst feeding tank, and the H2 feed rate was maintained at 80 Nm 3 / h, reactor temperature 250℃, superficial gas velocity 0.4m / s, constant temperature reaction for 8h.

[0103] During the catalyst reduction reaction, the second gas-solid fluidized bed reactor is heated to 120°C under an N2 atmosphere, and the pressure is controlled to be 1.0MPa. After the reduction reaction is completed, the catalyst in the first gas-solid fluidized bed reactor is transferred to the second gas-solid fluidized bed reactor by pneumatic conveying using the pressure difference. The first gas-solid fluidized bed reactor stops H2 feeding, maintains the hydrogen atmosphere circulation, and reduces the reactor temperature to within the range of 150-200°C. After the transfer is completed, nitrogen is introduced into the second gas-solid fluidized bed reactor for replacement until the H2 content in the reactor is lower than 1%. At the same time, the temperature of the second gas-solid fluidized bed reactor is required to be maintained within the range of 120-200°C and the pressure is maintained within the range of 1.0-1.2MPa. After the replacement is completed, a nitrogen + oxygen mixed gas containing 0.5% oxygen is introduced, where the oxygen flow rate is 20Nm 3 / h, reactor pressure 1.0-1.2MPa, temperature 200-220℃, after 5 hours of oxidation reaction, nitrogen feed is changed to make the oxygen content of the oxidation passivation system lower than 0.1% and then the passivation operation is completed.

[0104] Before the catalyst is oxidatively passivated, the pressure of the fluidized bed reactor used for carbonization is maintained at 1.0-1.2 MPa. After the passivation reaction is completed, the catalyst is returned to the fluidized bed reactor by nitrogen transport. Hydrogen is introduced to boost the system pressure to 1.5 MPa, and the feed is switched to CO+H2, with a hydrogen-to-carbon ratio of 80 and a total CO+H2 flow rate of 200 Nm 3 / h; Simultaneously, the reactor temperature was raised to 300°C at a rate of 10°C / h and maintained at this temperature for 12 hours to complete catalyst carbonization and form precipitated pure-phase precipitated iron carbide. During the heating and temperature-maintaining processes, the flow rate of the circulating compressor was adjusted to maintain the superficial velocity of the inlet gas into the fluidized bed reactor within the range of 0.4-0.45 m / s during the carbonization process. After the carbonization reaction was complete, the temperature was lowered to 120°C, the CO+H2 feed was stopped, and hydrogen circulation was maintained.

[0105] During the carbonization reaction, the catalyst storage tank was replaced with a nitrogen atmosphere at a controlled pressure of 0.7 MPa. After the carbonization reaction was completed, the catalyst was transferred to a finished catalyst tank. X-ray diffractometer analysis revealed that the resulting product was a monoclinic χ-iron carbide phase. The resulting target χ-iron carbide exhibited excellent crystallinity, closely matching all characteristic peaks of χ-iron carbide, and was free of impurities. The resulting iron carbide was a pure-phase χ-iron carbide catalyst. Catalyst losses during reduction, passivation, and carbonization were minimal, resulting in a catalyst yield of 99.8% based on elemental iron.

[0106] Example 3

[0107] Similar to Example 1, the reactor size and catalyst throughput were further increased. The fluidized bed reactor used for reduction and carbonization had a diameter of 1.5 m and a height of 20 m, with a catalyst throughput of 7,000-10,000 kg / batch. The reactor used for oxidation and passivation had a diameter of 1.8 m and a height of 12 m.

[0108] First, the atmosphere of the reduction reaction system was replaced with H2 atmosphere, and the pressure was controlled at 3.0 MPa. 8000 kg of catalyst was added to the reactor through the catalyst feeding tank, and the H2 feed rate was maintained at 20000 Nm 3 / h, reactor temperature 350℃, superficial gas velocity 0.2m / s, constant temperature reaction for 10h.

[0109] During the catalyst reduction isothermal reaction, the second gas-solid fluidized bed reactor is heated to 120°C under N2 atmosphere and the pressure is controlled at 2.5MPa. After the reduction isothermal reaction is completed, the catalyst in the first gas-solid fluidized bed reactor is transferred to the second gas-solid fluidized bed reactor by pneumatic conveying using the pressure difference. The H2 feed is stopped in the first gas-solid fluidized bed reactor, the hydrogen atmosphere circulation is maintained, and the reactor temperature is reduced to the range of 200-250°C.

[0110] After the transfer is completed, nitrogen is introduced into the second gas-solid fluidized bed reactor for replacement until the H2 content in the reactor is less than 1%. After the replacement is completed, a nitrogen + oxygen mixed gas containing 1% oxygen is introduced, where the oxygen flow rate is 1800Nm 3 / h, reactor pressure 2.5MPa, reaction temperature 160-180℃, after 10 hours of oxidation reaction, nitrogen feed is changed to complete the passivation operation after the oxygen content of the oxidation passivation system is lower than 0.1%.

[0111] Before the catalyst is oxidatively passivated, the pressure of the first fluidized bed reactor to be used for carbonization is maintained at 2.0-2.2 MPa. After the passivation reaction is completed, the catalyst is returned to the fluidized bed reactor by nitrogen delivery. Hydrogen is introduced to boost the system pressure to 3.0 MPa, and the feed is switched to CO+H2, with a hydrogen-to-carbon ratio of 50 and a total CO+H2 flow rate of 24,000 Nm 3 / h; simultaneously, the reactor temperature was raised to 380°C at a rate of 10°C / h and maintained at this temperature for 12 hours to complete catalyst carbonization and form precipitated pure-phase precipitated iron carbide. During the heating and temperature-maintaining processes, the flow rate of the circulating compressor was adjusted to maintain the superficial velocity of the inlet gas into the fluidized bed reactor within the range of 0.4-0.5m / s. After the carbonization reaction was completed, the temperature was lowered to 120°C, the CO+H2 feed was stopped, and hydrogen circulation was maintained. The pressure was reduced to 1.5MPa. The circulating gas during the carbonization process did not pass through the dehydration adsorption tank.

[0112] During the carbonization reaction, the catalyst storage tank was replaced with a nitrogen atmosphere at a controlled pressure of 1.0 MPa. After the carbonization reaction was completed, the catalyst was transferred to a finished catalyst tank. The catalyst in the catalyst storage tank was then analyzed by X-ray diffractometer, revealing a monoclinic χ-iron carbide phase. The resulting target χ-iron carbide exhibited excellent crystallinity, corresponding closely to all characteristic peaks of χ-iron carbide, and was free of any impurities. The resulting iron carbide was a pure-phase χ-iron carbide catalyst. Catalyst losses during reduction, passivation, and carbonization were minimal, resulting in a catalyst yield of 99.6% based on elemental iron.

[0113] Comparative Example 1

[0114] The difference from Example 1 is that, referring to the manufacturing process in CN112569992A, the temperature during the oxidation treatment is controlled at 30° C. and the oxygen content of the oxidizing atmosphere is 3%. The rest are the same.

[0115] X-ray diffractometer analysis of the catalyst from the catalyst storage tank revealed a predominantly monoclinic χ-iron carbide phase, with approximately 5% of other crystalline phases present. The target χ-iron carbide exhibited excellent crystallinity, corresponding well to all characteristic peaks of χ-iron carbide, and was free of any impurities. The resulting iron carbide was a pure-phase χ-iron carbide catalyst. The catalyst yield, calculated based on elemental iron, was 97.4%.

[0116] However, compared with Example 1, the overall operation process needs to be cooled and depressurized first, and then warmed and pressurized after the batch change is completed, the operation process is complex, the overall operation period is significantly increased, and the catalyst yield is lower than that of Example 1, which is not conducive to large-scale production operation.

Claims

1. An industrial activation system for forming a pure phase χ iron carbide catalyst, characterized in that The industrial activation system includes a catalyst feeding tank, a first gas-solid fluidized bed reactor, a second gas-solid fluidized bed reactor, a catalyst delivery pipe and a catalyst storage tank; wherein, The catalyst feeding tank is connected to the first gas-solid fluidized bed reactor and is used to supply the first fluidized bed reactor with the oxidized catalyst to be activated; The bottom of the first gas-solid fluidized bed reactor is equipped with a first air inlet pipe for supplying the gas required in the first gas-solid fluidized bed reactor, and the top is equipped with a first tail gas pipe for discharging the gas in the first gas-solid fluidized bed reactor; The second gas-solid fluidized bed reactor is equipped with a second gas inlet pipe at the bottom to supply the gas required in the second gas-solid fluidized bed reactor, and a second tail gas pipe at the top to discharge the gas in the second gas-solid fluidized bed reactor; One end of the catalyst delivery pipe is connected to the lower part of the first gas-solid fluidized bed, and the other end is connected to the lower part of the second gas-solid fluidized bed, and is used to deliver the catalyst treated with hydrogen reduction in the first gas-solid fluidized bed reactor to the second gas-solid fluidized bed reactor, and to return the catalyst treated with oxygen-containing atmosphere in the second gas-solid fluidized bed reactor to the first gas-solid fluidized bed reactor for carbonization treatment with synthesis gas to prepare a pure phase x-iron carbide catalyst; The catalyst storage tank is connected to the lower part of the first gas-solid fluidized bed reactor and is used to receive the pure phase χ iron carbide catalyst product prepared from the first gas-solid fluidized bed reactor.

2. The industrial activation system according to claim 1, characterized in that: The industrial activation system also includes a tail gas separation unit and a tail gas circulation pipe. The tail gas separation unit is used to separate the gaseous reaction products produced by the reaction in the first gas-solid fluidized bed reactor from the tail gas discharged from the first tail gas pipe. The tail gas circulation pipe is connected to the tail gas separation unit and is used to send at least part of the tail gas separated by the tail gas separation unit into the first gas-solid fluidized bed reactor as circulating gas.

3. The industrial activation system according to claim 2, characterized in that The tail gas separation unit comprises: a heat exchanger for performing heat exchange and cooling on the tail gas leaving the first gas-solid fluidized bed reactor and transported through the first tail gas pipe to recover heat; The first gas-liquid separation tank is used to separate the gas and liquid of the tail gas after being cooled by the first heat exchanger to remove the liquid phase; A cooler for cooling the remaining water and organic matter in the tail gas from the first gas-liquid separation tank for condensation; a second gas-liquid separation tank, used for performing gas-liquid separation on the tail gas from the cooler to remove the liquid phase to obtain circulating gas; and The dehydration adsorption tank is used to dry and dehydrate at least a portion of the circulating gas from the second gas-liquid separation tank to obtain dried circulating gas.

4. The industrial activation system according to claim 3, characterized in that The heat exchanger is used to exchange heat and cool the tail gas transported through the first tail gas pipe and the intake air transported through the first intake pipe to enter the first gas-solid fluidized bed reactor to recover heat; The industrial activation system also includes: a first heater, disposed on the first air intake pipe, for heating the intake air from the heat exchanger; a second heater, provided on the second air inlet pipe, for heating the air to be introduced into the second gas-solid fluidized bed reactor; and The circulating compressor is arranged on the exhaust gas circulation pipe, and is used for compressing the exhaust gas from the exhaust gas circulation pipe and sending the pressurized exhaust gas into the first intake pipe.

5. A method for activating a pure phase x-iron carbide catalyst using the industrial activation system according to any one of claims 1 to 4, the activation method comprising: (1) In the first gas-solid fluidized bed reactor, the oxidized catalyst is reduced using hydrogen; (2) transferring the catalyst after the reduction treatment in step (1) to a second gas-solid fluidized bed reactor, and performing an oxidative passivation treatment on the catalyst using an oxygen-containing atmosphere; (3) The catalyst after the oxidation passivation treatment in step (2) is transferred back to the first gas-solid fluidized bed reactor, and the catalyst is carbonized using synthesis gas to obtain a pure phase χ iron carbide catalyst.

6. The activation method according to claim 5, characterized in that The conditions for the reduction treatment in step (1) are: the reaction atmosphere is H2 atmosphere, the reaction temperature is 240-350°C, the reaction pressure is 1.0-4.0 MPa, and the reaction time is 1-24 hours; The oxidation passivation treatment conditions in step (2) are: the reaction atmosphere is oxygen-containing nitrogen with an oxygen volume content of 0.5-2%, the reaction temperature is 150-300°C, the reaction pressure is 1.0-4.0 MPa, and the reaction time is 1-24 hours; The carbonization treatment conditions in step (3) are as follows: the reaction atmosphere is synthesis gas with a hydrogen-to-carbon ratio of 10-100:1, the reactor pressure is 1.0-4.0 MPa, the reaction temperature is 300-500°C, and the reaction time is 1-36 hours.

7. The activation method according to claim 6, characterized in that During the reduction treatment in step (1), the H2 air velocity is 600-6000Nm 3 / h / t, during the reduction process, the fluidized bed reactor adopts a bubbling state, and the superficial velocity of the gas at the inlet is 0.1-0.5m / s; During the oxidation passivation treatment in step (2), the oxygen space velocity is 100-800Nm 3 / h / t, the oxidation reactor adopts bubbling reaction, and the superficial gas velocity is 0.1-0.4m / s; During the carbonization treatment in step (3), the synthesis gas space velocity is 500-5000Nm 3 / h / t, during the carbonization process, the bed of the fluidized bed reactor adopts a turbulent state, and the superficial velocity of the gas at the inlet is 0.2-0.7m / s.

8. The activation method according to claim 6, characterized in that During the reduction treatment in step (1), the H2 air velocity is 1000-4000Nm 3 / h / t, during the reduction process, the fluidized bed reactor adopts a bubbling state, and the superficial velocity of the gas at the inlet is 0.15-0.3m / s; During the oxidation passivation treatment in step (2), the oxygen space velocity is 200-500Nm 3 / h / t, the oxidation reactor adopts bubbling reaction, and the superficial gas velocity is 0.15-0.3m / s; During the carbonization treatment in step (3), the synthesis gas space velocity is 1000-3000Nm 3 / h / t, during the carbonization process, the bed of the fluidized bed reactor adopts a turbulent state, and the superficial velocity of the gas at the inlet is 0.3-0.6m / s.

9. The activation method according to any one of claims 6 to 8, characterized in that When heating in step (1), the temperature is raised to 100-150°C at a rate of 15-25°C / h and maintained for 1-2 hours, and then the temperature is continued to be raised at a rate of 10-20°C / h to reach the reaction temperature of the reduction treatment; In steps (2) and (3), when heating, the heating rate is 10-20°C / h to reach the reaction temperature.

10. The activation method according to claim 9, characterized in that When the catalyst is transferred between different devices, the catalyst is delivered to the target device by nitrogen delivery using the pressure difference.

11. The activation method according to claim 10, characterized in that The catalyst in the catalyst feeding tank is maintained at 40-60° C. under N 2 conditions, and then the catalyst is added to the first gas-solid fluidized bed reactor through N 2 gas transportation using pressure difference.

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