Catalyst regeneration

Abstract

The catalyst is regenerated by the method of the present invention, which uses a heat exchange fluid such as superheated steam to remove heat during the process, which relies on efficient heat transfer (e.g., made possible by microchannel reactor structures) compared to prior art heat exchanges that rely on phase changes, for example, between water and (partially or fully vaporized) steam, allowing for a simplified protocol and allowing for high temperature transitions between steps, which leads to a reduced duration of the regeneration process and avoids the potential risk of water hammer.

Description

[Technical field]

[0001] The present invention relates to a method for regenerating a catalyst, such as a Fischer-Tropsch (FT) catalyst. [Background technology]

[0002] The Fischer-Tropsch process is widely used to produce fuels from carbon monoxide and hydrogen and is represented by the formula: (2n+1)H 2 +nCO→C n H 2n+2 +nH 2 O It can be expressed as follows:

[0003] The reaction is highly exothermic and is catalyzed by a Fischer-Tropsch catalyst, typically a cobalt-based catalyst, under conditions of elevated temperature (typically at least 180° C., e.g. 200° C. or higher) and pressure (e.g. at least 10 bar). A product mixture is obtained, where n typically encompasses the range of 10 to 120. It is desirable to minimize the selectivity of light gases (e.g. methane) in the product mixture, i.e. the proportion of methane (n=1), and to maximize the selectivity to C5 and higher (n≧5) paraffins, typically to a level of 85% or greater. It is also desirable to maximize the conversion of carbon monoxide.

[0004] The hydrogen and carbon monoxide feedstock is typically synthesis gas.

[0005] Syngas can be produced by gasifying carbonaceous materials at elevated temperatures, for example, about 700° C. or higher. Carbonaceous materials can include any carbon-containing material that can be gasified to produce syngas. Carbonaceous materials can include biomass (e.g., plant or animal matter, biodegradable waste, etc.), food sources (e.g., corn, soybeans, etc.), and / or non-food sources such as coal (e.g., low rank coal, high rank coal, clean coal, etc.), oil (e.g., crude oil, heavy oil, tar sands oil, shale oil, etc.), solid waste (e.g., municipal solid waste, hazardous waste), refuse derived fuels (RDF), tires, petroleum coke, garbage, food waste, biogas, sewage sludge, animal waste, agricultural waste (e.g., corn stover, switchgrass, grass clippings), construction debris, plastic materials (e.g., plastic waste), cotton gin waste, mixtures of two or more thereof, and the like.

[0006] Alternatively, syngas can be produced by other means, such as by reforming natural or landfill gas, or gas produced by an anaerobic digestion process. Syngas can be produced by the conversion of CO2 to hydrogen using electrolysis as the hydrogen source. 2 It can also be produced by reforming (eg, the so-called "electricity-to-fuels" process).

[0007] The generated synthesis gas, as described above, may be converted to H by steam reforming (e.g., reacting methane with steam in the presence of a Fischer-Tropsch catalyst (a steam methane reforming (SMR) reaction); partial oxidation; autothermal reforming; carbon dioxide reforming; or a combination of two or more thereof) in preparation to provide fresh synthesis gas. 2 It can be treated to adjust the molar ratio to CO.

[0008] H in new syngas 2The molar ratio of CO is desirably in the range of about 1.6:1 to about 2.2:1, or about 1.8:1 to about 2.10:1, or about 1.95:1 to about 2.05:1.

[0009] The new synthetic gas is H 2 and CO. The tail gas can be optionally combined with recycled tail gas (e.g., recycled FT tail gas) that also contains H and CO to form the reactant mixture. 2 and CO, H 2 The molar ratio to CO optionally ranges from about 0.5:1 to about 2:1, or from about 0.6:1 to about 1.8:1, or from about 0.7:1 to about 1.2:1.

[0010] The combined FT syngas feed (consisting of fresh syngas combined with recycled tail gas) is preferably H 2 and CO in a molar ratio ranging from about 1.4:1 to about 2.1:1, or from about 1.7:1 to about 2.0:1, or from about 1.7:1 to about 1.9:1.

[0011] When recycled tail gas is used, the volumetric ratio of fresh syngas to recycled tail gas used to form the reactant mixture can be in the range of, for example, from about 1:1 to about 20:1, or from about 1:1 to about 10:1, or from about 1:1 to about 6:1, or from about 1:1 to about 4:1, or from about 3:2 to about 7:3, or about 2:1.

[0012] During the Fischer-Tropsch reaction, the catalyst gradually deactivates, its effectiveness decreases, and a gradual increase in temperature is required to maintain acceptable carbon monoxide conversion. This is described by Steynberg et al. "Fischer-Tropsch catalyst deactivation in commercial microchannel reactor operation" Catalysis Today 299 (2018) pp10-13.

[0013] Eventually, regenerating the catalyst becomes essential to restore its effectiveness. It is known to regenerate catalysts in situ.

[0014] Several different reactor types are known for carrying out Fischer-Tropsch synthesis, including fixed bed reactors, slurry bubble-column reactors (SBCRs) and microchannel reactors (Rytter et al, "Deactivation and Regeneration of Commercial Type Fischer-Tropsch Co-Catalysts - A Mini-Review" Catalysts 2015, 5, pp 478-499 at pp 482-483).

[0015] Microchannel reactors have been disclosed in WO 2016 / 201218(A) in the name of the applicant, which is incorporated by reference, and also in LeViness et al "Velocys Fischer-Tropsch Synthesis Technology - New Advances on State-of-the-Art" Top Catal 2014 57 pp518-525. Such reactors have the particular advantage that due to the high ratio of heat exchange surface area to microchannel (and therefore catalyst) volume, very efficient heat removal can be achieved.

[0016] However, as stated by Rytter et al (ibid.), p. 490, paragraph 2, lines 6 and 7: "Microchannel reactors present special challenges depending on the placement of the catalyst. In situ regeneration is optional, or the catalyst can be removed either by draining the catalyst particles or by removing the multichannel trays with the catalyst attached for external treatment."

[0017] The present invention relates to in situ catalyst regeneration in a microchannel reactor.

[0018] It is known, for example from Example 6 of WO 2016 / 201218(A), to regenerate a Fischer-Tropsch catalyst in a microchannel reactor by a three-stage process involving dewaxing with hydrogen at 350° C.-375° C., oxidation starting with air introduction by cooling to 70° C., and then reduction with hydrogen at about 350° C.

[0019] Heating and cooling is accomplished throughout the temperature range through the use of circulating cooling water as well as superheated steam. The transition from circulating cooling water to superheated steam, and vice versa, typically occurs in the range of 150° C. to 200° C., and if the correct procedures are not followed, there can be potential problems with the possibility of water hammering the reactor / steam drum, causing equipment damage resulting in downtime and repair costs.

[0020] In the above process, it is considered essential to cool the reactor to about 70° C. to avoid large exotherms from the reaction of the catalyst with oxygen and to eliminate the potential for reaction between hydrogen and oxygen (with inadequate purging between steps), but this limits the cooling and heating rates, thereby increasing the duration of the regeneration process. [Prior art documents] [Patent documents]

[0021] [Patent Document 1] International Publication WO2016 / 201218(A) Brochure [Non-patent literature]

[0022] [Non-Patent Document 1] Steynberg et al. “Fischer-Tropsch catalyst deactivation in commercial microchannel reactor operation” Catalysis Today 299 (2018) pp10-13 [Non-Patent Document 2] Rytter et al, “Deactivation and Regeneration of Commercial Type Fischer-Tropsch Co-Catalysts - A Mini-Review” Catalysts 2015, 5, pp 478-499 at pp 482-483 [Non-Patent Document 3] LeViness et al “Velocys Fischer-Tropsch Synthesis Technology - New Advances on State-of-the-Art” Top Catal 2014 57 pp518-525 [Non-Patent Document 4] Rytter et al (ibid), page 490 para 2 lines 6 and 7 Summary of the Invention [Problem to be solved by the invention]

[0023] It is an object of the present invention to overcome or mitigate the above-mentioned disadvantages of the prior art. [Means for solving the problem]

[0024] Thus, in a first aspect, the present invention provides a method for in situ regenerating a catalyst in a reactor equipped with heat exchange channels, preferably a microchannel reactor, comprising the steps of: a) dewaxing the catalyst by treating it at elevated temperature with a hydrogen-containing dewaxing gas stream flowing through the process microchannels of the reactor; b) oxidizing the resulting dewaxed catalyst by treating it with an oxidizing gas stream flowing through the process microchannels of the reactor at an elevated temperature; and c) reducing the resulting oxidized catalyst by treating it with a reducing gas stream flowing through the process microchannels of the reactor at an elevated temperature; in the transition from step a) to step b), reducing the temperature inside the process microchannels and / or the heat exchange channels from a temperature sufficient for dewaxing to a first lower limit of at least 90°C, preferably at least 100°C, more preferably between 140°C and 180°C, most preferably between 145°C and 155°C; In step b), increasing the temperature inside the process microchannels and / or the heat exchange channels to a temperature sufficient to oxidize the catalyst; in the transition from step b) to step c), reducing the temperature inside the process microchannels and / or the heat exchange channels from a temperature sufficient to cause oxidation to a first lower limit of at least 90°C, preferably at least 100°C, more preferably between 140°C and 180°C, most preferably between 145°C and 155°C; Then in step c), increasing the temperature inside the process microchannels and / or the heat exchange channels to a value sufficient to reduce the catalyst; wherein the temperature inside the process microchannels and / or the heat exchange channels is controlled by a heat exchange fluid flowing through the heat exchange channels of the microchannel reactor, and the entirety of the heat exchange fluid does not undergo a phase change.

[0025] In a preferred embodiment, the entire heat exchange fluid does not undergo a phase change in the method of the present invention. However, the method of the present invention can also be realized when the heat exchange fluid comprises multiple phases, only one of which does not undergo a phase change in the operation of the method of the present invention. For example, the heat exchange fluid may comprise only superheated steam, in which case no phase change occurs in the heat exchange fluid during the method of the present invention. This embodiment of the present invention is illustrated below in Example 5. As an alternative embodiment, the heat exchange fluid may comprise saturated steam (a mixture of steam and water), in which case only a portion of the heat exchange fluid (steam) does not undergo a phase change during the method of the present invention. This latter embodiment is illustrated below in Example 6.

[0026] The method according to the invention may also be used to preferably regenerate catalysts in situ in many chemical processes which require catalyst regeneration by dewaxing, oxidation and reduction, Fischer-Tropsch being one such chemical process.

[0027] In a second aspect, the present invention provides a method for in situ regeneration of a catalyst in a reactor equipped with heat exchange channels, preferably a microchannel reactor, comprising the steps of: x) oxidizing the catalyst by treating it at elevated temperature with an oxidizing gas stream flowing through the process microchannels of the reactor; and y) reducing the resulting oxidized catalyst by treating it with a reducing gas stream flowing through the process microchannels of the reactor at an elevated temperature; In step x), increasing the temperature inside the process microchannels and / or the heat exchange channels to a temperature sufficient to oxidize the catalyst; In the transition from step x) to step y), reducing the temperature inside the process microchannels and / or the heat exchange channels from a temperature sufficient to cause oxidation to a first lower limit of at least 90°C, preferably at least 100°C, more preferably between 140°C and 180°C, most preferably between 145°C and 155°C; Then in step y), increasing the temperature inside the process microchannels and / or the heat exchange channels to a value sufficient to reduce the catalyst; wherein the temperature inside the process microchannels and / or the heat exchange channels is controlled by a heat exchange fluid flowing through the heat exchange channels of the microchannel reactor, and the entire heat exchange fluid does not undergo a phase change.

[0028] The method according to the invention may also be suitably used to regenerate catalysts in situ in a number of chemical processes which require the regeneration of catalysts by oxidation and reduction. The synthesis of methanol is one such chemical process. Others may include the oxidative regeneration of hydroprocessing catalysts, the methanation of carbon monoxide to produce synthetic natural gas, and the redox regeneration of Fischer-Tropsch catalysts where the dewaxing step is carried out by physical means such as solvent extraction.

[0029] Preferably, the heat exchange fluid is water vapor.

[0030] Preferably, the catalyst is a metal-based catalyst, e.g., a Fischer-Tropsch catalyst such as a cobalt or iron-containing catalyst. In the following description, preferred temperatures for dewaxing, oxidation and reduction are given for a cobalt-based Fischer-Tropsch catalyst, although it will be recognized that different types of catalysts may require alternative temperatures to be used, the selection of which is well within the purview of one skilled in the art.

[0031] Preferably, the catalyst is disposed on a porous support.

[0032] Preferably, the oxidizing gas stream comprises oxygen and a non-oxidizing diluent gas. Preferably, the oxygen content of the oxidizing gas stream is less than 21% by volume, preferably less than 15% by volume, more preferably less than 10%, even more preferably less than 5%, most preferably between 1% and 4%. This feature minimizes the risk of uncontrolled exothermic reactions during the oxidation step.

[0033] The temperature of the gas stream is controlled by a heat exchange fluid flowing through the heat exchange channels of the microchannel reactor. Preferably, the heat exchange fluid is water vapor.

[0034] In a preferred embodiment, step a) begins upon cooling down the reactor from synthesis (e.g. FT synthesis) mode to a transition temperature of approximately 170°C for optional nitrogen purging and introduction of hydrogen-containing gas. Hydrocracking occurs during this step, which causes the formation of light hydrocarbons from the residual hydrocarbons in the catalyst bed. The gas environment is maintained at a concentration of 75% or more hydrogen in the reducing gas, preferably 80% to 90% hydrogen.

[0035] Preferably, the dewaxing gas stream comprises hydrogen and may optionally comprise a diluent gas, which may comprise (or be) nitrogen, methane or light hydrocarbons, for example.

[0036] It is recommended that the heat-up under hydrogen-containing gas to a maximum temperature, initiated with liquid water flow in the coolant circuit (when in Fischer-Tropsch synthesis mode), is enabled by an intermediate pressure steam header, while the cool-down typically starts towards the lowest temperature at which superheated steam is available for the transition from liquid water to steam (vapor) flow in the coolant circuit, typically in the range of 140°C to 180°C, more preferably 145°C to 155°C.

[0037] Once the flow of water vapor is established, in a preferred embodiment the temperature of the catalyst bed / reactor / hydrogen containing gas stream is increased to a hold temperature of between 300°C and 400°C, preferably between 330°C and 380°C, most preferably between 340°C and 360°C, and held at or near (preferably within 15°C of) the hold temperature for a period of between 1 hour and 24 hours, preferably between 10 and 20 hours, more preferably between 10 and 15 hours.

[0038] Upon completion of step a), the temperature of the catalyst bed / reactor / gas stream is preferably reduced from the dewaxing temperature to the lowest temperature at which superheated steam is available, typically in the range of 140° C. to 180° C., more preferably 145° C. to 155° C., for inert gas (e.g. nitrogen) purging and subsequent introduction of oxidation gas. This feature minimizes the time required for regeneration and the risk of water hammer in either the reactor or the associated steam drums and piping.

[0039] After completion of the dewaxing step, purging with an inert gas (eg, nitrogen) is completed prior to introducing the oxidizing gas in step b).

[0040] Preferably, the oxidizing gas stream comprises oxygen and a diluent gas. Preferably, the oxygen content of the oxidizing gas stream is less than 21% by volume, preferably less than 15% by volume, more preferably less than 10%, even more preferably less than 5%, most preferably between 1% and 4%. This feature minimizes the risk of uncontrolled exothermic reactions during the oxidation step at temperatures increased by the flow of superheated steam in the coolant channels.

[0041] The diluent gas may include (or be) air, nitrogen, argon, helium, or carbon dioxide, for example.

[0042] Preferably in step b) the temperature of the catalyst bed / reactor / oxidation gas stream is raised to a temperature at which the catalyst is fully oxidized, between 250°C and 325°C, more preferably between 280°C and 300°C. The temperature of the final hold is preferably held at or near (preferably within 15°C of) the hold temperature for a period of between 1 hour and 24 hours, preferably between 10 and 20 hours, more preferably between 10 and 15 hours. Once the hold is completed the temperature is then preferably reduced to the lowest temperature at which superheated steam is available, typically in the range of 140°C to 180°C, more preferably between 145°C to 155°C. This feature minimises the time required for regeneration.

[0043] Preferably, after completion of the oxidation step, purging with an inert gas (eg, nitrogen) is completed before introducing the reducing gas in step c).

[0044] Preferably, in step c), the temperature of the reducing gas stream is raised to a hold temperature of between 300°C and 400°C, preferably between 330°C and 380°C, most preferably between 340°C and 360°C, and held at or near (preferably within 15°C of) the hold temperature for a period of between 1 hour and 24 hours, preferably between 10 and 20 hours, more preferably between 10 and 15 hours.

[0045] Preferably, the reducing gas stream comprises hydrogen, and optionally a diluent gas, which may comprise (or be) for example nitrogen, methane, light hydrocarbons, carbon dioxide or carbon monoxide.

[0046] Preferably, the temperature of the oxidizing gas stream in step b) or step x), or the temperature of the reducing gas stream in step a) or step c) or step y) is changed (increased or decreased) at a rate of 5°C to 30°C per hour, preferably 10°C to 20°C per hour, most preferably 12°C to 18°C ​​per hour.

[0047] Preferably, the temperature in the process microchannels is within 10° C., preferably 5° C., more preferably 2° C., and most preferably 1° C., of the temperature in the adjacent heat transfer channels. This feature minimizes the risk of uncontrolled reaction of the catalyst.

[0048] Preferably, the internal maximum transverse dimension of the process microchannels is no greater than 12 mm, preferably no greater than 5 mm, more preferably no greater than 2 mm, and most preferably no greater than 1 mm. These ranges maximize heat transfer, thereby minimizing the risk of uncontrolled reaction of the catalyst.

[0049] The invention also provides in a second aspect a Fischer-Tropsch process comprising reacting in a Fischer-Tropsch reactor a gas mixture comprising carbon monoxide and hydrogen, and periodically regenerating a catalyst in the Fischer-Tropsch reactor by a method as defined above.

[0050] Preferably, the gas mixture flows in parallel flow paths through multiple Fischer-Tropsch reactors or through multiple Fischer-Tropsch reactor cores of one or more Fischer-Tropsch reactors, the flow paths being isolated in a cyclical fashion, and the dewaxing, oxidation and reduction gas streams of steps a), b) and c) being fed sequentially through the isolated flow paths to regenerate the Fischer-Tropsch catalyst in that flow path while the Fischer-Tropsch reaction occurs in the remaining flow paths. This feature allows for continuous production and avoids plant downtime.

[0051] In a preferred embodiment, the syngas mixture is produced by gasifying biomass and / or municipal or solid waste products, which may be subsequently reformed. Other feedstocks, such as landfill gas or natural gas, may be directly reformed without prior gasification.

[0052] The present invention also provides, in a third aspect, a method according to the above for in situ regeneration of a cobalt-containing, iron-containing or ruthenium-containing Fischer-Tropsch catalyst in a microchannel reactor equipped with heat exchange channels.

[0053] The present invention also provides in a fourth aspect a method according to the above for in situ regeneration of a hydrocarbon processing catalyst in a microchannel reactor equipped with heat exchange channels.

[0054] The present invention also provides, in a fifth aspect, a method for regenerating any catalyst using at least one treatment in a hydrogen-containing process stream and one treatment in a process stream containing oxygen. For example, some chemical processes may not require a dewaxing step; others may achieve dewaxing through physical means such as solvent extraction, in which case regeneration can be completed with subsequent oxidation and reduction steps according to the present invention. For example, a methanol synthesis catalyst can be regenerated with oxidation and reduction steps x) and y) according to the present invention. [Brief description of the drawings]

[0055] Preferred embodiments of the present invention will now be described, by way of example only, with reference to Figures 1 to 7 of the accompanying drawings. [Figure 1] 1 is a temperature plot during a catalyst regeneration process using a heat exchange fluid under conditions of heat transfer involving transition from a liquid phase to a vapor phase or vice versa in the heat exchange fluid (i.e., conventional process). [Diagram 2] 2 is a schematic diagram of a comparative temperature plot illustrating a catalyst regeneration method according to the present invention and according to the method of FIG. 1. [Diagram 3] FIG. 1 is a schematic diagram of a microchannel reactor used in a preferred embodiment. [Figure 4] FIG. 4 is a schematic diagram of a reactor core utilized in the reactor of FIG. [Diagram 5] FIG. 5 is a schematic diagram of a heat exchange unit utilized in the reactor core of FIG. [Figure 6] FIG. 5 is a schematic diagram of a catalytic unit containing process microchannels that is utilized in the reactor core of FIG. 4. [Figure 7] FIG. 1 is a schematic diagram of a Fischer-Tropsch island (facility) with five different reactor trains (each containing one or more microchannel reactors) showing two stages, A) and B), in the operation of the reactor trains, where different reactor trains 200C and 200D are isolated from the Fischer-Tropsch synthesis process for catalyst regeneration. DETAILED DESCRIPTION OF THE PREFERRED EMBODIMENTS

[0056] A microchannel reactor was used that contained two process layers (each containing approximately 500 process channels per layer, as shown in FIG. 6) and three coolant layers (each containing approximately 175 channels per layer, as shown in FIG. 5). The reactor was loaded with a cobalt-based FT catalyst and operated in FT synthesis mode with synthesis gas derived from natural gas (using a steam reforming process) for a period of 815 hours, generating approximately 1.75 H using the membrane. 2 The :CO ratio was then adjusted to 0.05 to 0.15% CO. This was then subjected to the wax removal, oxidation and reduction (WROR) process as summarized in FIG.

[0057] FIG. 1 shows a temperature plot of the regeneration process of the cobalt-based Fischer-Tropsch catalyst described above in the microchannel reactor with cooling with water and steam as the heat exchange fluid (i.e. with phase change and consequent heat removal as latent heat).

[0058] As shown, a three-step process is involved, which includes wax removal, oxidation and reduction (WROR) phases, requiring heating up and cooling down of the catalyst bed (in the reactor) for each phase.

[0059] First, the synthesis is stopped by lowering the reactor temperature to approximately 170° C., then shutting off the synthesis gas (STOP SYNGAS). This is followed by nitrogen and then hydrogen to ensure the environment for the wax removal step. Temperature ramps for wax removal are initiated between WR START, 2 and WR DONE, 3. The initial heat-up is performed with the liquid coolant flow running to a temperature of approximately 210° C. The reactor is then cooled to approximately 170° C., the cooling medium is switched to superheated steam, and the heat-up, hold and cool-down follow the profile shown in FIG. 1. Once cooled down to approximately 150° C., the liquid coolant medium (water) is reintroduced and the reactor is cooled to approximately 70° C. This is followed by purging with nitrogen and starting the gradual controlled introduction of oxidizing gas with OX START, 4, then increasing the oxygen concentration in the system by 1%. Once the final environment is reached, the oxidation temperature ramp begins and ends with OX DONE, 5. Once again, during the heat-up phase, the transition from liquid water coolant to superheated steam coolant is made at a temperature of approximately 150° C., and the reverse transition from superheated steam to liquid coolant is made at approximately the same temperature. Once the oxidation temperature ramp is completed, the reactor is at approximately 70° C. under oxygen-containing gas. This is then purged with nitrogen and then hydrogen to ensure an environment for the reduction step. Third, the reduction phase temperature ramp begins at Rstart, 6, and ends at Rcomplete, 7, when the hydrogen supply is cut off. Once again, during the heat-up phase, the transition from liquid water coolant to superheated steam coolant is made at a temperature of approximately 150° C., and the reverse transition from superheated steam to liquid coolant is made at a synthesis start temperature of approximately 170° C. Regeneration is then completed and the synthesis gas is restarted (START SYNGAS).

[0060] The recovery of catalytic activity following this comparative protocol is illustrated in Table 1 below:

[0061] [Table 1]

[0062] Each of these phases carries the risk of exothermic reactions due to the exothermic nature of the reactions, hydrocracking in the wax removal step (mild) and cobalt oxidation in the oxidation step (heavy), and the potential for reaction between hydrogen and oxygen (with inadequate purging between steps). To mitigate these risks, the transitions between each of these steps are carried out at approximately 70-80°C, although the final holding temperatures in these steps are often in the range of 300-375°C. Heating and cooling over the entire range of temperatures involves the use of circulating cooling water as well as superheated steam. The transitions from the cooling water circulation to superheated steam and vice versa are typically carried out in the range of 150-200°C, exposing the reactor / steam drum to potential water hammer.

[0063] FIG. 2 shows an idealized version of the same temperature profile as FIG. 1 as plot 1, but also shows temperature plot 10 achievable according to the present invention with a cobalt-based Fischer-Tropsch catalyst in the same microchannel reactor. In this case, the heat exchange medium that can be used is superheated steam. The lowest temperature at which superheated steam can be utilized is 150° C., and as a result, the transition between steps occurs at 150° C. instead of 70° C. The heating and cooling rates in plots 1 and 10 were essentially identical at 15° C. / hr. The method of the present invention was found to reduce the time taken for WROR (wax removal, oxidation, reduction) by approximately 24 hr (1 day) over the original process which took 7 days, as illustrated in plot 10. Assuming regeneration every 60 days, or about 6 times per year, the method of the present invention reduces the time taken for regeneration by approximately 6 days, or increases the availability of the Fischer-Tropsch reactor by approximately 2%.

[0064] The results of testing the individual components of the method of the present invention are set forth below: The wax removal parts of temperature plots 1 and 10 are essentially identical, therefore no changes to the practice of the wax removal protocol are necessary. [Example 1-2]

[0065] (About oxidation) The oxidation step in regeneration is the one that is most sensitive to the rate of oxygen introduction. 2 Increasing the inlet temperature is expected to increase reactivity (in the cobalt reoxidation reaction) and the initial O 2 Heat release during introduction is considered.

[0066] A single channel kilopocket reactor was used to measure the modified O 2 The introduction protocol was tested. 2 In the introduction, the thermal response (measured as a temperature spike at a reactor wall thermocouple located in the middle of the wall between the process and coolant channels) and the pressure drop across the catalyst bed were used as indicators to evaluate good air introduction.

[0067] A fresh cobalt-based Fischer-Tropsch catalyst was first activated by reduction in hydrogen and held at a temperature of 150° C., then O 2 was introduced.

[0068] Table 2 shows the O 2 The results of the implementation are summarized and show good agreement with the comparison protocol in terms of the observed maximum temperature rise (measured by the thermowell described above) and the change in pressure drop. 2 The amount of O that needs to be adjusted 2 Depending on the size of the process channel, oxygen is controlled through changes in concentration (exemplified) or flow (not shown) to deliver the correct amount of oxygen. Using a moving front heat release model of a repeating unit (single process and single coolant layer), a detailed mechanical analysis performed according to ASME Section VIII Division 2 was used to evaluate the thermal impact of the oxygen introduction step and to verify the acceptable fatigue life (>1000 thermal cycles) of the reactor.

[0069] [Table 2]

[0070] The maximum temperature rise and pressure drop presented in the examples are within 5% of the comparative values ​​and within the measurement uncertainty of the instruments. Performance is therefore comparable, but regeneration times are significantly reduced and the risk of water hammer is significantly reduced. [Example 3-4]

[0071] (About the reduction) The reduction of the catalyst was investigated with an onset temperature of 150° C. Using a single channel Kilo pocket reactor, the modified reduction protocol was tested after the dewaxing and oxidation steps to confirm acceptable performance.

[0072] To expedite the testing, catalysts previously used in synthesis and WROR work were used in this test. The catalysts were activated by the target protocol, the comparative protocol, and the protocol modified for activation according to the invention, and then the FT synthesis was carried out by H 2 The start-up was at operating conditions corresponding to: CO = 1.82, 41% inerts, 2.41 MPa (350 psig) inlet pressure, and 356 ms contact time. The reactor temperature was initially set at 201°C and then increased to target 75±1% CO conversion. Each protocol was tested in triplicate, and the performance of the catalysts activated by the comparative and modified protocols were found to be statistically indistinguishable.

[0073] A cobalt-based FT catalyst previously used in a synthesis run and subsequently subjected to a dewaxing and oxidation treatment was activated by reduction in incoming hydrogen and synthesis gas. The reactor temperature was set at 201°C and the CO conversion was compared @ 24 hours on stream. The reactor temperature was then increased to cause catalyst deactivation and maintain approximately 75 ± 0.5% CO conversion @ 48-72 hours on stream.

[0074] Table 3 summarizes the FT synthesis performance results. The comparative protocol and the inventive protocol are statistically indistinguishable.

[0075] [Table 3]

[0076] The performance is therefore comparable, but the regeneration time is significantly reduced and the risk of water hammer is significantly reduced. [Example 5]

[0077] (About the whole process) An example of a detailed regeneration protocol for a cobalt-based FT catalyst, implemented as a multi-step process of dewaxing, oxidation and reduction phases, is as follows:

[0078] Wax removal The reducing gas flow was set to the target flow and H 2 Purity is targeted to be greater than 85 mol%. The reactor is pressurized to the target pressure and allowed to heat up from 170°C to 350°C at a rate of ≦15°C / hr. Upon completing the heat up to a temperature of approximately 220°C, a transition is made from liquid water flow to superheated steam as the coolant medium and heat up to the hold temperature is resumed. Once the target hold temperature is reached, the reducing environment is maintained at constant temperature for a period of 12 hr and then cooled down to the target 150°C transition temperature at a rate of ≦15°C / hr.

[0079] Oxidation Before starting the oxidation process, the reactor is purged with nitrogen to eliminate combustible gases (e.g., H used during wax removal). 2 ) must be eliminated. This is done by eliminating pressure-vacuum cycles or N 2 This can be accomplished by purging with 500 psi. Note that the air used for the oxidation process should have a dew point temperature of -40° F., less than <0.1 ppmw particulates, and should be essentially free of S and N contaminants.

[0080] The nitrogen gas flow is set to the target flow and the reactor is pressurized to the target pressure. While maintaining the total flow rate (GHSV), small amounts of air are introduced to raise the oxygen concentration to about 0.1 mol% and hold for a predefined time. Air introduction is continued to raise the oxygen concentration by, for example, 0.1% (with or without hold) until the final target oxygen concentration (e.g., approximately 3 mol%) is reached. The final O 2 After reaching concentration, begin heating the reactor to a temperature of 150° C. to 300° C. at a rate of ≦15° C. / hr. After completing the 12 hr hold period, begin cooling the reactor down from 300° C. to 150° C. at a rate of ≦15° C. / hr. The reactor is purged with nitrogen in preparation for the final reduction step.

[0081] reduction H 2 A long initial purge with H was performed for a period of 4 hours, initially at a transition temperature of 150° C. The reduction gas flow was set to the target flow, and H 2 Purity was targeted to exceed 99.6 mol%. The reactor was pressurized to the target pressure, heated up from 150°C to 350°C at ≦15°C / hr, then held at 350°C for 12 hr, and finally cooled down to the syngas inlet temperature of about 170°C at a rate of ≦15°C / hr. At this state, a switch was made from superheated steam to liquid water as the cooling medium in preparation for FT synthesis. [Example 6]

[0082] (About the whole process) The reactor / catalyst / heat exchange cooldown to 70°C in the comparative process involves two steps in commercial practice, from the final holding temperature for each step with superheated steam heat exchange medium to the saturated steam temperature (where steam drum pressure controls the temperature in the steam drum based on the saturated steam curve). Falling below this temperature requires an additional flush of the steam drum with fresh water make-up and blowdown, or by waiting an extended period to allow the temperature to cool down by natural convection in the comparative process. In the process described by FIG. 1 above, the target rate of cooldown was achieved through a combination of steam drum make-up-blowdown as described above, and the use of removable insulation. In commercial practice, a process utilizing the method of the present invention, albeit on a relatively small scale, has been found to be able to achieve the target benefits of the method of the present invention, with an example transition temperature of about 99-105°C (steam saturation temperature based on operating the steam drum at about 2 psi above ambient pressure).

[0083] An example of a detailed regeneration protocol for a cobalt-based FT catalyst, implemented as a multi-step process of dewaxing, oxidation and reduction phases as described in Example 5, shows that the minimum achievable transition temperature (based on saturated steam pressure, operating the steam drum at approximately 2 psi above ambient pressure, taking into account the local climatic conditions) is 99-105°C.

[0084] Details of a suitable microchannel reactor are provided below in connection with FIGS.

[0085] 3, the microchannel reactor 200 includes a containment vessel 210 that contains or houses three microchannel reactor cores 220. In other embodiments, the containment vessel 210 may be used to contain or house from 1 to about 12 microchannel reactor cores, or from 1 to about 8 microchannel reactor cores, or from 1 to about 4 microchannel reactor cores. The containment vessel 210 may be a pressurizable vessel. The containment vessel 210 includes inlets and outlets 230 that allow reactants to flow to the microchannel reactor core 220, products to flow from the microchannel reactor core 220, and heat exchange fluids to and from the microchannel reactor core 220.

[0086] One of the inlets 245 may be connected to a header or manifold (not shown) provided in each of the microchannel reactor cores 220 for flowing reactants into the process microchannels. One of the inlets 230 is connected to a header or manifold (not shown) provided in each of the microchannel reactor cores 220 for flowing heat exchange fluid, such as superheated steam, into the heat exchange channels. One of the outlets 245 is connected to a manifold or footer (not shown) provided in each of the microchannel reactor cores 220 for flowing products from the process microchannels. One of the outlets 230 is connected to a manifold or footer (not shown) provided in each of the microchannel reactor cores 220 for flowing heat exchange fluid from the heat exchange channels.

[0087] The containment vessel 210 may be constructed using any suitable material sufficient to withstand the operating pressures that may be developed within the microchannel reactor core 220. For example, the shell 240 and reinforcing ribs 242 of the containment vessel 210 may be constructed of cast steel. The flanges 245, couplings and pipes may be constructed of, for example, 316 stainless steel.

[0088] 4, 5 and 6, the microchannel reactor core 220 contains alternating stacks of layered units 300 of process microchannels 310 and layered units 350 of heat exchange channels 355.

[0089] The microchannel reactor core 220 may include a number of plates in a stack that defines a number of process layers and a number of heat exchange layers, each plate having a periphery, and the periphery of each plate or shim being welded to the periphery of the next adjacent plate to provide a peripheral seal in the stack, as shown in US 2012 / 0095268 A1, which is incorporated herein by reference.

[0090] The microchannel reactor core 220 may have the form of a square or rectangular 6-sided solid block. The microchannel reactor core 220 may have the same cross section along its length. The microchannel reactor core 220 may be in the form of a parallelepiped, or cubic block, or prism.

[0091] The Fischer-Tropsch catalyst 500 is located in the process microchannel 310 and may be in any form, including a fixed bed of particulate solids or a variety of structured catalyst forms.

[0092] 4 shows a corrugated sheet 315 sandwiched between plates 316 and 317, defining process microchannels 310 on either side of the sheet 315. For clarity, Fischer-Tropsch catalyst 500 is shown in only three of these microchannels, but in reality each microchannel 310 is filled with catalyst 500. Further details of the construction are disclosed in WO 2008 / 030467(A), which is incorporated herein by reference.

[0093] The Fischer-Tropsch catalyst 500 may comprise cobalt and a support. The catalyst may have a Co loading in the range of about 10 to about 60% by weight, or about 15 to about 60% by weight, or about 20 to about 60% by weight, or about 25 to about 60% by weight, or about 30 to about 60% by weight, or about 32 to about 60% by weight, or about 35 to about 60% by weight, or about 38 to about 60% by weight, or about 40 to about 60% by weight, or about 40 to about 55% by weight, or about 40 to about 50% cobalt.

[0094] The Fischer-Tropsch catalyst 500 may further comprise a precious metal. The precious metal support may be one or more of Pd, Pt, Rh, Ru, Re, Ir, Au, Ag, and Os. The precious metal may be one or more of Pd, Pt, Rh, Ru, Ir, Au, Ag, and Os. The precious metal may be one or more of Pt, Ru, and Re. The precious metal may be Ru. Alternatively, or in addition, the precious metal may be Pt. The Fischer-Tropsch catalyst may comprise from about 0.01 to about 30% total precious metal (based on the total weight of all precious metals present as a total weight percent of the catalyst precursor or activated catalyst), or from about 0.05 to about 20% total precious metal, or from about 0.1 to about 5% total precious metal, or about 0.2% total precious metal.

[0095] The Fischer-Tropsch catalyst 500 may also include one or more other metal-based components as promoters or modifiers. These metal-based components may be present in the catalyst precursor and / or activated catalyst as carbides, oxides, or elemental metals. Suitable metals for the one or more other metal-based components may be, for example, one or more of Zr, Ti, V, Cr, Mn, Ni, Cu, Zn, Nb, Mo, Tc, Cd, Hf, Ta, W, Re, Hg, Tl, and a 4f-block lanthanide. Suitable 4f-block lanthanides may be La, Ce, Pr, Nd, Pm, Sm, Eu, Gd, Tb, Dy, Ho, Er, Tm, Yb, and / or Lu. The metals for the one or more other metal-based components may be, for example, one or more of Zn, Cu, Mn, Mo, and W. The metal for the one or more other metal-based components can be, for example, one or more of Re and Pt. The catalyst may contain from about 0.01 to about 10% total other metals (based on the combined weight of all other metals, as a total weight percent of the catalyst precursor or activated catalyst), or may contain from about 0.1 to about 5% total other metals, or may contain up to about 3% total other metals.

[0096] The Fischer-Tropsch catalyst 500 may be derived from a catalyst precursor that may be activated to produce the Fischer-Tropsch catalyst, for example, by heating the catalyst precursor in hydrogen and / or a hydrocarbon gas (e.g., methane) or in hydrogen or a hydrocarbon gas diluted with another gas, such as nitrogen and / or methane, to convert at least a portion of the carbide or oxide to elemental metal. In the active catalyst, the cobalt may be at least partially in the form of its carbide or oxide.

[0097] The Fischer-Tropsch catalyst precursor may be activated using a carboxylic acid as a reducing agent. The carboxylic acid may be selected to minimize the fracture of the catalyst precursor, but still ultimately produce an effective catalyst. A mixture of two or more carboxylic acids may be used. The carboxylic acid may be an alpha-hydroxy carboxylic acid, such as citric acid, glycolic acid, lactic acid, mandelic acid, or a mixture of two or more thereof.

[0098] The Fischer-Tropsch catalyst 500 may include a catalyst support. The support may include a refractory metal oxide, carbide, carbon, nitride, or a mixture of two or more thereof. The support may include alumina, zirconia, silica, titania, or a mixture of two or more thereof. The surface of the support may be modified by treatment with silica, titania, zirconia, magnesia, chromia, alumina, or a mixture of two or more thereof. The materials used for the support and the materials used to modify the support may be different. The support may include silica, the surface of which may be treated with a refractory solid oxide, such as titania, which is an oxide. The materials used to modify the support may be used to increase the stability of the supported catalyst (e.g., by reducing deactivation). The catalyst support may contain up to about 30% by weight, or such as from about 1% to about 30% by weight, or from about 5% to about 30% by weight, or from about 5% to about 25% by weight, or from about 10% to about 20% by weight, or from about 12% to about 18% by weight of an oxide (e.g., silica, titania, magnesia, chromia, alumina, or a mixture of two or more thereof) used to modify the surface of the support. The catalyst support may be in the form of a structured shape, pellets, or powder. The catalyst support may be in the form of a particulate solid. Without wishing to be bound by theory, it is believed that the surface treatments presented herein help protect the Co from sintering during the operation of the Fischer-Tropsch process.

[0099] The deactivation rate of the Fischer-Tropsch catalyst 500 may be such that it can be used for Fischer-Tropsch synthesis for greater than about 300 hours, or greater than about 3,000 hours, or greater than about 12,000 hours, or greater than about 15,000 hours, all before recovery or regeneration of the catalyst is required.

[0100] The Fischer-Tropsch catalyst 500 may be used for extended periods of time (e.g., greater than 300 hours) with a deactivation rate of less than about 1.4% per day, or less than about 1.2% per day, or between about 0.1% and about 1% per day, or between about 0.03 and about 0.15% per day.

[0101] The Fischer-Tropsch catalyst 500 may have any size and geometric configuration that fits within the process microchannel 310. The catalyst may be in the form of a particulate solid (e.g., pellets, powder, fiber, etc.) having a median particle size of about 1 to about 1000 μm (microns), or about 10 to about 750 μm, or about 25 to about 500 μm. The median particle size may range from 50 to about 500 μm, or about 100 to about 500 μm, or about 125 to about 400 μm, or about 170 to about 300 μm. In one embodiment, the catalyst may be in the form of a fixed bed of particulate solid.

[0102] The microchannel reactor core 220 may contain, for example, six layers 350 of heat exchange channels 355 .

[0103] 6, each unit 300 of the process microchannel 310 may have, for example, a height (h) of 6.35 mm and a width (w) of 165 mm. The length of each process microchannel may be, for example, 600 mm.

[0104] Referring to FIG. 5, each unit 350 of the heat exchange channel 355 may have, for example, a height (h) of 6.35 mm, a width (w) of 6.35 mm and a length (l) of 600 mm.

[0105] Each unit 300 of process microchannels 310 may have, for example, 165 process microchannels 310. The process microchannels 310 may have a cross-section having any shape, for example, square, rectangular, circular, semicircular, etc. The internal height of each process microchannel 310 is considered to be the smaller of the internal dimensions perpendicular to the direction of flow of the reactants and products through the process microchannel. Each of the process microchannels 310 may have, for example, an internal height of 6.35 mm and a width of 1 mm.

[0106] Each unit 350 of heat exchange channels 355 may have, for example, 168 heat exchange channels. The heat exchange channels 355 may be microchannels or they may have larger dimensions that would classify them as not being microchannels. Each of the heat exchange channels 355 may have, for example, an inner height or width of 6.35 mm.

[0107] The microchannel reactor core 220 may be made of any material that provides sufficient strength, dimensional stability, and heat transfer properties to enable the operation of the desired process. These materials may include, for example, aluminum; titanium; nickel; platinum; rhodium; copper; chromium; alloys of any of the foregoing metals; brass; steel (e.g., stainless steel); quartz; silicon; or combinations of two or more thereof. Each microchannel reactor may be constructed of stainless steel and one or more corrugated copper or aluminum sheets used to form the channels.

[0108] The microchannel reactor core 220 may be fabricated using known techniques including, for example, wire electrical discharge machining, conventional machining, laser cutting, photochemical machining, electrochemical machining, molding, water jet, stamping, etching (e.g., chemical, photochemical or plasma etching), and combinations thereof.

[0109] The microchannel reactor core 220 may be constructed by forming plates with portions removed to allow flow through. Stacks of plates can be assembled, for example, via diffusion bonding, laser welding, diffusion brazing, and similar methods to form an integrated device. The microchannel reactor can be assembled, for example, using a combination of plates and partial plates or strips. In this method, channels or void areas can be formed by assembling strips or partial plates to reduce the amount of material required.

[0110] The microchannel reactor core 220 may include multiple plates in a stack defining multiple process layers and multiple heat exchange layers, each plate having a periphery, and the periphery of each plate or shim being welded to the periphery of the next adjacent plate to provide a peripheral seal for the stack, as shown in US 2012 / 0095268 A1, which is incorporated herein by reference.

[0111] The containment vessel 210 may include a control mechanism for maintaining the pressure within the containment vessel at a level at least as high as the internal pressure within the microchannel reactor core 220. The internal pressure within the containment vessel 210 may range from about 10 to about 60 atmospheres, or from about 15 to about 30 atmospheres during operation of a synthesis gas conversion process (e.g., a Fischer-Tropsch process). The control mechanism for maintaining the pressure within the containment vessel may include a check valve and / or a pressure regulator. The check valve or regulator may be programmed to effectuate any internal pressure desired for the containment vessel. Either or both of these may be used in conjunction with a system of pipes, valves, controllers, etc. to ensure that the pressure in the containment vessel 210 is maintained at a level at least as high as the internal pressure within the microchannel reactor core 220. This is done in part to protect the welds used to form the microchannel core 220. A significant drop in pressure within the containment vessel 210 could result in costly rupture of welds in the microchannel reactor core 220 without a corresponding drop in internal pressure within the microchannel reactor core 220. The control mechanism may be designed to allow bypass of one or more process gases to the containment vessel when pressure exerted by the containment vessel gases is reduced.

[0112] Fischer-Tropsch microchannels are characterized by having a bulk flow path. The term "bulk flow path" refers to an open pathway (continuous bulk flow region) within a process microchannel. The continuous bulk flow region allows rapid fluid flow through the channel without a large pressure drop. In one embodiment, the fluid flow in the bulk flow region is laminar. The bulk flow region within each process microchannel may range from about 0.05 to about 10,000 mm 2 , or about 0.05 to about 5000 mm 2 , or about 0.1 to about 2500 mm 2 The bulk flow region may comprise from about 5% to about 95%, or from about 30% to about 80% of the cross section of the process microchannel.

[0113] Contact time of the reactants with the catalyst may range up to about 3600 milliseconds (ms), or up to about 2000 ms, or from about 10 to about 2600 ms, or from about 10 ms to about 2000 ms, or from about 20 ms to about 500 ms, or from about 200 to about 450 ms, or from about 240 to about 350 ms.

[0114] The space velocity (or gas hourly space velocity (GHSV)) for fluid flow in the process microchannels is at least about 1000 hr -1 (typical liters of feed / hour / liters of volume in the process microchannel), or at least about 1800 hr -1 , or about 1000 to about 1,000,000 hours -1 , or about 5000 to about 20,000 hours -1 may be also possible.

[0115] The pressure in the process microchannels may be up to about 100 atmospheres, or may range from about 1 to about 100 atmospheres, or from about 1 to about 75 atmospheres, or from about 2 to about 40 atmospheres, or from about 2 to about 10 atmospheres, or from about 10 to about 50 atmospheres, or from about 20 to about 30 atmospheres.

[0116] The pressure drop as the fluid flows through the process microchannel may range from about 30 atmospheres per meter of channel length (atm / m), or about 25 atm / m, or about 20 atm / m. The pressure drop may range from about 10 to about 20 atm / m.

[0117] In a preferred embodiment, the reactor has a heat transfer surface (or heat transfer wall) for removing heat of reaction from the reactor (or process microchannel layer), and the ratio of the surface area of ​​the heat transfer surface to the volume of catalyst in the reactor is at least about 300 square meters of heat transfer surface per cubic meter of catalyst, for example about 300 to about 5000, or preferably about 1000 to 3000 m. 2 / m 3 It is a catalyst.

[0118] The heat flux for heat exchange in the microchannel reactor core 220 is about 0.01 to about 500 watts per square centimeter of surface area (W / cm) of one or more heat transfer walls of the process microchannels in the microchannel reactor. 2 ) or from about 0.1 to about 250 W / cm 2 , or about 1 to about 125 W / cm 2 , or about 1 to about 100 W / cm 2 , or about 1 to about 50 W / cm 2 , or about 1 to about 25 W / cm 2 , or about 1 to about 10 W / cm 2 The range may be from about 0.2 to about 5 W / cm. 2 , or about 0.5 to about 3 W / cm 2 , or about 1 to about 2 W / cm 2 may be also possible.

[0119] 7, a chain of microchannel reactors 200A to 200E is shown in two states A) and B). The microchannel reactors are each fed in parallel with synthesis gas (SYNGAS) from a common supply line, and the products (FT products) are combined in parallel as shown.

[0120] In state A), reactor 200C is isolated and its catalyst is regenerated according to the protocol of plot 10 of Figure 2. When this regeneration is complete, the catalyst is returned to the Fischer-Tropsch operation by restarting the flow of SYNGAS and into the FT product line, and a similar regeneration is performed on the catalyst in reactor 200D, as shown in state B). The regeneration is cycled through each of reactors 200A through 200D, so that four of the reactors are in service of the Fischer-Tropsch process at any one time, while the remaining reactor is having its catalyst regenerated.

[0121] The superficial velocity of the fluid flowing within the process microchannel may be at least about 0.01 meters per second (m / s), or at least about 0.1 m / s, or in the range of about 0.01 to about 100 m / s, or in the range of about 0.01 to about 10 m / s, or in the range of about 0.1 to about 10 m / s, or in the range of about 1 to about 100 m / s, or in the range of about 1 to about 10 m / s.

[0122] The free stream velocity of the fluid flowing within the process microchannel may be at least about 0.001 m / s, or at least about 0.01 m / s, or in the range of about 0.001 to about 200 m / s, or in the range of about 0.01 to about 100 m / s, or preferably in the range of about 0.01 to about 200 m / s.

[0123] The conversion of CO from fresh syngas may be about 70% or more, or about 75% or more, or about 80% or more, or about 90% or more, or about 91% or more, or about 92% or more, or about 88% to about 95%, or about 90% to about 94%, or about 91% to about 93%. If tail gas recycle is used, the one-pass conversion of CO relative to the CO in the reactant mixture (i.e., fresh syngas and recycled tail gas) may range from about 50% to about 90%, or from about 60% to about 85%.

[0124] The selectivity to methane in the Fischer-Tropsch (FT) product may range from about 0.01 to about 10%, or from about 1% to about 5%, or from about 1% to about 10%, or from about 3% to about 9%, or from about 4% to about 8%.

[0125] The Fischer-Tropsch product may include a gaseous product fraction and a liquid product fraction. The gaseous product fraction may include hydrocarbons boiling below about 350° C. at atmospheric pressure (e.g., tail gas to middle distillates). The liquid product fraction (condensate fraction) may include hydrocarbons boiling above about 350° C. (e.g., vacuum gas oil to heavy paraffins).

[0126] The Fischer-Tropsch product fraction boiling below about 350° C. may be separated into a tail gas fraction and a condensate fraction, e.g., normal paraffins and higher boiling hydrocarbons of about 5 to about 20 carbon atoms, using, for example, a high pressure and / or low pressure vapor-liquid separator, or a low pressure separator, or a combination of separators. The fraction boiling above about 350° C. (condensate fraction) may be separated into a wax fraction boiling in the range of about 350° C. to about 650° C., after removal of one or more fractions boiling above about 650° C. The wax fraction may contain linear paraffins of about 20 to about 50 carbon atoms along with relatively small amounts of higher boiling branched paraffins. Separation may be accomplished using fractional distillation.

[0127] The Fischer-Tropsch products may include methane, wax and other heavy, high molecular weight products. The products may include olefins such as ethylene, normal and iso-paraffins, and combinations thereof. They may include hydrocarbons in the distillate fuel range, including the jet or diesel fuel range.

[0128] Branching can be advantageous in some end uses, especially where high octane numbers and / or low pour points are desired. The degree of isomerization may be greater than about 1 mole of isoparaffins per mole of n-paraffins, or about 3 moles of isoparaffins per mole of n-paraffins. When used in diesel fuel compositions, the product may include a hydrocarbon mixture having a cetane number of at least about 60.

Claims

1. 1. A method for in situ catalyst regeneration in a microchannel reactor equipped with heat exchange channels, comprising: a) dewaxing said catalyst by treating it at elevated temperature with a hydrogen-containing dewaxing gas stream flowing through the process microchannels of said microchannel reactor; b) oxidizing the resulting dewaxed catalyst by treating it with an oxidizing gas stream flowing through the process microchannels of the microchannel reactor at elevated temperature; and c) reducing the resulting oxidized catalyst by treating it with a reducing gas stream flowing through the process microchannels of the microchannel reactor at an elevated temperature; in transition from step a) to step b), reducing the temperature inside the process microchannels and / or the heat exchange channels from a temperature sufficient for dewaxing to a first lower limit of from 140° C. to 180° C.; In step b), increasing the temperature inside the process microchannels and / or the heat exchange channels to a temperature sufficient to oxidize the catalyst; In a transition from step b) to step c), reducing the temperature inside the process microchannels and / or the heat exchange channels from a temperature sufficient to cause oxidation to a first lower limit of between 140° C. and 180° C.; Then in step c), increasing the temperature inside the process microchannels and / or the heat exchange channels to a value sufficient to reduce the catalyst; The method, wherein the temperature inside the process microchannels and / or the heat exchange channels is controlled by a heat exchange fluid flowing through the heat exchange channels of the microchannel reactor, and the entirety of the heat exchange fluid does not undergo a phase change.

2. 2. The method of claim 1, wherein step a) is initiated upon cooling down the microchannel reactor from synthesis mode to a transition temperature of approximately 170° C. for the introduction of the hydrogen-containing dewaxing gas.

3. 3. A process according to claim 1 or 2, wherein in step a) the temperature of the catalyst bed of the microchannel reactor and / or the dewaxed gas stream is raised to a temperature of from 250°C to 400°C and held at or within 15°C of that holding temperature for a period of from 1 hour to 24 hours.

4. The method according to any one of claims 1 to 3, wherein the heat exchange fluid is water vapor.

5. The method according to any one of claims 1 to 3, wherein the catalyst is a metal-based catalyst.

6. The process of claim 5, wherein the catalyst is a Fischer-Tropsch catalyst.

7. 7. The process of claim 5 or 6, wherein the catalyst is a cobalt or iron containing catalyst.

8. The method according to any one of claims 1 to 7, wherein the catalyst is disposed on a porous support.

9. The method according to any of claims 1 to 8, wherein the temperature of each gas stream is controlled by a heat exchange fluid flowing through the heat exchange channels of the microchannel reactor.

10. A process according to any of the preceding claims, wherein in step b) the temperature of the catalyst bed of the microchannel reactor and / or the oxidation gas stream is raised to a temperature between 250°C and 325°C at which the catalyst is fully oxidized.

11. 11. A method according to any of the preceding claims, wherein in step c) the temperature of the reducing gas stream is increased to a hold temperature of between 300°C and 400°C and held at or within 15°C of the hold temperature for a period of between 1 hour and 24 hours.

12. 12. A Fischer-Tropsch process comprising reacting in a Fischer-Tropsch reactor a gas mixture comprising carbon monoxide and hydrogen, and periodically regenerating a catalyst in said Fischer-Tropsch reactor by a method according to any one of claims 1 to 11.

13. The method according to any of the preceding claims, wherein the entire heat exchange fluid does not undergo a phase change in the method.

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