Microchannel reactor and method of making sustainable aviation fuel

Abstract

The application provides a kind of microchannel reactor and the method for preparing sustainable aviation fuel, the microchannel reactor includes the reaction channel layer and the heat exchange channel layer of alternative superposition, the reaction channel layer includes the microchannel array of parallel, the heat exchange channel layer includes heat exchange medium;The reaction channel layer includes feed inlet and discharge outlet;The application is realized by the integrated design of the reaction channel layer and the heat exchange channel layer of alternative superposition, realizes the efficient heat transfer mass transfer and precision isothermal control of mixed dehydrogenation and C9~C16 aromatic hydrocarbon hydrogenation process, significantly improves C2~C6 olefin selectivity and naphthene purity, the content of prepared SAF naphthene component is greatly improved, fully meets the performance requirements of ASTM D7566 standard to SAF naphthene / aromatic component, can realize more than 50% mixing ratio even 100% pure bio-based application, solves the core bottleneck of traditional preparation SAF naphthene content deficiency.

Description

Technical Field

[0001] This invention relates to the field of aviation fuel technology, and in particular to a microchannel reactor and a method for preparing sustainable aviation fuel. Background Technology

[0002] Sustainable aviation fuel (SAF), as a core alternative fuel for reducing carbon emissions in the aviation industry, boasts significant advantages in its production process, including a synergistic aromatization and hydrodeoxygenation route using naphtha and liquefied petroleum gas. This route offers advantages such as a wide range of raw material sources, high aromatic content, and the ability to achieve 100% pure combustion. The core steps of this route include the mixed dehydrogenation of light alkanes to produce C3-C4 olefins and the hydrogenation saturation of C9-C16 aromatics to produce cycloalkanes. Both are strongly endothermic or exothermic reactions, requiring extremely high precision in temperature control, mass and heat transfer efficiency, and catalyst stability.

[0003] In existing technologies, the above two-step reaction typically employs a traditional fixed-bed reactor. Fixed-bed reactors suffer from the following significant technical drawbacks: First, they have low heat and mass transfer efficiency; the mixed dehydrogenation section (strongly endothermic) is prone to cold spots leading to decreased conversion, while the C9-C16 aromatic hydrogenation section (strongly exothermic) is prone to hot spots causing coking, over-cracking, and rapid catalyst deactivation. Second, they have a wide residence time distribution, increasing side reactions; the selectivity for C3-C4 olefins is typically only 75%-85%, and the selectivity for cycloalkanes in aromatic hydrogenation is difficult to stabilize above 90%. Third, they have a large pressure drop (typically 0.2-0.5 MPa), limiting space velocity increases, resulting in bulky equipment and high energy consumption. Fourth, they are difficult to achieve efficient thermal coupling with the aromatization and hydrodeoxygenation units in the process, leading to low overall energy utilization.

[0004] The aforementioned defects have led to the mixed dehydrogenation and C9-C16 aromatic hydrogenation stages becoming bottlenecks in the entire naphtha and LPG co-aromatization coupled hydrodeoxygenation process: the low olefin yield directly affects the subsequent alkylation yield, and incomplete aromatic hydrogenation affects the purity and antioxidant stability of SAF cycloalkane components, ultimately limiting the production and quality improvement of sustainable aviation fuel and the failure to meet the octane number requirements, making it difficult to meet the stringent requirements for cycloalkane content in aviation fuel specifications (ASTM D7566).

[0005] Therefore, there is an urgent need to develop a new type of reactor that can achieve isothermal operation, high mass and heat transfer efficiency, low pressure drop, and modular scale-up, fundamentally solving the temperature control problem of mixed dehydrogenation and C9~C16 aromatic hydrogenation steps, achieving seamless thermal integration with upstream aromatization and downstream separation units, and providing key equipment support for the industrial application of naphtha and liquefied petroleum gas synergistic aromatization coupled hydrodeoxygenation process. Summary of the Invention

[0006] The technical problem to be solved by this invention is that existing traditional reactors have low heat and mass transfer efficiency, large pressure drop, and low overall energy utilization.

[0007] To address the aforementioned technical problems, in a first aspect, the present invention provides a microchannel reactor, the microchannel reactor comprising alternating stacked reaction channel layers and heat exchange channel layers, the reaction channel layers comprising a parallel array of microchannels, the heat exchange channel layers comprising a heat exchange medium; the reaction channel layers comprising an inlet and an outlet; The channel diameter of the microchannel array is 100μm to 2mm, for example, it can be 100μm, 1mm, 1.5mm, 2mm, etc.; the number of channels in the microchannel array is 80 to 10000, for example, it can be 80, 1000, 5000 or 10000, etc.

[0008] The inner wall of the microchannel array includes a porous support layer and a catalyst layer.

[0009] The microchannel reactor provided by this invention is suitable for the co-aromatization of bio-based naphtha and liquefied petroleum gas to produce sustainable aviation fuel. In this invention, the reaction channel layer is a parallel array of microchannels formed by precision etching, machining, or 3D printing of metal or ceramic plates. The design of hundreds to thousands of parallel channels achieves uniform material distribution, avoiding reaction imbalances caused by excessively high or low local material concentrations. Simultaneously, the parallel connection of multiple channels increases the reaction throughput, adapting to the needs of industrial-scale production. Compared with traditional fixed-bed reactors, the microchannel structure of this invention significantly increases the contact area between the material and the catalyst, accelerating the reaction rate.

[0010] The inner wall of the microchannel array comprises a porous support layer and a catalyst layer. The porous support layer enhances hydrothermal stability and wear resistance, and also regulates the acidity distribution on the support surface, reducing strong acid sites and preventing side reactions. The catalyst layer forms a structured wall, effectively improving the catalytic effect.

[0011] The heat exchange channel layer is closely adjacent to and alternately arranged with the reaction channel layer, and is arranged in a counter-current or cross-flow manner. Oil, molten salt, steam or water flows in the channel as the heat exchange medium to achieve isothermal operation in the reaction channel layer.

[0012] The reaction channel layer includes an inlet and an outlet. The inlet is used to introduce the reactants (mixed dehydrogenation or hydrogenation materials), and the outlet is used to discharge the post-reaction stream. The structural design of the inlet and outlet is adapted to the distribution of the microchannel array, ensuring that the materials enter each microchannel quickly and uniformly, while avoiding the retention of the post-reaction stream, reducing side reactions, and improving product purity. The microchannel reactor of this invention is connected to the upstream aromatization reactor, the hydrodeoxygenation reactor, and the downstream separation unit via thermally coupled pipelines.

[0013] The porous support layer includes Al2O3 or SiO2, and the thickness of the porous support layer is 5~50μm, for example, it can be 5μm, 10μm, 20μm, 30μm, 40μm or 50μm, etc.

[0014] The catalyst layer comprises any one or a combination of at least two of the following: Pt-Sn supported Al2O3, Pt-Sn-Zn supported Al2O3, Ni-Mo supported Al2O3, or Pd-based catalyst.

[0015] Preferably, the heat exchange medium includes any one of oil, molten salt, steam, or water.

[0016] In a second aspect, the present invention provides a method for mixed dehydrogenation in the preparation of sustainable aviation fuel, the method comprising: feeding mixed dehydrogenation material through the feed port of the microchannel reactor described in the first aspect, discharging it from the discharge port after the mixed dehydrogenation reaction, thereby completing the mixed dehydrogenation.

[0017] Preferably, the mixed dehydrogenation material comprises a C1-C6 light alkane component and a hydrogen component.

[0018] The total conversion rate of the C1-C6 light alkane components is 38%-42%, for example, it can be 38%, 39%, 40%, 41% or 42%, etc.

[0019] The temperature of the mixed dehydrogenation reaction is 480~520℃, and the pressure is 0.1~0.4MPa.

[0020] The core of this invention lies in utilizing the highly efficient heat transfer characteristics of a microchannel reactor to control reaction temperature fluctuations within a very small range, avoiding side reactions of light alkane cracking caused by high temperatures. At the same time, the efficient mass transfer characteristics within the microchannels ensure sufficient contact between the light alkane components and the catalyst, significantly improving the dehydrogenation reaction rate and conversion rate. Furthermore, the targeted design of the catalyst layer further enhances the selectivity of the dehydrogenation reaction, reduces the generation of by-products, and lowers the cost of subsequent separation and purification, providing high-quality dehydrogenation feedstock for the subsequent synthesis of sustainable aviation fuels.

[0021] Thirdly, the present invention provides a method for hydrogenation in the preparation of sustainable aviation fuel, the method comprising: feeding hydrogenation material through the feed inlet of the microchannel reactor described in the first aspect, and discharging it from the outlet after the hydrogenation reaction, thereby completing the hydrogenation process.

[0022] Preferably, the hydrogenated material comprises C9-C16 aromatic hydrocarbon components and hydrogen components.

[0023] The aromatic conversion rate of the C9-C16 aromatic components is 65%-85%, for example, it can be 65%, 70%, 75%, 80% or 85%, etc.

[0024] The hydrogenation reaction is carried out at a temperature of 320~380℃ and a pressure of 4~7MPa.

[0025] The hydrogenation materials include C9-C16 aromatic components and hydrogen components. The C9-C16 aromatic components are key intermediates in the synthesis of sustainable aviation fuels, and the thoroughness of their hydrogenation reaction directly affects the quality of aviation fuels.

[0026] The technical advantage of this method lies in the fact that the compact structure of the microchannel reactor can effectively withstand the high-pressure conditions of the hydrogenation reaction, avoiding the safety hazards of high-pressure leakage in traditional reactors; at the same time, the efficient heat and mass transfer characteristics enable the aromatic components to fully contact with hydrogen and catalyst, accelerating the hydrogenation reaction rate and improving the conversion rate of aromatics.

[0027] Fourthly, the present invention provides a method for operating the microchannel reactor described in the first aspect, the method comprising: before operating the microchannel reactor, purging the reactor with nitrogen until the oxygen volume content is no higher than 0.1%, then introducing hydrogen to reduce and activate the catalyst layer at atmospheric pressure and 200-300°C for 2-4 hours; after activation, gradually raising the temperature to the target reaction temperature, while introducing a heat exchange medium to start temperature control; after the temperature stabilizes, introducing hydrogen, and feeding the reactants from the inlet of the microchannel reactor; the reaction stream is discharged from the outlet of the microchannel reactor; during operation, adjusting the flow rate of the heat exchange medium in the heat exchange channel layer in real time through an online monitoring system to maintain the temperature fluctuation in the reaction channel layer to less than 2°C; after the reaction stabilizes, adjusting the flow rate of the heat exchange medium through an online monitoring system to maintain the temperature fluctuation in the reaction channel layer to less than 2°C; when the microchannel reactor is stopped, first cutting off the feed, purging and replacing with nitrogen and cooling to no higher than 100°C, and then regenerating the catalyst with air.

[0028] Implementing this invention has the following beneficial effects: This invention achieves efficient heat and mass transfer and precise isothermal control in the mixed dehydrogenation and C9-C16 aromatic hydrogenation processes through an integrated design of alternating reaction channel layers and heat exchange channel layers. It significantly improves the selectivity of C2-C6 olefins and the purity of cycloalkanes, resulting in a substantial increase in the cycloalkane content of the prepared SAF. This fully meets the performance requirements of ASTM D7566 standard for SAF cycloalkane / aromatic components, enabling blending ratios of over 50% and even 100% pure bio-based applications. This solves the core bottleneck of insufficient cycloalkane content in traditional SAF preparation methods.

[0029] This invention enhances the mixed dehydrogenation and aromatic hydrogenation steps in the synergistic aromatization process of naphtha and liquefied petroleum gas, achieving high-value conversion of bio-based naphtha byproducts and bio-based LPG, avoiding the waste of using low-value naphtha directly as fuel, and significantly improving feedstock utilization and overall yield.

[0030] This invention reduces pressure drop and facilitates thermal coupling with aromatization and hydrodeoxygenation units through structured wall catalysts and modular design, significantly reducing external steam and cooling water consumption, lowering overall energy consumption by more than 20%, extending catalyst life by more than 30%, and reducing operating costs.

[0031] This invention can be applied to the capacity expansion and retrofitting of existing HEFA (hydrodeoxygenation fuel exchange) units, requiring only the addition or replacement of mixed dehydrogenation and aromatics hydrogenation units to achieve process upgrades. It can also be used in the construction of new integrated biorefining and chemical processing (SAF) plants, offering flexible process layouts and promising prospects for industrial application. The entire route primarily uses bio-based feedstocks, resulting in significant reductions in carbon emissions throughout its lifecycle, thus offering both environmental and economic benefits. Attached Figure Description

[0032] Figure 1 The present invention is a microchannel reactor, wherein 1-the entire microchannel reactor, 2-reaction channel layer, 3-heat exchange channel layer, 4-microchannel array, 5-catalyst layer, 6-feed inlet, and 7-discharge outlet. Detailed Implementation

[0033] To make the objectives, technical solutions, and advantages of the embodiments of the present invention clearer, the technical solutions in the embodiments of the present invention will be clearly and completely described below. Obviously, the described embodiments are only some, not all, of the embodiments of the present invention. Based on the embodiments of the present invention, all other embodiments obtained by those skilled in the art without creative effort are within the scope of protection of the present invention.

[0034] The feed and discharge flow of the microchannel reactor of this invention is as follows, and the specific structural diagram is shown in the figure. Figure 1 As shown: The microchannel reactor 1 is composed of stacked reactor modules, including alternating reaction channel layers 2 and heat exchange channel layers 3; a parallel microchannel array 4 is formed within the reaction channel layers, and the inner walls of the channels are coated with a structured wall catalyst layer 5 (coated on top of the porous support layer). When performing the mixed dehydrogenation function: the C1~C6 light alkanes mixture from the separation unit and the circulating hydrogen enter from the upper reactant inlet 6 and undergo a dehydrogenation reaction under specific conditions; the post-reaction mixture stream (containing C2~C6 olefins, unconverted alkanes, a small amount of hydrogen and by-products) is discharged from the lower product outlet 7 and directly sent to the downstream alkylation unit.

[0035] When performing the C9~C16 aromatic hydrocarbon hydrogenation function: the C9~C16 aromatic hydrocarbon stream and hydrogen from the separation unit enter from the upper reactant inlet 6 and undergo hydrogenation reaction under specific conditions; the post-reaction stream (mainly C9~C16 cycloalkanes) is discharged from the lower product outlet 7 and sent to the final separation unit to obtain the SAF cycloalkane component.

[0036] Example 1

[0037] Preparation of sustainable aviation fuel using the microchannel reactor provided by this invention. (1) HEFA unit treatment: using catering waste oil as raw material, after dehydration, deacidification and demetallization pretreatment, it is sent to the HEFA reactor, using Ni-Mo supported Al2O3 catalyst, at a temperature of 350℃, a pressure of 5.0MPa and a liquid hourly space velocity of 2.0h. -1 Under certain conditions, a hydrodeoxygenation reaction is carried out, followed by isomerization cracking and distillation to obtain alkane-type SAF and bio-naphtha.

[0038] (2) Synergistic aromatization reaction: Bio-naphtha and bio-based LPG (biomass gasification separation product) are mixed and pretreated by desulfurization and dechlorination, and then fed into the aromatization reactor. Zn-supported ZSM-5 catalyst (Zn loading 3%) is used to react at a temperature of 480℃ and a pressure of 1.0MPa to separate C6~C9 aromatic components, C1~C6 light alkanes and by-product hydrogen. The purified hydrogen is then fed into the HEFA reactor.

[0039] (3) Dehydrogenation of light alkanes to olefins: C1~C6 light alkanes are fed into a microchannel reactor (mixed dehydrogenation function), and a Pt-Sn supported Al2O3 structured wall catalyst is used to carry out the reaction at a temperature of 500℃ and a normal pressure. The conversion rate of C1~C6 light alkanes is 40%, and the selectivity of C2~C6 monoolefins is 92%.

[0040] (4) Aromatic alkylation reaction: C6~C9 aromatics and C2~C6 monoolefins are mixed in a molar ratio of 5:1 and fed into an alkylation reactor. MCM-22 molecular sieve catalyst is used to react at a temperature of 200℃ and a pressure of 0.5MPa. C8~C16 alkyl aromatic fraction is obtained by distillation.

[0041] (5) SAF blending and molding: The alkane-type SAF is mixed with C8~C16 alkyl aromatic fractions and fed into a microchannel reactor (C9~C16 aromatic hydrogenation function). Using Ni-Mo supported Al2O3 structured wall catalyst, hydrogenation purification is carried out at a temperature of 350℃ and a pressure of 5.5MPa to obtain high aromatic SAF.

[0042] Testing revealed that the SAF prepared in this embodiment contained 32% aromatics and 48% cycloalkanes, with a distillation range of 180~280℃, fully complying with ASTM D7566 standards and can be used directly as 100% bio-based aviation fuel.

[0043] Example 2

[0044] Preparation of sustainable aviation fuel using the microchannel reactor provided by this invention. (1) HEFA unit processing: Animal and vegetable oils are used as raw materials. After dehydration, deacidification and demetallization pretreatment, they are fed into the HEFA reactor. Co-Mo supported Al2O3 catalyst is used at a temperature of 370℃, a pressure of 6.0MPa and a liquid hourly space velocity of 1.8h. -1 Under certain conditions, a hydrodeoxygenation reaction is carried out, followed by isomerization cracking and distillation to obtain alkane-type SAF and bio-naphtha.

[0045] (2) Synergistic aromatization reaction: Bio-naphtha and bio-based LPG (biogas purification product) are mixed and pretreated by desulfurization and dechlorination, and then fed into the aromatization reactor. Using a Ga-supported ZSM-5 catalyst (Ga loading 2.5%), the reaction is carried out at a temperature of 500℃ and a pressure of 1.2MPa to separate C6~C9 aromatic components, C1~C6 light alkanes and by-product hydrogen. The purified hydrogen is then fed into the HEFA reactor.

[0046] (3) Dehydrogenation of light alkanes to olefins: C1~C6 light alkanes are fed into a microchannel reactor (mixed dehydrogenation function), and a Pt-Sn-Zn supported Al2O3 structured wall catalyst is used to carry out the reaction at a temperature of 510℃ and 0.2MPa. The conversion rate of C1~C6 light alkanes is 41%, and the selectivity of C2~C6 monoolefins is 93.5%.

[0047] (4) Aromatic alkylation reaction: C6~C9 aromatics and C2~C6 monoolefins are mixed at a molar ratio of 4.5:1 and fed into an alkylation reactor. MCM-22 molecular sieve catalyst is used to react at a temperature of 210℃ and a pressure of 0.6MPa. C8~C16 alkyl aromatic fraction is obtained by distillation.

[0048] (5) SAF blending and molding: The alkane-type SAF is mixed with C8~C16 alkyl aromatic fractions and fed into a microchannel reactor (C9~C16 aromatic hydrogenation function). Using a Pd-based structured wall catalyst, hydrogenation purification is carried out at a temperature of 360℃ and a pressure of 6.0MPa to obtain high aromatic SAF.

[0049] Testing revealed that the SAF prepared in this embodiment contained 35% aromatics and 46% cycloalkanes, with a distillation range of 175~285℃, fully complying with ASTM D7566 standards and can be used directly as 100% bio-based aviation fuel.

[0050] Example 3

[0051] Preparation of sustainable aviation fuel using the microchannel reactor provided by this invention. (1) HEFA unit processing: Lignocellulosic oil is used as raw material. After dehydration, deacidification and demetallization pretreatment, it is fed into the HEFA reactor. Ni-Mo supported Al2O3 catalyst is used at a temperature of 380℃, a pressure of 6.5MPa and a liquid hourly space velocity of 1.6h. -1 Under certain conditions, a hydrodeoxygenation reaction is carried out, followed by isomerization cracking and distillation to obtain alkane-type SAF and bio-naphtha.

[0052] (2) Synergistic aromatization reaction: Bio-naphtha and bio-based LPG (biomass gasification separation product) are mixed and pretreated by desulfurization and dechlorination, and then fed into the aromatization reactor. Using a Ga-supported ZSM-5 catalyst (Ga loading 4%), the reaction is carried out at a temperature of 520℃ and a pressure of 1.5MPa to separate C6~C9 aromatic components, C1~C6 light alkanes and by-product hydrogen. The purified hydrogen is then fed into the HEFA reactor.

[0053] (3) Dehydrogenation of light alkanes to olefins: C1-C6 light alkanes are fed into the microchannel reactor (mixed dehydrogenation function) described in this invention, and Pt-Sn-Zn / Al2O3 structured wall catalyst is used to carry out the reaction at a temperature of 515℃ and 0.3MPa. The conversion rate of C1-C6 light alkanes is 39%, and the selectivity of C2-C6 monoolefins is 94%.

[0054] (4) Aromatic alkylation reaction: C6~C9 aromatics and C2~C6 monoolefins are mixed in a molar ratio of 4:1 and fed into an alkylation reactor. MCM-22 molecular sieve catalyst is used to react at a temperature of 220℃ and a pressure of 0.7MPa. C8~C16 alkyl aromatic fraction is obtained by distillation.

[0055] (5) SAF blending and molding: The alkane-type SAF is mixed with C8~C16 alkyl aromatic fractions and fed into a microchannel reactor (C9~C16 aromatic hydrogenation function). Using Ni-Mo / Al2O3 structured wall catalyst, hydrogenation purification is carried out at a temperature of 370℃ and a pressure of 6.5MPa to obtain high aromatic SAF.

[0056] Testing revealed that the SAF prepared in this embodiment contained 34% aromatics and 47% cycloalkanes, with a distillation range of 170~290℃, fully complying with ASTM D7566 standards and can be used directly as 100% bio-based aviation fuel.

[0057] Comparative Example 1 This comparative example employs a co-aromatization process using naphtha and liquefied petroleum gas, but both the mixed dehydrogenation and C9-C16 aromatic hydrogenation utilize conventional fixed-bed reactors. The specific steps are as follows: (1) HEFA unit processing: consistent with step (1) of Example 1, to obtain alkane-type SAF and bio-naphtha.

[0058] (2) Synergistic aromatization reaction: consistent with step (2) of Example 1, C6~C9 aromatic components, C1~C6 light alkanes and by-product hydrogen are separated.

[0059] (3) Dehydrogenation of light alkanes to olefins: C1~C6 light alkanes are fed into a conventional fixed-bed dehydrogenation reactor and reacted with Pt-Sn / Al2O3 catalyst at 500℃ and atmospheric pressure to produce C2~C6 monoolefins. The selectivity of monoolefins is only 81% (due to the increase of side reactions caused by the cold point).

[0060] (4) Aromatic alkylation reaction: consistent with step (4) of Example 1, to obtain C8~C16 alkyl aromatic fraction.

[0061] (5) SAF blending and molding: The alkane-type SAF is mixed with the C8~C16 alkyl aromatic fraction and fed into a conventional fixed-bed hydrogenation reactor. The mixture is then hydrogenated and refined using a Ni-Mo / Al2O3 catalyst at a temperature of 350℃ and a pressure of 5.5MPa to obtain SAF.

[0062] Testing revealed that the SAF prepared in this comparative example contained 30% aromatics and 44% cycloalkanes, with a distillation range of 182°C to 288°C. However, due to coking caused by hot spots in the fixed bed, the catalyst life was shortened by 45% compared to Example 1, the pressure drop was higher, the equipment volume was approximately 2.8 times larger than that of the microchannel reactor, the overall energy consumption increased by 25% compared to Example 1, and the raw material utilization rate decreased by 9%.

[0063] Comparative Example 2 In this comparative example, it is basically the same as in Example 1, except that the diameter of the microchannel array channel of the microchannel reactor is changed to 3 mm (exceeding the range of 100 μm to 2 mm). Other structures, catalyst layers, heat exchange channel layers and operating conditions are the same.

[0064] (1) HEFA unit processing: consistent with step (1) of Example 1, to obtain alkane-type SAF and bio-naphtha.

[0065] (2) Synergistic aromatization reaction: consistent with step (2) of Example 1, C6~C9 aromatic components, C1~C6 light alkanes and by-product hydrogen are separated.

[0066] (3) Dehydrogenation of light alkanes to olefins: C1~C6 light alkanes are fed into a modified reactor (mixed dehydrogenation function), and a Pt-Sn supported Al2O3 structured wall catalyst is used to carry out the reaction at a temperature of 500℃ and atmospheric pressure. The conversion rate of C1~C6 light alkanes is 36%, the selectivity of C2~C6 monoolefins is 85%, and the temperature fluctuation in the reaction channel reaches ±7℃ (due to the decrease in mass transfer efficiency caused by the increase in channel diameter).

[0067] (4) Aromatic alkylation reaction: consistent with step (4) of Example 1, to obtain C8~C16 alkyl aromatic fraction.

[0068] (5) SAF blending and molding: The alkane-type SAF is mixed with C8~C16 alkyl aromatic fractions and fed into the modified reactor (C9~C16 aromatic hydrogenation function). Using Ni-Mo supported Al2O3 structured wall catalyst, hydrogenation purification is carried out at a temperature of 350℃ and a pressure of 5.5MPa. The aromatic conversion rate is only 69%, and the final SAF has an aromatic content of 29% and a cycloalkanes content of 42%.

[0069] Testing revealed that although the SAF prepared in this comparative example basically met the requirements for distillation range distribution, the mass and heat transfer efficiency decreased due to the microchannel diameter exceeding the suitable range. The overall cycloalkane content decreased by 6 percentage points compared to Example 1, the catalyst life was shortened by about 30%, the pressure drop increased, and the energy consumption increased by about 15% compared to Example 1.

[0070] Comparative Example 3 In this comparative example, it is basically the same as in Example 1, except that the alternating stacked heat exchange channel layers are removed from the microchannel reactor, and only the reaction channel layer (including the parallel microchannel array and catalyst layer) is retained. The reaction temperature is controlled by an external electric heating furnace and a cooling jacket. Other structures and operating conditions are the same.

[0071] (1) HEFA unit processing: consistent with step (1) of Example 1, to obtain alkane-type SAF and bio-naphtha.

[0072] (2) Synergistic aromatization reaction: consistent with step (2) of Example 1, C6~C9 aromatic components, C1~C6 light alkanes and by-product hydrogen are separated.

[0073] (3) Dehydrogenation of light alkanes to olefins: C1~C6 light alkanes are fed into the modified reactor (mixed dehydrogenation function), and the reaction is carried out at 500℃ and atmospheric pressure using a Pt-Sn supported Al2O3 structured wall catalyst. The conversion rate of C1~C6 light alkanes is 37%, and the selectivity of C2~C6 monoolefins is 82%. The temperature fluctuation in the reaction channel reaches ±9℃ (due to the lack of integrated heat exchange layer, the isothermal control fails and obvious cold spots appear).

[0074] (4) Aromatic alkylation reaction: consistent with step (4) of Example 1, to obtain C8~C16 alkyl aromatic fraction.

[0075] (5) SAF blending and molding: The alkane-type SAF is mixed with C8~C16 alkyl aromatic fractions and fed into the modified reactor (C9~C16 aromatic hydrogenation function). Ni-Mo supported Al2O3 structured wall catalyst is used for hydrogenation purification at a temperature of 350℃ and a pressure of 5.5MPa. The aromatic conversion rate is 71%, and the final SAF has an aromatic content of 28% and a cycloalkanes content of 41%.

[0076] Testing revealed that the content of cycloalkane components in the SAF prepared in this comparative example was 7 percentage points lower than that in Example 1. Due to poor temperature control, side reactions increased, coking intensified, catalyst life was shortened by about 40%, the equipment size was larger, and the overall energy consumption increased by about 22% compared to Example 1. This fully demonstrates that the integrated design without alternating stacked heat exchange channel layers cannot achieve precise isothermal control.

[0077] Finally, it should be noted that the above embodiments are only used to illustrate the technical solutions of the present invention, and not to limit them; although the present invention has been described in detail with reference to the foregoing embodiments, those skilled in the art should understand that modifications can still be made to the technical solutions described in the foregoing embodiments, or equivalent substitutions can be made to some of the technical features; and these modifications or substitutions do not cause the essence of the corresponding technical solutions to deviate from the spirit and scope of the technical solutions of the embodiments of the present invention.

Claims

1. A microchannel reactor, characterized in that, The microchannel reactor comprises alternating layers of reaction channels and heat exchange channels. The reaction channel layer includes a parallel array of microchannels, and the heat exchange channel layer includes a heat exchange medium. The reaction channel layer includes an inlet and an outlet. The microchannel array has a channel diameter of 100 μm to 2 mm and a channel number of 80 to 10,000. The inner wall of the microchannel array includes a porous support layer and a catalyst layer.

2. The microchannel reactor according to claim 1, characterized in that, The porous support layer comprises Al2O3 or SiO2, and the thickness of the porous support layer is 5~50μm; The catalyst layer comprises any one or a combination of at least two of the following: Pt-Sn supported Al2O3, Pt-Sn-Zn supported Al2O3, Ni-Mo supported Al2O3, or Pd-based catalyst.

3. The microchannel reactor according to claim 1, characterized in that, The heat exchange medium includes any one of oil, molten salt, steam, or water.

4. A method for mixed dehydrogenation during the preparation of sustainable aviation fuel, characterized in that, The method includes: feeding the mixed dehydrogenation material through the feed port of the microchannel reactor according to any one of claims 1 to 3, and discharging it from the discharge port after the mixed dehydrogenation reaction, thereby completing the mixed dehydrogenation.

5. The method according to claim 4, characterized in that, The mixed dehydrogenation material includes C1-C6 light alkane components and hydrogen components; The total conversion rate of the C1-C6 light alkane components is 38%-42%; The temperature of the mixed dehydrogenation reaction is 480~520℃, and the pressure is 0.1~0.4MPa.

6. A method for hydrogenation in the preparation of sustainable aviation fuel, characterized in that, The method includes: feeding the hydrogenated material through the feed port of the microchannel reactor according to any one of claims 1 to 3, and discharging it from the discharge port after the hydrogenation reaction, thereby completing the hydrogenation process.

7. The method according to claim 6, characterized in that, The hydrogenated material includes C9-C16 aromatic hydrocarbon components and hydrogen components; The aromatic conversion rate of the C9-C16 aromatic components is 65%-85%; The hydrogenation reaction is carried out at a temperature of 320~380℃ and a pressure of 4~7MPa.

8. A method for operating a microchannel reactor according to any one of claims 1 to 3, characterized in that, The operating method includes: before operating the microchannel reactor, purging the reactor with nitrogen until the oxygen volume content is no higher than 0.1%, then introducing hydrogen to reduce and activate the catalyst layer at atmospheric pressure and 200-300℃ for 2-4 hours; after activation, gradually raising the temperature to the target reaction temperature while introducing a heat exchange medium to start temperature control; after the temperature stabilizes, introducing hydrogen and feeding the reactants from the inlet of the microchannel reactor; the reaction residue is discharged from the outlet of the microchannel reactor; during operation, the flow rate of the heat exchange medium in the heat exchange channel layer is adjusted in real time through an online monitoring system to maintain the temperature fluctuation in the reaction channel layer to less than 2℃; after the reaction stabilizes, the flow rate of the heat exchange medium is adjusted through an online monitoring system to maintain the temperature fluctuation in the reaction channel layer to less than 2℃; when the microchannel reactor stops operating, first cut off the feed, purge and replace with nitrogen and cool to no higher than 100℃, then regenerate the catalyst with air.

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