Industrial activation system and activation method for forming pure-phase chi iron carbide catalyst

Abstract

The invention discloses an industrial activation system and an activation method for forming a pure-phase chi iron carbide catalyst. The industrial activation system 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 intermittent operation mode is adopted, the reduction reaction and carbonization process of the catalyst are achieved in one reactor, the oxidation and passivation process is completed on another independent gas-solid fluidized bed, 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 to be carbonized after passivation is completed. The pure-phase chi iron carbide is suitable for industrial production.

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 an activation method for forming a pure phase x-iron carbide catalyst. Background Art

[0002] my country's energy structure is characterized by rich coal, poor oil and little gas. Therefore, the conversion of coal into high-quality liquid fuels and high-value-added chemicals has always received great attention. Indirect coal liquefaction is a process in which coal is used as raw material, and after gasification and purification processes, it becomes clean synthesis gas, and then through the Fischer-Tropsch synthesis reaction, it becomes large molecular hydrocarbons and alcohols. It is an important way to achieve the rational, efficient and clean use of coal.

[0003] The Fischer-Tropsch synthesis reaction is the core part 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 main active phase of the iron-based Fischer-Tropsch synthesis catalyst is iron carbide. Usually, the prepared Fischer-Tropsch synthesis iron-based catalysts are prepared in an oxidized state. Before the Fischer-Tropsch synthesis reaction, it is necessary to use H2, carbon monoxide, low-carbon olefins and other reducing gases for activation treatment to make the catalyst form a stable iron carbide active phase. Therefore, the activation process has a great influence on the reaction performance of the Fischer-Tropsch synthesis iron-based catalyst.

[0004] However, since there are many types and structural forms of iron carbide, their activities 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 types of iron carbides in the catalyst, and it is not a pure phase structure. CN103551207A and CN100404137C etc. propose to use gas-solid fluidized bed reactor to activate precipitated iron catalyst, and the catalyst after activation is also mixed phase; CN 107149948A discloses a gas-solid fluidized bed Fischer-Tropsch synthesis catalyst activation method, in order to control the water content in the reactor, a dehydration tank is set in front of the reactor, and the activation temperature rise process is divided into different stages, mainly at 120°C constant temperature for 2-8 hours, and in order to ensure the stability of the reaction, at 220-230°C constant temperature for 3-10h; reaction pressure 0.5-2MPa, empty tower gas velocity 0.04-0.12m / s, the final constant temperature is 260-280°C. The reducing gas is H2 and CO, the molar ratio is 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 the 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 hydrocarbon products such as oil wax.

[0006] CN112569992A and CN112569975A disclose methods for preparing precipitated χ iron carbide and a mixture of χ iron carbide and other types of phase-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 disclose methods for preparing supported χ iron carbide, and CN112569975A discloses a method for first forming pure phase χ iron carbide, pure phase θ iron carbide and pure phase ε / ε ’ Iron carbide is then mixed with a 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, but the overall process is complicated. 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 the pure phase χ iron carbide is prepared under synthesis gas conditions.

[0008] In the patents published in CN112569992A and CN112569975A, the process of forming precipitated pure phase x iron carbide is embedded in the entire catalyst preparation process and is not suitable for large-scale continuous production. The catalyst produced on a large scale is in an oxidized state when it leaves the factory. It is generally transported to the reaction device site, and the iron oxide is activated into an active iron phase with Fischer-Tropsch synthesis reaction ability by an independent activation device, and then moved to the Fischer-Tropsch synthesis reactor to participate in the reaction.

[0009] Although the methods for preparing pure phase iron carbide precursors in the above 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, and then oxygen passivated, and finally 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 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 perform N2 replacement operation to ensure that oxygen does not contact H2 and CO to avoid danger. There are problems in 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, so 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 fixed fluidized bed or gas-solid bubbling bed Fischer-Tropsch catalyst reduction activation system and process, and is provided with a cyclone separator to recover the catalyst entrained in the gas, a catalyst metering scale to control the mass of the catalyst entering the Fischer-Tropsch reactor, and a gas mixer to adjust the reducing gas atmosphere. This technology is suitable for the activation of general catalysts, that is, forming mixed phase active components, and does not point out 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, which divides the reduction process into a reduction stage and an adjustment stage according to different reaction heat stages, wherein the reduction stage corresponds to a heating stage, and the adjustment stage corresponds to a reaction exothermic stage. The catalyst reduction temperature of this technology is 260-450°C, the pressure is 2.0-5.0MPa, the inlet gas velocity is 0.15-0.7m / s, etc., and the reduction operation is carried out in an atmosphere of H2, CO or a high hydrogen-carbon ratio (H2+CO). This technology is also not aimed at the formation of pure phase iron carbide, nor does it point out the key points and technical methods for forming pure phase iron carbide.

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

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

[0015] The iron-based catalysts for Fischer-Tropsch synthesis produced on a large scale are generally in an oxidized state when leaving the factory and have no catalytic activity. They need to undergo a separate activation treatment to make them have catalytic performance for Fischer-Tropsch synthesis.

[0016] At present, the activation process of iron-based catalysts is a general method of converting iron oxide into a mixture of various active phases such as ferroferric oxide, elemental iron, and multiphase iron carbide. There is no activation process suitable for large-scale operation to form pure phase iron carbide as the main active component of the catalyst. Summary of the invention

[0017] In view of this, the object of the present invention is to provide an industrial activation system and activation method for forming a pure phase χ iron carbide catalyst, so as to solve the problem of industrial large-scale preparation of pure phase χ iron carbide catalyst.

[0018] To achieve one aspect of the above-mentioned invention object, the present invention adopts the following technical solution:

[0019] An industrial activation system for forming a pure phase iron carbide catalyst, the industrial activation system comprises 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,

[0020] 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 an oxidized catalyst to be activated;

[0021] The first gas-solid fluidized bed reactor is equipped with a first air inlet pipe at the bottom to supply the gas required in the first gas-solid fluidized bed reactor, and a first tail gas pipe at the top to discharge 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 treated by reduction in the first gas-solid fluidized bed reactor to the second gas-solid fluidized bed reactor, and to deliver the catalyst treated by the oxygen-containing atmosphere in the second gas-solid fluidized bed reactor back 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, used 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, used to cool 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 invention, preferably, the heat exchanger is used to perform heat exchange and temperature reduction on the tail gas transported through the first tail gas pipe and the intake gas transported through the first intake pipe to enter the first gas-solid fluidized bed reactor to recover heat.

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

[0034] a first heater, disposed on the first air intake pipe, for heating the intake air from the heat exchanger;

[0035] a second heater, disposed on the second air inlet pipe, for heating the air to be introduced into the second gas-solid fluidized bed reactor; and

[0036] A circulation 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.

[0037] In order to achieve another aspect of the above-mentioned invention object, the present invention also provides an activation method for forming a pure phase x iron carbide catalyst using the above-mentioned industrial activation system, the activation method comprising:

[0038] (1) In a first gas-solid fluidized bed reactor, reducing the oxidized catalyst using hydrogen;

[0039] (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;

[0040] (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.

[0041] According to the activation method of the present invention, preferably, 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 1.0-4.0Mpa, and the reaction time is 1-24h.

[0042] According to the activation method of the present invention, preferably, 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 1.0-4.0Mpa, and the reaction time is 1-24h.

[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 oxidation passivation treatment in step (2), the oxygen space velocity is 100-800Nm 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 gas velocity of the gas at the inlet is 0.2-0.7m / s, preferably 0.3-0.6m / 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 delivery using the pressure difference; the catalyst storage tank is a 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 advanced pure phase iron carbide catalysts to be used in Fischer-Tropsch synthesis devices; based on the main parameter characteristics of forming 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 on 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 returns to the first gas-solid fluidized bed reactor of the fluidized bed for carbonization to form precipitated pure phase χ iron carbide, and the pure phase iron carbide 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 the 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 an 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 in conjunction with the embodiments and drawings, but the present invention is not limited to the listed embodiments, but also includes equivalent improvements and modifications 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 invention is an iron-based Fischer-Tropsch synthesis catalyst in which iron is in an oxidized state. For example, the catalyst precursor prepared in CN112569982A, CN112569993A or CN112569987A can be used, in which the Fe component is in an oxidized state; the average particle size of the finished oxidized catalyst particles used as the raw material is preferably 70-80 microns, which belongs to the Geldart A type particle range in the gas-solid fluidization field; preferably, the catalyst attrition of the raw material is less than 4g / h.

[0060] In one embodiment, Figure 1 As shown, the industrial activation device of the present invention comprises 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 and is used to supply the first fluidized bed reactor with the oxidized catalyst to be activated as a catalyst raw material; the specific conveying method is well known in the art, such as screw conveying or pneumatic conveying.

[0062] The bottom of the first gas-solid fluidized bed reactor 3 is equipped with a first air 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 tail gas after the reaction.

[0063] A second air inlet pipe 15 is provided at the bottom of the second gas-solid fluidized bed reactor 10 to supply the required gas to the second gas-solid fluidized bed reactor to fluidize the material in the reactor, and a second tail gas pipe 16 is provided at the top to discharge the gas in the second gas-solid fluidized bed reactor as tail gas after the 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, which is used to send the catalyst after reduction treatment in the first gas-solid fluidized bed reactor into the second gas-solid fluidized bed reactor and send the catalyst after treatment in the oxygen-containing atmosphere 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 catalyst can be smoothly transported 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 also includes a tail gas separation unit and a tail gas circulation pipe 14, wherein the tail gas separation unit is used to separate 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 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 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 transported through the first tail gas pipe 32 to recover heat;

[0069] The first gas-liquid separation tank 5 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;

[0070] A cooler 6 is 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 can be understood 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 heat and cool the tail gas transported through the first tail gas pipe 32 and 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, disposed 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, disposed on the second air inlet pipe, for heating the air to be introduced into the second gas-solid fluidized bed reactor, so as to adjust the temperature in the second gas-solid fluidized bed; and

[0078] The circulating compressor 9 is arranged 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 intake pipe.

[0079] During operation, the industrial activation system of the present invention first completes the gas circulation operation under the condition of N2 and stabilizes it. After N2 is introduced into the system and the replacement is completed, the circulation compressor 9 is started, and the fresh N2 is mixed with the circulation gas from the circulation compressor 9 to form the inlet gas, and then passes through the heat exchanger 1 and the first heater 2 in sequence, and then enters the first gas-solid fluidized bed reactor 3. The gas out of the reactor is used as the circulation gas, and 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 in sequence, and then enters the circulation compressor 9, and is mixed with the fresh gas N2 to form the inlet gas, and the cycle is repeated. When the system pressure reaches the set value, the gas circulation is performed to keep the system pressure stable. In this cycle, the pressure range in the first gas-solid fluidized bed reactor is 1.0-4.0MPa, such as 2, 2.5, 3 or 3.5MPa.

[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] Under the hydrogen atmosphere of H2 or a mixture of H2+N2, the temperature of the first gas-solid fluidized bed reactor is increased, for example, the temperature is increased to 100-150°C, such as 110, 120, 130 or 140°C at 15-25°C 20°C / h and maintained for 1-2 hours, and then the temperature is continued to be increased at a rate of 10-20°C / h. When the temperature reaches the set temperature, it is treated for 1-24 hours, such as 4, 8, 12, 16 or 20 hours. The reactor temperature is in the range of 240-350°C, such as 260, 280, 300 or 330°C.

[0083] After the reduction is completed, the catalyst is transferred to a separate second gas-solid fluidized bed reactor through a catalyst delivery pipe by nitrogen delivery for oxidation passivation treatment. Nitrogen is first introduced for replacement to ensure that the H2 content 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 maintained for 1-24 hours, such as 4, 8, 12, 16 or 20 hours. The pressure range is 1.0-4.0MPa, such as 2, 2.5, 3 or 3.5MPa to complete the oxidation passivation treatment.

[0084] After the oxidation passivation is completed, the second gas-solid fluidized bed reactor is replaced with nitrogen and returned to the first gas-solid fluidized bed reactor by nitrogen transportation for carbonization treatment.

[0085] The fresh gas in the carbonization process is H2 and carbon monoxide. The synthesis gas mixed with the circulating gas is used as the inlet gas, which passes through the first heat exchanger and the heater in sequence and enters the fluidized bed reactor. The gas out of the reactor is used as the tail gas, which passes through the heat exchanger, the first gas-liquid separator, the water cooler, the second gas-liquid separator and the dehydration adsorption tank in sequence, and enters the circulation compressor as the circulating gas, which is mixed with the fresh gas (H2+CO) as the inlet gas, and the cycle is repeated. The reactor pressure in the carbonization process is 1.0-4.0MPa, such as 2, 2.5, 3 or 3.5MPa, and the reactor temperature is increased to 300-500℃, such as 320, 350, 380, 400, 450 or 480℃ at a rate of 10-20℃ / h, 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 H2 space velocity as fresh gas feed is controlled to be 600-6000Nm according to the amount of catalyst. 3 / h / t, preferably 1000-4000Nm 3 / h / tFor example 2000 or 3000Nm 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 as fresh gas feed is mixed with oxygen, in which the proportion of oxygen is 0.5-2%, such as 1% or 1.5%, and the oxygen space velocity is controlled at 100-800Nm according to the amount of catalyst. 3 / h / t, preferably 200-500Nm 3 / h / tFor example 300 or 400Nm 3 / h / t; preferably, the second gas-solid fluidized bed reactor adopts a bubbling morphology 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, and the synthesis gas space velocity is controlled to 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 gas 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 required, the catalyst can be continuously 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 in conjunction with 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 replacing with H2 atmosphere, 100 kg of catalyst was added to the first gas-solid fluidized bed reactor through the catalyst feeding tank, and the temperature was raised to 120°C at a rate of 20°C / h and kept at a constant temperature for 1 hour. The temperature was continued to be raised at a rate of 10°C / h and H2 was introduced. The hydrogen flow rate was 200 Nm 3 / h, and after the temperature reaches 300°C, it is kept warm for 12 hours to complete the catalyst reduction. During the reduction process, the reactor pressure is maintained at 2.0 MPa and the inlet superficial gas velocity is 0.2 m / s.

[0094] After the reduction treatment 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 to the first gas-solid fluidized bed reactor is stopped, the hydrogen atmosphere circulation is maintained, and the reactor temperature is reduced to the range of 150-200°C.

[0095] 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 0.6% oxygen is introduced, with an oxygen flow rate of 30Nm 3 / h, reactor pressure 2.0-2.2MPa, reaction temperature 180-200°C, after 12 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%.

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

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

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

[0099] The catalyst in the catalyst storage tank is taken for X-ray diffractometer detection, and the obtained product is a monoclinic χ-iron carbide phase; the generated target product χ-iron carbide has good crystallinity, completely corresponds to all characteristic peaks of χ-iron carbide, and is free of impurities. The obtained iron carbide is a pure phase χ-iron carbide catalyst; at the same time, the loss of the catalyst in the reduction, passivation and carbonization processes is very small, and the catalyst yield calculated based on the iron element is 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 to 1.5 MPa. 100 kg of catalyst was added to the reactor through the catalyst feeding tank, and the H2 feed amount 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 in a 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 feeding H2, maintains the circulation of hydrogen atmosphere, 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 less than 1%. At the same time, the temperature of the second gas-solid fluidized bed reactor is 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 complete the passivation operation after the oxygen content of the oxidation passivation system is lower than 0.1%.

[0104] Before the catalyst is oxidized and passivated, the pressure of the fluidized bed reactor used for carbonization is maintained at 1.0-1.2MP. After the passivation reaction is completed, the catalyst is returned to the fluidized bed reactor by nitrogen delivery. Hydrogen is introduced to increase the system pressure to 1.5MPa, and the feed is switched to CO+H2, with a hydrogen-to-carbon ratio of 80 and a total CO+H2 flow rate of 200Nm 3 / h; at the same time, the reactor temperature is raised to 300℃ at a rate of 10℃ / h, and the constant temperature time is maintained for 12 hours to complete the carbonization of the catalyst and form a precipitated pure phase precipitated χ iron carbide. During the heating and constant temperature process, the flow rate of the circulating compressor is adjusted to maintain the apparent gas velocity of the inlet gas of the fluidized bed reactor in the range of 0.4-0.45m / s during the carbonization process. After the carbonization reaction is completed, the temperature is lowered to 120℃, the CO+H2 feed is stopped, and the hydrogen circulation is maintained.

[0105] During the carbonization reaction, the catalyst storage tank is replaced with N2 atmosphere, and the pressure is controlled to 0.7MPa. After the carbonization reaction is completed, it is transferred to the catalyst finished product tank, and the catalyst in the catalyst storage tank is taken for X-ray diffractometer detection. The obtained product is a monoclinic χ-iron carbide phase; the generated target product χ-iron carbide has good crystallinity, corresponds well to all characteristic peaks of χ-iron carbide, has no impurities, and the obtained iron carbide is a pure phase χ-iron carbide catalyst. The loss of the catalyst reduction, passivation and carbonization process is very small, and the catalyst yield calculated based on the iron element is 99.8%.

[0106] Example 3

[0107] Similar to Example 1, the scale of the reactor and the catalyst processing capacity are further improved. The fluidized bed reactor used for reduction and carbonization has a diameter of 1.5m and a height of 20m, and a catalyst processing capacity of 7000-10000kg / batch. The reactor used for oxidation passivation has a diameter of 1.8m and a height of 12m.

[0108] First, the atmosphere of the reduction reaction system was replaced with H2 atmosphere, and the pressure was controlled to 3.0 MPa. 8000 kg of catalyst was added to the reactor through the catalyst feeding tank, and the H2 feed amount was maintained at 20000 Nm 3 / h, reactor temperature 350℃, superficial gas velocity 0.2m / s, constant temperature reaction 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 to 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 2.0-2.2MP. After the passivation reaction is completed, the catalyst is returned to the fluidized bed reactor by nitrogen delivery. Hydrogen is introduced to increase the system pressure to 3.0MPa, and the feed is switched to CO+H2, with a hydrogen-to-carbon ratio of 50 and a total CO+H2 flow rate of 24000Nm 3 / h; at the same time, the reactor temperature is raised to 380℃ at a heating rate of 10℃ / h, and the constant temperature time is maintained for 12 hours to complete the catalyst carbonization and form a precipitated pure phase precipitated χ iron carbide. During the heating and constant temperature process, the flow rate of the circulating compressor is adjusted to maintain the apparent gas velocity of the inlet gas of the fluidized bed reactor in the range of 0.4-0.5m / s during the carbonization process. After the carbonization reaction is completed, it is reduced to 120℃, CO+H2 feeding is stopped, hydrogen circulation is maintained, and the pressure is reduced to 1.5MPa. The circulating gas in the carbonization process does not pass through the dehydration adsorption tank.

[0112] During the carbonization reaction, the catalyst storage tank is replaced with N2 atmosphere, and the pressure is controlled to be 1.0MPa. After the carbonization reaction is completed, it is transferred to the catalyst finished product tank, and the catalyst in the catalyst storage tank is taken for X-ray diffractometer detection. The obtained product is a monoclinic χ-iron carbide phase; the generated target product χ-iron carbide has good crystallinity, which corresponds well to all characteristic peaks of χ-iron carbide, and has no other impurities. The obtained iron carbide is a pure phase χ-iron carbide catalyst. The loss of the catalyst reduction, passivation and carbonization process is very small, and the catalyst yield calculated based on the iron element is 99.6%.

[0113] Comparative Example 1

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

[0115] The catalyst in the catalyst storage tank was tested by X-ray diffractometer. The obtained product was mostly monoclinic χ-iron carbide phase, but there were about 5% other crystal phases. The generated target product χ-iron carbide had good crystallinity, well corresponding to all characteristic peaks of χ-iron carbide, without any other impurities. The obtained iron carbide was a pure phase χ-iron carbide catalyst. The catalyst yield calculated based on iron element was 97.4%.

[0116] However, compared with Example 1, the overall operation process requires cooling and reducing pressure first, and then heating and increasing pressure after the sudden change. The operation process is complicated, the overall operation cycle is significantly increased, and the catalyst yield is lower than that of the embodiment, which is not conducive to large-scale production operations.

Claims

1. An industrial activation system for forming a pure phase x iron carbide catalyst, characterized in that The industrial activation system 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; 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 an oxidized catalyst to be activated; The first gas-solid fluidized bed reactor is equipped with a first air inlet pipe at the bottom to supply the gas required in the first gas-solid fluidized bed reactor, and a first tail gas pipe at the top to discharge 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 by reduction in the first gas-solid fluidized bed reactor to the second gas-solid fluidized bed reactor, and to deliver the catalyst treated by the oxygen-containing atmosphere in the second gas-solid fluidized bed reactor back to the first gas-solid fluidized bed reactor for treatment 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 x-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 contained in 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, used for exchanging heat and cooling 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, used to cool 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 perform heat exchange and temperature reduction on the tail gas transported through the first tail gas pipe and the intake gas 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, disposed on the second air inlet pipe, for heating the air to be introduced into the second gas-solid fluidized bed reactor; and A circulation 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. An activation method for forming a pure phase x iron carbide catalyst using an industrial activation system according to any one of claims 1 to 4, the activation method comprising: (1) In a first gas-solid fluidized bed reactor, reducing the oxidized catalyst 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 1.0-4.0Mpa, and the reaction time is 1-24h; 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 h; 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.

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, 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; During the oxidation passivation treatment in step (2), the oxygen space velocity is 100-800Nm 3 / h / t, preferably 200-500Nm 3 / h / t, the oxidation reactor adopts bubbling reaction, the superficial gas velocity is 0.1-0.4m / s, preferably 0.15-0.3m / s; 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 gas velocity of the gas at the inlet is 0.2-0.7m / s, preferably 0.3-0.6m / s.

8. The activation method according to claim 6 or 7, 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.

9. The activation method according to claim 8, 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.

10. The activation method according to claim 9, characterized in that: 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 delivery using pressure difference.

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