PROCESS FOR OPERATING A RADIAL REACTOR

CROSS-REFERENCE TO RELATED APPLICATIONS

[0001] This application claims the benefit of priority to U.S. Provisional Application

No. 62/323,395, filed April 15, 2016, the entirety of which is hereby incorporated by reference for any and all purposes.

FIELD

[0002] The present technology generally relates to processes for conversion of synthesis gas in a tubular reactor to produce hydrocarbons. The tubular reactor includes one or more catalyst carriers that contain a solid synthesis catalyst, such as a particulate Fischer- Tropsch catalyst.

SUMMARY

[0003] A process is provided for the conversion of synthesis gas. The process includes contacting in a tubular reactor a gaseous stream with a solid synthesis catalyst to produce a synthetic product. The gaseous stream includes synthesis gas. The tubular reactor of the process includes (a) a reactor inlet in fluid communication with one or more reactor tubes wherein each reactor tube includes a tube inlet, a tube outlet located downstream of the tube inlet, an inner tube wall, an outer tube wall, and one or more catalyst carriers within the reactor tube, (b) a reactor outlet located downstream of the reactor inlet in fluid

communication with the one or more reactor tubes, and (c) a cooling medium in contact with the one or more reactor tubes. The one or more reactor tubes of the process may each independently have a diameter of about 30 mm to about 300 mm. The process may include introducing the gaseous stream through the reactor inlet at a pressure from about 250 psig to about 1,000 psig.

[0004] The catalyst carrier within the reactor tube may be in an "outward flow" configuration or an "inward flow" configuration. The outward flow catalyst carrier may include (i) an annular container, the annular container comprising a perforated inner wall defining a central tube, a perforated outer wall, a top surface closing the annular container and a bottom surface closing the annular container; wherein the central tube has a top and bottom according to the top surface and the bottom surface and the annular container holds the solid synthesis catalyst; (ii) a surface closing the bottom of the central tube; (iii) a skirt extending upwardly from the perforated outer wall of the annular container from a position at or near the bottom surface of the annular container to a position below the location of a seal; where (iv) the seal is located at or near the top surface and extending from the container by a distance which extends beyond an outer surface of the skirt. The inward flow catalyst carrier may include (i) an annular container, the annular container comprising a perforated inner wall defining a central tube, a perforated outer wall, a top surface closing the annular container and a bottom surface closing the annular container; wherein the central tube has a top and bottom according to the top surface and the bottom surface and the annular container holds the solid synthesis catalyst; (ii) a surface closing the top of the central tube; (iii) a skirt extending downwardly from the perforated outer wall of the annular container from a position at or near the top surface of the annular container to a position above the location of a seal; where (iv) the seal is located at or near the bottom surface and extending from the container by a distance which extends beyond an outer surface of the skirt.

[0005] The solid synthesis catalyst may be a particulate Fischer-Tropsch catalyst.

The particulate Fischer-Tropsch catalyst may have a Co loading from about 20 wt% to about 56 wt%; and a catalyst gas hourly space velocity of the gaseous stream in the tubular reactor may be from about 5,000 hr ^-1 to about 15,000 hr ^-1 . The cooling medium may also be at a temperature of about 160 °C to about 240 °C. The process may involve maintaining at least about 55% CO conversion per pass in the one or more reactor tubes.

[0006] An average annular space between the outer surface of the skirt and the inner surface of the tube wall may be from about 1.3 mm to about 2.6 mm in one or more of the catalyst carriers. A total combined surface area of the outer surface of the skirt and the inner tube wall, and optionally a heat transfer structure, per volume of the solid synthesis catalyst in the catalyst carrier (the "SA/V") may be from about 400 m ² /m ³ to about 40,000 m ² /m ³ ;

[0007] As noted above, a heat transfer structure may be included, where the heat transfer structure includes at least one of (i) a plurality of fins extending radially from the perforated inner wall to the perforated outer wall;

(ii) a plurality of fins extending radially from the perforated inner wall to an inner surface of the skirt; and

(iii) a network of heat conducting surfaces in conductive thermal contact with a

portion the inner tube wall and a portion of the outer surface of the skirt.

[0008] A further related aspect is provided where the tubular reactor of the process includes (a) a reactor inlet in fluid communication with one or more reactor tubes wherein each reactor tube includes a tube inlet, a tube outlet located downstream of the tube inlet, an inner tube wall, an outer tube wall, and one or more catalyst carriers within the reactor tube, (b) a reactor outlet located downstream of the reactor inlet in fluid communication with the one or more reactor tubes, and (c)a cooling medium in contact with the one or more reactor tubes. The one or more reactor tubes of the related process each independently have a diameter of about 40 mm to about 65 mm, and the process may include introducing the gaseous stream through the reactor inlet at a pressure from about 250 psig to about 1,000 psig. For each catalyst carrier independently, an average annular space between the outer surface of the skirt and the inner surface of the tube wall is from about 1.3 mm to about 10.0 mm.

BRIEF DESCRIPTION OF THE DRAWINGS

[0009] FIG. 1 illustrates a top view of a catalyst carrier ("can / ^' ") utilized in reactor modeling, according to the examples.

[0010] FIG. 2 A illustrates a cross-sectional view of can i with outward flow directions and process variables used in the reactor modeling analysis, according to the examples. FIG. 2B illustrates a cross-section view of a catalyst carrier with inward flow directions indicated.

[0011] FIG. 3 illustrates the effect that generating a higher temperature gradient (the difference between the highest temperature from a gaseous stream exiting a catalyst container and the heat exchanger tube wall temperature) exhibits on the CO conversion at an inlet pressure of 1,000 psig, according to the examples.

[0012] FIG. 4 illustrates the effect of the higher temperature gradient on the methane

(CH ₄ ) selectivity at an inlet pressure of 1,000 psig, according to the examples.

[0013] FIG. 5 illustrates the effect of the temperature gradient on the CO conversion at an inlet pressure of 750 psig, according to the examples.

[0014] FIG. 6 plots the methane selectivity as the temperature gradient is increased at an inlet pressure of 750 psig, according to the examples.

[0015] FIG. 7 plots the temperature gradient versus CO conversion at an inlet pressure of 350 psig, according to the examples.

[0016] FIG 8 plots methane selectivity as the temperature gradient is increased at a reactor inlet pressure of 350 psig, according to the examples.

[0017] FIG. 9 plots the CO conversion versus the thermal gradient for a first nonuniform catalyst profile at 750 psig, according to the examples.

[0018] FIG. 10 plots the methane selectivity versus the thermal gradient for a first non-uniform catalyst profile at 750 psig, according to the examples.

[0019] FIG. 11 plots the CO conversion versus the thermal gradient for a second nonuniform catalyst profile at 750 psig, according to the examples.

[0020] FIG. 12 plots the methane selectivity versus the thermal gradient for a second non-uniform catalyst profile at 750 psig, according to the examples.

[0021] FIG. 13 plots the CO conversion versus the thermal gradient for a third nonuniform catalyst profile at 750 psig, according to the examples.

[0022] FIG. 14 plots the methane selectivity versus the thermal gradient for a third non-uniform catalyst profile at 750 psig, according to the examples. [0023] FIG. 15 plots the overall CO conversion and selectivity to methane and C ₅₊ species versus can i required to replicate the productivity reported in Gamlin, according to the examples.

[0024] FIG. 16 illustrates the can inlet and outlet temperatures (left axis) and can heat production (right axis) plotted versus can number I for Example 4.

[0025] FIG. 17 illustrates inlet and heat exchanger temperatures used to maintain a maximum can exit temperature less than 240 °C as plotted versus gas hourly space velocity for heat exchanger gaps between 1.0 to 5.0 millimeters (provided in the accompanying legend) for the 80 can reactor of Example 3.

[0026] FIG. 18 plots the number of tubes required to achieve a C ₅ + production of

10,000 BPD as a function of gas hourly space velocity for heat exchanger gaps between 1.0 to 5.0 millimeters for the 80 can reactor of Example 4.

[0027] FIG. 19 provides CO conversion plotted versus gas hourly space velocity for heat exchanger gaps between 1.0 to 5.0 millimeters for the 80 can reactor of Example 4.

[0028] FIG. 20 provides maximum can exit temperature to heat exchanger wall temperature difference plotted versus gas hourly space velocity for heat exchanger gaps between 1.0 to 5.0 millimeters for the 80 can reactor of Example 4.

[0029] FIG. 21 illustrates reactor pressure drop plotted versus gas hourly space velocity for gas hourly space velocity for heat exchanger gaps between 1.0 to 5.0 millimeters for the 80 can reactor of Example 4.

[0030] FIG. 22 provides reactor C ₅₊ yield plotted versus gas hourly space velocity for heat exchanger gaps between 1.0 to 5.0 millimeters for the 80 can reactor of Example 4.

[0031] FIG. 23 provides inlet and heat exchanger temperatures used to maintain a maximum can exit temperature less than 240 °C plotted versus gas hourly space velocity for heat exchanger gaps between 1.0 to 5.0 millimeters for the 200 can reactor of Example 5. [0032] FIG. 24 provides number of tubes required to achieve a C ₅ + production of

10,000 BPD as a function of gas hourly space velocity for heat exchanger gaps between 1.0 to 5.0 millimeters for the 200 can reactor of Example 5.

[0033] FIG. 25 provides CO conversion plotted versus gas hourly space velocity for heat exchanger gaps between 1.0 to 5.0 millimeters for the 200 can reactor of Example 5.

[0034] FIG. 26 provides C ₅₊ yield plotted versus gas hourly space velocity for heat exchanger gaps between 1.0 to 5.0 millimeters for the 200 can reactor of Example 5.

[0035] FIG. 27 provides maximum can exit temperature to heat exchanger wall temperature difference plotted versus gas hourly space velocity for heat exchanger gaps between 1.0 to 5.0 millimeters for the 200 can reactor of Example 5.

[0036] FIG. 28 provides reactor pressure drop plotted versus gas hourly space velocity for heat exchanger gaps between 1.0 to 5.0 millimeters for the 200 can reactor of Example 5.

[0037] FIG. 29 illustrates the results provided by modeling the flow distribution of the gaseous stream exiting the perforated outer wall of three different annular containers, where the annular space and the inner gap were varied, according to Example 6.

[0038] FIG. 30 illustrates a plurality of fins extending radially from the perforated inner wall to the perforated outer wall of catalyst carriers of certain embodiments of the present technology.

DETAILED DESCRIPTION

I. Definitions

[0039] The following terms are used throughout as defined below.

[0040] As used herein and in the appended claims, singular articles such as "a" and

"an" and "the" and similar referents in the context of describing the elements (especially in the context of the following claims) are to be construed to cover both the singular and the plural, unless otherwise indicated herein or clearly contradicted by context. Recitation of ranges of values herein are merely intended to serve as a shorthand method of referring individually to each separate value falling within the range, unless otherwise indicated herein, and each separate value is incorporated into the specification as if it were individually recited herein. All methods described herein can be performed in any suitable order unless otherwise indicated herein or otherwise clearly contradicted by context. The use of any and all examples, or exemplary language (e.g., "such as") provided herein, is intended merely to better illuminate the embodiments and does not pose a limitation on the scope of the claims unless otherwise stated. No language in the specification should be construed as indicating any non-claimed element as essential.

[0041] As used herein, "about" will be understood by persons of ordinary skill in the art and will vary to some extent depending upon the context in which it is used. If there are uses of the term which are not clear to persons of ordinary skill in the art, given the context in which it is used, "about" will mean up to plus or minus 10% of the particular term.

[0042] As will be understood by one skilled in the art, for any and all purposes, particularly in terms of providing a written description, all ranges disclosed herein also encompass any and all possible sub-ranges and combinations of sub-ranges thereof. Any listed range can be easily recognized as sufficiently describing and enabling the same range being broken down into at least equal halves, thirds, quarters, fifths, tenths, etc. As a non- limiting example, each range discussed herein can be readily broken down into a lower third, middle third and upper third, etc. As will also be understood by one skilled in the art all language such as "up to," "at least," "greater than," "less than," and the like include the number recited and refer to ranges which can be subsequently broken down into sub-ranges as discussed above. Finally, as will be understood by one skilled in the art, a range includes each individual member. Thus, for example, a group having 1-3 atoms refers to groups having 1, 2, or 3 atoms. Similarly, a group having 1-5 atoms refers to groups having 1, 2, 3, 4, or 5 atoms, and so forth.

[0043] The "activity" of the solid synthesis catalysts described herein refers to the primary catalytic metal surface area per unit of packed bed volume in the solid synthesis catalyst. For example, a particulate Fischer- Tropsch catalyst with a higher Co surface area per packed bed volume may be described as having a "higher activity" than a particulate Fischer- Tropsch catalyst with a lower Co surface area per packed bed volume.

[0044] The "primary catalytic metal" of a solid synthesis catalyst refers to the solid synthesis catalysts with more than one catalytic metal, where the primary catalytic metal is present at a weight percent (wt%) greater than the other catalytic metals.

[0045] The term "catalyst gas hourly space velocity" refers to the total gaseous feed flow at standard conditions (0 °C, 1 atm) divided by the reactor volume that contains catalyst and typically recorded in units of hr ^"1 . Moreover, discussion herein using the term "gas hourly space velocity" is to be understood as referring to catalyst gas hourly space velocity.

[0046] A "portion" of a composition or stream, as used herein, means from about 1% to about 100% by volume of the composition or stream, or any range including or in between any two integers from about 1% to about 100%. A "portion" of a surface means from about 1%) to about 100%) by surface area of the surface, or any range including or in between any two integers from about 1%> to about 100%>.

[0047] The term "fluid" refers to a gas, a liquid, or a mixture of a gas and a liquid, wherein the fluid may further contain dispersed solids, liquid droplets and/or gaseous bubbles. The droplets and/or bubbles may be irregularly or regularly shaped and may be of similar or different sizes.

[0048] The term "thermal contact" refers to two bodies, for example, two metals, that may or may not be in physical contact with each other or adjacent to each other but still exchange heat with each other. One body in thermal contact with another body may heat or cool the other body.

[0049] The term "conductive thermal contact" refers to two bodies, for example, two metals, where at least some minimal amount of physical contact with each other is present such that there is a conductive heat flow path between the two bodies. [0050] As used herein, C _m -C _n , such as C ₁ -C ₁₂ , Ci-C ₈ , or Ci-C ₆ when used before a group refers to that group containing m to n carbon atoms. For example, C ₁ -C4 refers to a group that contains 1, 2, 3, or 4 carbon atoms.

II. The Present Technology

[0051] For highly exothermic synthetic processes that utilize synthesis gas, heat removal is a primary concern. For example, the Fischer-Tropsch reaction is highly exothermic and heat removal is essential to avoid thermal runaway and/or poor product selectivity. The thermal gradient is the difference between the maximum and minimum temperatures to which the catalyst is exposed. It is well appreciated to operate the Fischer- Tropsch reaction with as low a temperature as possible while achieving the target conversion and with low thermal gradients to keep the maximum catalyst temperature low. Different styles of reactors have emerged to manage heat removal over the past century.

[0052] The fixed bed reactor is one style where heat is transferred radially to the walls of small diameter tubes (e.g., 2-5 cm) while also transferring heat to co-produced and recycled liquids. A traditional fixed bed reactor has a thermal gradient of around 10-20 °C, while the highly intensified microchannel reactor platform reduces thermal gradients to on the order of 1-5 °C to provide exceptional productivity. In general, lower temperature gradients are beneficial to improve selectivity to desired hydrocarbons while minimizing methane production.

[0053] U. S. Pat. Pub. No. 2014/0187653 Al to Gamlin (hereinafter, "Gamlin") is related to yet another reactor design, where low conversion is achieved in a series catalyst carriers, or "cans," (where each catalyst carrier is essentially an adiabatic stage in the Fischer- Tropsch reactor) within tubes in a reactor before transferring heat at the exterior of each can in an annular space. The cans include an annular space from 3.0 mm to 10.0 mm between the exterior surface of the can and the heat transfer surface. Gamlin teaches that within a single tube, there may be about 80 to 200 cans operating in series so that the reactor does not experience thermal runaway. Multiple tubes are operated in parallel to achieve the desired plant capacity. Gamlin teaches that when utilizing either a particulate bed or monolith catalyst structure within the catalyst carriers, the tubular reactor will have a maximum temperature difference between the catalyst container exit gas and the heat transfer fluid of 10-40 °C and a reasonable temperature rise of 10 to 20 °C over each catalyst carrier - the same thermal gradient observed in typical Fischer- Tropsch fixed bed reactors. Gamlin teaches further that for one arrangement, the reaction temperature ranges from 190 to 250 °C.

[0054] De Geugd (De Geugd, R.M. "Fischer- Tropsch Synthesis Revisited; Efficiency and Selectivity Benefits from Imposing Temporal and/or Spatial Structure in the Reactor," Jan. 27, 2004 (P onsen & Looijen B.V., Wageningen, The Netherlands) similarly teaches the benefits of a 10 to 20 °C thermal gradient. De Geugd describes a physical structure for Fischer Tropsch based on a monolithic catalyst that itself has poor radial heat transfer due to a low effective thermal conductivity through the catalyst particulate bed and operates in essentially an adiabatic mode, similar to Gamlin' s reactor. The de Geugd monolith is closely followed by a heat exchanger to remove the exothermic reaction heat produced in the monolith. The de Geugd reactor is operated in a recycle mode in order to achieve an overall high conversion. De Geugd discloses a maximum temperature gradient of 23 °C and teaches that a lower gradient leads to improved performance and greater selectivity for desired higher carbon number hydrocarbon products (C ₅ + hydrocarbons). At higher temperature gradients, the selectivity to the undesired methane increases substantially while reducing the selectivity to the desired C ₅ + or wax fractions.

[0055] In the present technology, significantly better performance is surprisingly achieved by providing higher thermal gradients, including those in the range from 30 °C to 80 °C. Significantly better performance is also provided by utilizing catalyst carriers with annular spaces from about 1.1 mm to about 2.6 mm and/or by utilizing the herein disclosed heat transfer structures. Where hydrocarbons are formed, this performance is exemplified in significantly higher CO conversion and significantly higher C ₅ + hydrocarbon selectivity at reactor inlet pressures of about 250 psig or greater - results quite opposed to the conventional understanding in the art regarding temperature gradients and tubular reactors.

[0056] Accordingly, in an aspect, a process is provided for the conversion of synthesis gas. The process includes contacting in a tubular reactor a gaseous stream with a solid synthesis catalyst to produce a synthetic product. The gaseous stream includes synthesis gas. A catalyst gas hourly space velocity of the gaseous stream in the tubular reactor may be from about 5,000 hr ^-1 to about 15,000 hr ^-1 . The tubular reactor of the process includes (a) a reactor inlet in fluid communication with one or more reactor tubes wherein each reactor tube includes a tube inlet, a tube outlet located downstream of the tube inlet, an inner tube wall, an outer tube wall, and one or more catalyst carriers within the reactor tube, (b) a reactor outlet located downstream of the reactor inlet in fluid

communication with the one or more reactor tubes, and (c) a cooling medium at a temperature of about 160 °C to about 240 °C in contact with the one or more reactor tubes. The process may include introducing the gaseous stream through the reactor inlet at a pressure from about 50 psig to about 1,000 psig with a ratio of H ₂ /CO in the synthesis gas from about 1.6 to about 2.5, preferably at a pressure from about 250 psig to about 1,000 psig with a ratio of H ₂ /CO in the synthesis gas from about 1.6 to about 2.0. The pressure for introducing the gaseous stream through the reactor inlet may be about 50 psig, about 100 psig, about 150 psig, about 200 psig, about 250 psig, about 300 psig, about 350 psig, about 400 psig, about 450 psig, about 500 psig, about 550 psig, about 600 psig, about 650 psig, about 700 psig, about 750 psig, about 800 psig, about 850 psig, about 900 psig, about 950 psig, about 1,000 psig, or any range including and between any two of these values. For example, the process may preferably include introducing the gaseous stream through the reactor inlet at a pressure from about 500 psig to about 1,000 psig, more preferably at a pressure from about 500 psig to about 750 psig. The ratio of H ₂ /CO in the synthesis gas may be about 1.6, about 1.7, about 1.8, about 1.9, about 2.0, about 2.1, about 2.2, about 2.3, about 2.4, about 2.5, or any range including and/or in between any two of these values.

[0057] The catalyst carrier within the reactor tube may be in an "outward flow" configuration or an "inward flow" configuration. The outward flow catalyst carrier includes (i) an annular container, the annular container comprising a perforated inner wall defining a central tube, a perforated outer wall, a top surface closing the annular container and a bottom surface closing the annular container; wherein the central tube has a top and bottom according to the top surface and the bottom surface and the annular container holds the solid synthesis catalyst; (ii) a surface closing the bottom of the central tube; (iii) a skirt extending upwardly from the perforated outer wall of the annular container from a position at or near the bottom surface of the annular container to a position below the location of a seal; where (iv) the seal is located at or near the top surface and extending from the container by a distance which extends beyond an outer surface of the skirt. The inward flow catalyst carrier includes (i) an annular container, the annular container comprising a perforated inner wall defining a central tube, a perforated outer wall, a top surface closing the annular container and a bottom surface closing the annular container; wherein the central tube has a top and bottom according to the top surface and the bottom surface and the annular container holds the solid synthesis catalyst; (ii) a surface closing the top of the central tube; (iii) a skirt extending downwardly from the perforated outer wall of the annular container from a position at or near the top surface of the annular container to a position above the location of a seal; where (iv) the seal is located at or near the bottom surface and extending from the container by a distance which extends beyond an outer surface of the skirt. The "outer surface of the skirt" is the surface oriented away from the perforated outer wall. For each catalyst carrier independently, an average annular space between the outer surface of the skirt that is closest to the seal and the inner surface of the tube wall is from about 1.1 mm to about 10.0 mm. The average annular space for each catalyst carrier may independently be about 1.1 mm, about 1.2 mm, about 1.3 mm, about 1.4 mm, about 1.5 mm, about 1.6 mm, about 1.7 mm, about 1.8 mm, about 1.9 mm, about 2.0 mm, about 2.1 mm, about 2.2 mm, about 2.4 mm, about 2.6 mm, about 2.8 mm, about 3.0 mm, about 3.5 mm, about 4.0 mm, about 4.5 mm, about 5.0 mm, about 5.5 mm, about 6.0 mm, about 6.5 mm, about 7.0 mm, about 7.5 mm, about 8.0 mm, about 8.5 mm, about 9.0 mm, about 9.5 mm, about 10.0 mm, or any range including and between any two of these values.

[0058] The process may include at least at least one of the following features:

(1) one or more catalyst carriers with an average annular space between the outer surface of the skirt and the inner tube wall from about 1.3 mm to about 2.6 mm; and

(2) a total combined surface area of the outer surface of the skirt and the inner tube wall, and optionally a heat transfer structure, per volume of the solid synthesis catalyst in the catalyst carrier (the "SA/V") from about 400 m ² /m ³ to about 40,000 m ² /m ³ ;

wherein the heat transfer structure includes at least one of (i) a plurality of fins extending radially from the perforated inner wall to the perforated outer wall;

(ii) a plurality of fins extending radially from the perforated inner wall to an inner surface of the skirt; and

(iii) a network of heat conducting surfaces in conductive thermal contact with a portion the inner tube wall and a portion of the outer surface of the skirt.

[0059] Where the solid synthesis catalyst is a monolithic synthesis catalyst, it may be the monolithic synthesis catalyst has a channel extending longitudinally through the monolithic synthesis catalyst on a central axis of the monolithic synthesis catalyst, where the monolithic catalyst has a monolithic upper surface disposed toward the reactor inlet, a monolithic lower surface disposed toward the reactor outlet, a monolithic inner surface defined by the channel, and a monolithic outer surface disposed toward the inner tube wall. With such a monolithic synthesis catalyst, the channel may be sized such that the monolithic inner surface fits around the perforated inner wall. However, with such a monolithic synthesis catalyst, the catalyst carrier may omit the perforated inner wall. In an outward flow configuration, while the top surface of the catalyst carrier serves to ensure that the gaseous stream enters the channel for optimal radial flow, the outward flow catalyst carrier with a monolithic synthesis catalyst may omit a portion of, or all of, the top surface. If all of the top surface is omitted, the seal is located at the monolithic upper surface, preferably attached to the outer wall. In an inward flow configuration, the inward flow catalyst carrier with a monolithic synthesis catalyst may omit a portion of, or all of, the bottom surface. If all of the bottom surface is omitted, the seal is located at the monolithic lower surface, preferably attached to the outer wall.

[0060] The monolithic synthesis catalyst may not include a channel. In such embodiments, the catalyst container may not include a central tube and the outer wall may not be perforated. In outward flow configurations of such embodiments, feet on the upper face of the bottom surface are included such that, in use, the catalyst monolith is supported on the feet and there is a space between the monolithic lower surface and the upper face of the bottom surface of the catalyst carrier (an "outward flow monolithic gap") that is in flow communication with the annulus between the skirt and the outer wall of the annular container. The catalyst carrier preferably omits a portion of, or all of, the top surface. In such outward flow configurations, the bottom surface may be connected to the outer wall by two or more connections or the bottom surface may be connected to the outer wall by a perforated material, and the skirt extends from this connection at the bottom surface. The two or more connections may have a width of about 1 to about 20 mm (or any range including and between any two integers thereof). In inward flow configurations of embodiments where the monolithic synthesis catalyst does not include a channel, feet on the lower face of the top surface are included such that, in use, there is a space between the monolithic lower surface and the lower face of the top surface of the catalyst carrier (an "inward flow monolithic gap") that is in flow communication with the annulus between the skirt and the outer wall of the annular container. The catalyst carrier preferably omits a portion of, or all of, the bottom surface. In such inward flow configurations, the top surface may be connected to the outer wall by two or more connections or the top surface may be connected to the outer wall by a perforated material, and the skirt extends from this connection at the top surface. The two or more connections may have a width of about 1 to about 20 mm (or any range including and between any two integers thereof). The inward flow monolithic gap and outward flow monolithic gap may each independently be from about 2.0 mm to about 15.0 mm; thus, the monolithic gap may be about 2.0 mm, about 2.1 mm, about 2.2 mm, about 2.4 mm, about 2.6 mm, about 2.8 mm, about 3.0 mm, about 3.5 mm, about 4.0 mm, about 4.5 mm, about 5.0 mm, about 5.5 mm, about 6.0 mm, about 6.5 mm, about 7.0 mm, about 7.5 mm, about 8.0 mm, about 8.5 mm, about 9.0 mm, about 9.5 mm, about 10.0 mm, about 10.5 mm, about 11.0 mm, about 11.5 mm, about 12.0 mm, about 12.5 mm, about 13.0 mm, about 13.5 mm, about 14.0 mm, about 14.5 mm, about 15.0 mm, or any range including and between any two of these values.

[0061] It is to be understood that the average annular space may be different for each catalyst carrier. For example, a first carrier group may include at least two catalyst carriers with the average annular space from about 1.3 mm to about 3.0 mm, and a second carrier group of at least two catalyst carriers with an average annular space larger than the first carrier group and that is between about 1.5 mm to about 10.0 mm. In any embodiment including a second carrier group, the catalyst carriers of the second carrier group may include an average annular space between about 3.0 mm to about 10.0 mm. The catalyst carriers of the second carrier group may contain a solid synthesis catalyst with higher activity than the solid synthesis catalyst contained by the catalyst carriers of the first carrier group. For example, it may be that the activity of the solid synthesis catalyst of the second carrier group is at least 5% higher, at least 10% higher, at least 20% higher, at least 30% higher, at least 40% higher, or at least 50% higher than the activity of the solid synthesis catalyst of the first carrier group. It is further understood that there may be a plurality of catalyst carrier groups, where each carrier group may contain a solid synthesis catalyst of differing activity.

[0062] In regard to the reactor tube(s), each reactor tube may include from 10 to about 500 catalyst carriers within the reactor tube, preferably about 20 to about 200 catalyst carriers within the reactor tube, more preferably about 40 to about 200 catalyst carriers within the reactor tube, even more preferably about 80 to about 200 catalyst carriers within the reactor tube and/or about 40 to about 160 catalyst carriers within the reactor tube, such as from about 100 to about 160 catalyst carriers. The reactor tube(s) may each independently be from about 1 meter (m) to about 20 m in length. The one or more reactor tubes may each independently have a diameter of about 40 mm to about 150 mm, and may each

independently have a diameter of about 40 mm to about 65 mm. Thus, the one or more reactor tubes may each independently have a diameter of about 40 mm, about 45 mm, about 50 mm, about 55 mm, about 60 mm, about 65 mm, about 70 mm, about 75 mm, about 80 mm, about 85 mm, about 90 mm, about 95 mm, about 100 mm, about 110 mm, about 120 mm, about 130 mm, about 140 mm, about 150 mm, or any range including and between any two of these values.

[0063] The tubular reactor may contain from 1 to about 20,000 reactor tubes. It may be the tubular reactor includes at least 100 reactor tubes. The tubular reactor may preferably include less than about 10,000 reactor tubes, more preferably less than about 5,000 reactor tubes, and even more preferably less than about 2,200 reactor tubes. It is especially preferred if the tubular reactor includes less than about 2,000 reactor tubes. The reactor tubes of the tubular reactor preferably operate in parallel. Thus, the process may be performed in such reactor tubes running in parallel, where increasing the number of individual reactor tubes may raise the capacity and/ or throughput for the process. Furthermore, the inner tube wall of the reactor rubes may include shapes to ensure thermal contact between the skirt and the inner tube wall, and/or enhance heat transfer between the fluid exiting the catalyst carrier and the inner tube wall. Such shapes for ensuring thermal contact and/or enhanced heat transfer may include channels, indented features, and/or protruded features that increase surface area for heat transfer and/or increase the convective heat transfer coefficient to remove heat from the exiting fluid. The shapes may be arranged longitudinally along the length of one or more reactor tubes. Such shapes include (when viewed as a cross-section as depicted in FIG. 1) square, rectangle, and triangular.

[0064] In a related aspect, a process is provided for the conversion of synthesis gas.

The process includes contacting in a tubular reactor a gaseous stream with a solid synthesis catalyst to produce a synthetic product. The gaseous stream includes synthesis gas. The tubular reactor of the process includes (a) a reactor inlet in fluid communication with one or more reactor tubes wherein each reactor tube includes a tube inlet, a tube outlet located downstream of the tube inlet, an inner tube wall, an outer tube wall, and one or more catalyst carriers within the reactor tube, (b) a reactor outlet located downstream of the reactor inlet in fluid communication with the one or more reactor tubes, and (c) a cooling medium in contact with the one or more reactor tubes. The process may include introducing the gaseous stream through the reactor inlet at a pressure from about 50 psig to about 1,000 psig with a ratio of H ₂ /CO in the synthesis gas from about 1.6 to about 2.5, preferably a pressure from about 250 psig to about 1,000 psig with a ratio of H ₂ /CO in the synthesis gas from about 1.6 to about 2.0. The pressure for introducing the gaseous stream through the reactor inlet may be about 50 psig, about 100 psig, about 150 psig, about 200 psig, about 250 psig, about 300 psig, about 350 psig, about 400 psig, about 450 psig, about 500 psig, about 550 psig, about 600 psig, about 650 psig, about 700 psig, about 750 psig, about 800 psig, about 850 psig, about 900 psig, about 950 psig, about 1,000 psig, or any range including and between any two of these values. For example, the process may preferably include introducing the gaseous stream through the reactor inlet at a pressure from about 500 psig to about 1,000 psig, more preferably at a pressure from about 500 psig to about 750 psig. The ratio of H ₂ /CO in the synthesis gas may be about 1.6, about 1.7, about 1.8, about 1.9, about 2.0, about 2.1, about 2.2, about 2.3, about 2.4, about 2.5, or any range including and/or in between any two of these values. [0065] The catalyst carrier within the reactor tube may be of the "outward flow" configuration or the "inward flow" configuration described previously. For monolithic synthesis catalysts, the catalyst carrier may be, but is not required to be, according to any one of the configurations for monolithic synthesis catalysts described herein.

[0066] The process includes at least one of the following features:

(1) an average annular space between the outer surface of the skirt and the inner tube wall from about 1.3 mm to about 2.6 mm; and

(2) a total combined surface area of the outer surface of the skirt and the inner tube wall, and optionally a heat transfer structure, per volume of the solid synthesis catalyst in the catalyst carrier (the "SA/V") from about 400 m ² /m ³ to about 40,000 m ² /m ³ ;

wherein the heat transfer structure includes at least one of

(i) a plurality of fins extending radially from the perforated inner wall to the perforated outer wall;

(ii) a plurality of fins extending radially from the perforated inner wall to an inner surface of the skirt; and

(iii) a network of heat conducting surfaces in conductive thermal contact with a portion the inner tube wall and a portion of the outer surface of the skirt.

Thus, it may be that an average annular space between the outer surface of the skirt that is closest to the seal and the inner surface of the tube wall is from about 1.3 mm to about 2.6 mm in one or more of the catalyst carriers. The average annular space may be about 1.3 m, about 1.4 mm, about 1.5 mm, about 1.6 mm, about 1.7 mm, about 1.8 mm, about 1.9 mm, about 2.0 mm, about 2.1 mm, about 2.2 mm, about 2.3 mm, about 2.4 mm, about 2.5 mm, about 2.6 mm, or any range including and between any two of these values. Thus, the average annular space may preferably be from about 1.5 mm to about 2.0 mm.

[0067] It is to be understood that the average annular space may be different for each catalyst carrier, and it may be for processes including (1) above that not all catalyst carriers have an annular space of about 1.3 mm to about 2.6 mm. A first carrier group may include at least two catalyst carriers with the average annular space from about 1.3 mm to about 2.6 mm, and a second carrier group of at least two catalyst carriers with an average annular space larger than the first carrier group and that is between about 1.5 mm to about 10.0 mm. In any embodiment including a second carrier group, the catalyst carriers of the second carrier group may include an average annular space between about 3.0 mm to about 10.0 mm. The catalyst carriers of the second carrier group may contain a solid synthesis catalyst with higher activity than the solid synthesis catalyst contained by the catalyst carriers of the first carrier group. For example, it may be that the activity of the solid synthesis catalyst of the second carrier group is at least 5% higher, at least 10% higher, at least 20% higher, at least 30% higher, at least 40% higher, or at least 50% higher than the activity of the solid synthesis catalyst of the first carrier group. It is further understood that there may be a plurality of catalyst carrier groups, where each carrier group may contain a solid synthesis catalyst of differing activity.

[0068] In regard to the reactor tube(s), each reactor tube may include from 10 to about 500 catalyst carriers within the reactor tube, preferably about 20 to about 200 catalyst carriers within the reactor tube, more preferably about 40 to about 200 catalyst carriers within the reactor tube, even more preferably about 80 to about 200 catalyst carriers within the reactor tube and/or about 40 to about 160 catalyst carriers within the reactor tube, such as from about 100 to about 160 catalyst carriers. The reactor tube(s) may each independently be from about 1 meter (m) to about 20 m in length. The one or more reactor tubes may each independently have a diameter of about 40 mm to about 300 mm, and may each

independently have a diameter of about 40 mm to about 65 mm. Thus, the one or more reactor tubes may each independently have a diameter of about 40 mm, about 45 mm, about 50 mm, about 55 mm, about 60 mm, about 65 mm, about 70 mm, about 75 mm, about 80 mm, about 85 mm, about 90 mm, about 95 mm, about 100 mm, about 110 mm, about 120 mm, about 130 mm, about 140 mm, about 150 mm, about 160 mm, about 170 mm, about 180 mm, about 190 mm, about 200 mm, about 220 mm, about 240 mm, about 260 mm, about 280 mm, about 300 mm, or any range including and between any two of these values.

[0069] The tubular reactor may contain from 1 to about 20,000 reactor tubes. The tubular reactor may include at least 100 reactor tubes. The tubular reactor may preferably include less than about 10,000 reactor tubes, more preferably less than about 5,000 reactor tubes, and even more preferably less than about 2,200 reactor tubes. It is especially preferred if the tubular reactor includes less than about 2,000 reactor tubes. The reactor tubes of the tubular reactor preferably operate in parallel. Thus, the process may be performed in such reactor tubes running in parallel, where increasing the number of individual reactor tubes may raise the capacity and/ or throughput for the process. Furthermore, the inner tube wall of the reactor rubes may include shapes to ensure thermal contact between the skirt and the inner tube wall, and/or enhance heat transfer between the fluid exiting the catalyst carrier and the inner tube wall. Such shapes for ensuring thermal contact and/or enhanced heat transfer may include channels, indented features, and/or protruded features that increase surface area for heat transfer and/or increase the convective heat transfer coefficient to remove heat from the exiting fluid. The shapes may be arranged longitudinally along the length of one or more reactor tubes. Such shapes include (when viewed as a cross-section as depicted in FIG. 1) square, rectangle, and triangular.

[0070] In a related aspect, a process is provided for the conversion of synthesis gas where the process includes contacting in a tubular reactor a gaseous stream with a solid synthesis catalyst to produce a synthetic product. The gaseous stream includes synthesis gas. The tubular reactor of the process includes (a) a reactor inlet in fluid communication with one or more reactor tubes wherein each reactor tube includes a tube inlet, a tube outlet located downstream of the tube inlet, an inner tube wall, an outer tube wall, and one or more catalyst carriers within the reactor tube, (b) a reactor outlet located downstream of the reactor inlet in fluid communication with the one or more reactor tubes, and (c) a cooling medium in contact with the one or more reactor tubes. The one or more reactor tubes each independently have a diameter of about 40 mm to about 65 mm, and the process may include introducing the gaseous stream through the reactor inlet at a pressure from about 50 psig to about 1,000 psig with a ratio of H ₂ /CO in the synthesis gas from about 1.6 to about 2.5, preferably at a pressure from about 250 psig to about 1,000 psig with a ratio of H ₂ /CO in the synthesis gas from about 1.6 to about 2.0. The diameter of the one or more reactor tubes may be about 40 mm, about 45 mm, about 50 mm, about 55 mm, about 60 mm, about 65 mm, or any range including and between any two of these values. Similarly, the pressure for introducing the gaseous stream through the reactor inlet may be about 50 psig, about 100 psig, about 150 psig, about 200 psig, about 250 psig, about 300 psig, about 350 psig, about 400 psig, about 450 psig, about 500 psig, about 550 psig, about 600 psig, about 650 psig, about 700 psig, about 750 psig, about 800 psig, about 850 psig, about 900 psig, about 950 psig, about 1,000 psig, or any range including and between any two of these values. For example, the process may preferably include introducing the gaseous stream through the reactor inlet at a pressure from about 500 psig to about 1,000 psig, more preferably at a pressure from about 500 psig to about 750 psig. The ratio of H ₂ /CO in the synthesis gas may be about 1.6, about 1.7, about 1.8, about 1.9, about 2.0, about 2.1, about 2.2, about 2.3, about 2.4, about 2.5, or any range including and/or in between any two of these values.

[0071] The catalyst carrier within the reactor tube may be of the "outward flow" configuration or an "inward flow" configuration described previously. For monolithic synthesis catalysts, the catalyst carrier may be, but is not required to be, according to any one of the configurations for monolithic synthesis catalysts described herein. For each catalyst carrier independently, an average annular space between the outer surface of the skirt that is closest to the seal and the inner surface of the tube wall is from about 1.1 mm to about 10.0 mm. The average annular space for each catalyst carrier may independently be about 1.1 mm, about 1.2 mm, about 1.3 mm, about 1.4 mm, about 1.5 mm, about 1.6 mm, about 1.7 mm, about 1.8 mm, about 1.9 mm, about 2.0 mm, about 2.1 mm, about 2.2 mm, about 2.4 mm, about 2.6 mm, about 2.8 mm, about 3.0 mm, about 3.5 mm, about 4.0 mm, about 4.5 mm, about 5.0 mm, about 5.5 mm, about 6.0 mm, about 6.5 mm, about 7.0 mm, about 7.5 mm, about 8.0 mm, about 8.5 mm, about 9.0 mm, about 9.5 mm, about 10.0 mm, or any range including and between any two of these values.

[0072] It is to be understood that the average annular space may be different for each catalyst carrier. For example, a first carrier group may include at least two catalyst carriers with the average annular space from about 1.3 mm to about 3.0 mm, and a second carrier group of at least two catalyst carriers with an average annular space larger than the first carrier group and between about 1.5 mm to about 10.0 mm. In any embodiment including a second carrier group, the catalyst carriers of the second carrier group may include an average annular space between about 3.0 mm to about 10.0 mm. The catalyst carriers of the second carrier group may contain a solid synthesis catalyst with higher activity than the solid synthesis catalyst contained by the catalyst carriers of the first carrier group. For example, it may be that the activity of the solid synthesis catalyst of the second carrier group is at least 5% higher, at least 10% higher, at least 20% higher, at least 30% higher, at least 40% higher, or at least 50% higher than the activity of the solid synthesis catalyst of the first carrier group. It is further understood that there may be a plurality of catalyst carrier groups, where each carrier group may contain a solid synthesis catalyst of differing activity.

[0073] Further, the process may include a heat transfer structure comprising at least one of

(1) an average annular space between the outer surface of the skirt and the inner tube wall from about 1.3 mm to about 2.6 mm; and

(2) a total combined surface area of the outer surface of the skirt and the inner tube wall, and optionally a heat transfer structure, per volume of the solid synthesis catalyst in the catalyst carrier (the "SA/V") from about 400 m ² /m ³ to about 40,000 m ² /m ³ ;

wherein the heat transfer structure includes at least one of

(i) a plurality of fins extending radially from the perforated inner wall to the perforated outer wall;

(ii) a plurality of fins extending radially from the perforated inner wall to an inner surface of the skirt; and

(iii) a network of heat conducting surfaces in conductive thermal contact with a portion the inner tube wall and a portion of the outer surface of the skirt.

[0074] In regard to the reactor tube(s), each reactor tube may include from 10 to about 500 catalyst carriers within the reactor tube, preferably about 20 to about 200 catalyst carriers within the reactor tube, more preferably about 40 to about 200 catalyst carriers within the reactor tube, even more preferably about 80 to about 200 catalyst carriers within the reactor tube and/or about 40 to about 160 catalyst carriers within the reactor tube, such as from about 100 to about 160 catalyst carriers. The reactor tube(s) may be from about 1 m to about 20 m in length. As noted above, the one or more reactor tubes may each independently have a diameter of about 40 mm to about 65 mm. Thus, the one or more reactor tubes may each independently have a diameter of about 40 mm, about 45 mm, about 50 mm, about 55 mm, about 60 mm, about 65 mm, or any range including and between any two of these values.

[0075] The tubular reactor may contain from 1 to about 20,000 reactor tubes. It may be the tubular reactor includes at least 100 reactor tubes. The tubular reactor may preferably include less than about 10,000 reactor tubes, more preferably less than about 5,000 reactor tubes, and even more preferably less than about 2,200 reactor tubes. It is especially preferred if the tubular reactor includes less than about 2,000 reactor tubes. The reactor tubes of the tubular reactor preferably operate in parallel. Thus, the process may be performed in such reactor tubes running in parallel, where increasing the number of individual reactor tubes may raise the capacity and/ or throughput for the process. Furthermore, the inner tube wall of the reactor rubes may include shapes to ensure thermal contact between the skirt and the inner tube wall, and/or enhance heat transfer between the fluid exiting the catalyst carrier and the inner tube wall. Such shapes for ensuring thermal contact and/or enhanced heat transfer may include channels, indented features, and/or protruded features that increase surface area for heat transfer and/or increase the convective heat transfer coefficient to remove heat from the exiting fluid. The shapes may be arranged longitudinally along the length of one or more reactor tubes. Such shapes include (when viewed as a cross-section as depicted in FIG. 1) square, rectangle, and triangular.

Features Related to all Aspects

[0076] By way of a non-limiting explanation, in outward radial flow processes of the present technology, introducing the gaseous stream through the reactor inlet allows for the gaseous stream (and the synthesis gas therein) to pass downwardly through the one or more reactor tubes to the top surface of a catalyst carrier, where the gaseous stream passes into the central tube before passing through the perforated inner wall of the catalyst carrier, contacting the solid synthesis catalyst. See FIG. 2A (discussed again in the Examples). In contacting the solid synthesis catalyst, at least a portion of the synthesis gas is reacted. The gaseous stream (now containing both reacted and unreacted synthesis gas) then passes through the perforated outer wall and then upwardly between an inner surface of the skirt and the perforated outer wall of the annular container until the gaseous stream reaches the seal. The "inner surface" of the skirt is the surface oriented toward the perforated outer wall of the annular container. Upon reaching the seal, the gaseous stream is directed over an end of the skirt and flows downwardly between the outer surface of the skirt and the inner surface of the tube wall. The steps repeat for each catalyst carrier that may be in each reactor tube. Upon exiting the final catalyst carrier in a reactor tube, the product (which includes hydrocarbons) exits the reactor outlet. For the avoidance of doubt, any discussion of orientation herein, for example terms such as upwardly, below, lower, and the like have, for ease of reference been discussed with regard to the orientation of the catalyst carrier as it relates to the location of the reactor inlet (e.g., "top" is disposed toward the reactor inlet; downwardly flows away from the reactor inlet) and as further illustrated in the accompanying drawings. However, where the tubes, and hence the catalyst carrier, are used in an alternative orientation, the terms should be construed accordingly.

[0077] For example, where the reactor inlet is located lower than the reactor inlet such that the gaseous stream is flowing against gravity (an "upflow reactor" design), the gaseous stream passes counter to gravity through the one or more reactor tubes to the top surface (disposed toward the reactor inlet) of a catalyst carrier, whereafter the gaseous stream passes into the central tube before passing through the perforated inner wall of the catalyst carrier, contacting the solid synthesis catalyst. With this understanding, terms "upwardly" and "downwardly" as well as "top" and "bottom" are construed in relation to the reactor inlet. An upflow reactor may provide enhanced heat transfer from the gaseous stream as it flows between the outer surface of the skirt and the inner surface of the tube wall due to increasing the average flow rate of the gaseous stream and increased liquid hold up at the heat exchange wall in comparison with a "downflow" reactor where the gaseous stream flows in the direction of gravitational pull.

[0078] By way of a non-limiting explanation regarding the inward radial flow processes of the present technology, introducing the gaseous stream through the reactor inlet allows for the gaseous stream (and the synthesis gas therein) to pass downwardly through the one or more reactor tubes to the top surface of a catalyst carrier, where the gaseous stream is directed to flow downwardly between the outer surface of the skirt and the inner surface of the tube wall until the gaseous stream reaches the seal (FIG. 2B). Upon reaching the seal, the gaseous stream is directed over the end of the skirt and flows upwardly between an inner surface of the skirt and the perforated outer wall where the gaseous stream is directed through the perforated outer wall and contacts the solid synthesis catalyst. In contacting the solid synthesis catalyst, at least a portion of the synthesis gas is reacted. The gaseous stream (now containing both reacted and unreacted synthesis gas) then passes through the perforated inner wall of the catalyst carrier and passes into the central tube, where it passes downwardly and exits the catalyst carrier. The steps repeat for each catalyst carrier that may be in each reactor tube in the inward radial flow reactor. Upon exiting the final catalyst carrier in a reactor tube, a product that includes the synthetic product exits the reactor outlet. As described for the outward radial flow reactor, for the avoidance of doubt any discussion of orientation herein, for example terms such as upwardly, below, lower, top, bottom, and the like have, for ease of reference, been discussed with regard to the orientation of the catalyst carrier as it relates to the location of the reactor inlet and as further illustrated in the accompanying drawings.

However, where the tubes, and hence the catalyst carrier, are used in an alternative orientation, the terms should be construed accordingly.

[0079] Similar to the outward radial flow reactor design, where the reactor inlet in the inward flow reactor is located lower than the reactor inlet such that the gaseous stream is flowing against gravity (an "upflow-inward flow reactor" design), the gaseous stream passes counter to gravity through the one or more reactor tubes to the top surface (disposed toward the reactor inlet) of a catalyst carrier. With this understanding, terms "upwardly" and

"downwardly" are construed in relation to the reactor inlet. An upflow reactor may provide enhanced heat transfer from the gaseous stream as it flows between the outer surface of the skirt and the inner surface of the tube wall due to increasing the average flow rate of the gaseous stream and increased liquid hold up at the heat exchange wall in comparison with a "downflow" inward flow reactor where the gaseous stream flows in the direction of gravitational pull.

[0080] Where a monolithic synthesis catalyst with a channel is employed in an outward flow configuration and the catalyst carrier omits a portion, or all, of the top surface, the gaseous stream first contacts the monolithic upper surface of the monolithic synthesis catalyst. Because the flow of the gaseous stream is impeded by the monolithic synthesis catalyst, it will generally take the easier path and enter the channel. Upon entering the channel of the monolithic synthesis catalyst, the gaseous stream passes radially through the monolithic synthesis catalyst towards the monolithic outer surface and to the skirt. In inward flow configurations where a monolithic synthesis catalyst within a channel is employed, the gaseous stream passes radially through the monolithic outer surface towards the monolithic inner surface. If a perforated inner wall is employed in the catalyst carrier, the gaseous stream passes through the monolithic inner surface and then through the perforated inner wall as described previously for the inward flow configurations; if a perforated inner wall is omitted, then the gaseous stream passes through the monolithic inner surface and passes into the channel, where it passes downwardly and exits the catalyst carrier.

[0081] Where a monolithic synthesis catalyst is employed that does not include a channel, the gaseous stream contacts the monolithic upper surface and flows through the monolithic synthesis catalyst in a direction parallel to the axis of the reactor tube. In an outward flow configuration, once the gaseous stream reaches the bottom surface of the catalyst carrier it is directed towards the skirt of the carrier; in an inward flow configuration, once the gaseous steam passes through the monolithic lower surface it effectively exits the catalyst carrier.

[0082] Each catalyst carrier will generally be sized such that it is of a smaller dimension than the internal dimension of the reactor tube into which it is placed. The seal is sized such that it interacts with the inner wall of the reactor tube when the catalyst carrier is in position within the reactor tube. The seal need not be perfect, provided that it is effective to cause the majority of the flowing gas to pass through, rather than around, the catalyst carrier. Further, the liquid or wax product produced during the process may effectively act to block small passages between the seal and the inner wall of the reactor tube, thus further ensuring the gaseous stream flows according to the intended flow path.

[0083] As discussed above, the process of any embodiment herein may include a first carrier group and a second carrier group, or a plurality of catalyst carrier groups. As a further example, in any embodiment herein, a first carrier group may include at least two catalyst carriers with the average annular space from about 1.5 mm to about 10.0 mm, and a second carrier group (downstream of the first carrier group) of at least two catalyst carriers with an average annular space smaller than the first carrier group and that is between about 1.3 mm to about 2.6 mm. A third carrier group may be included that is downstream of the second carrier group and includes at least two catalyst carriers with an average annular space larger than the second carrier group and that is between about 1.5 mm to about 10.0 mm. It is to be understood that the annular spaces of these carrier groups may independently be any subrange of the indicated ranges as described herein. The catalyst carriers of the first carrier group may contain a solid synthesis catalyst with higher activity than the solid synthesis catalyst contained by the catalyst carriers of the second carrier group. For example, it may be that the activity of the solid synthesis catalyst of the first carrier group is at least 5% higher, at least 10% higher, at least 20% higher, at least 30% higher, at least 40% higher, or at least 50%) higher than the activity of the solid synthesis catalyst of the second carrier group. When a third carrier group is employed, the third carrier group may contain a solid synthesis catalyst with higher activity than the solid synthesis catalyst contained by the catalyst carriers of the second carrier group; it may be that the activity of the solid synthesis catalyst of the third carrier group is at least 5% higher, at least 10% higher, at least 20% higher, at least 30% higher, at least 40% higher, or at least 50% higher than the activity of the solid synthesis catalyst of the second carrier group. As another example, in any embodiment herein, a first carrier group may include at least two catalyst carriers, and a second carrier group

(downstream of the first carrier group) of at least two catalyst carriers, where the first carrier group contains a solid synthesis catalyst with higher activity than the solid synthesis catalyst contained by the catalyst carriers of the second carrier group. The annular space may be of any dimension described herein. A third carrier group may be included that is downstream of the second carrier group and includes at least two catalyst carriers, where the third carrier group contains a solid synthesis catalyst with higher activity than the solid synthesis catalyst contained by the catalyst carriers of the second carrier group. The activity of the solid synthesis catalyst of the first and third carrier group may each independently be at least 5% higher, at least 10% higher, at least 20% higher, at least 30% higher, at least 40% higher, or at least 50% higher than the activity of the solid synthesis catalyst of the second carrier group. [0084] In embodiments of the outward flow reactor including a plurality of catalyst carriers, the gaseous stream (including unreacted synthesis gas and/or product) flows between the outer surface of the skirt of a first carrier and the inner surface of the reactor tube until the gaseous stream contacts the top surface and seal of a second carrier and is directed into the central tube of the second carrier. This flow path described for outward flow reactors is then repeated for each catalyst carrier.

[0085] In embodiments of the inward flow reactor including a plurality of catalyst carriers, the gaseous stream (including unreacted synthesis gas and/or product) flows out of the central tube until the gaseous stream contacts the upper surface of a second carrier and is directed to flow downwardly between the outer surface of the skirt and the inner surface of the tube wall until the gaseous stream reaches the seal. The flow path described for inward flow reactors is then repeated for each catalyst carrier.

[0086] Catalyst carriers may be formed of any suitable material, generally selected to withstand the operating conditions of the reactor. For example, the catalyst carrier may be fabricated from carbon steel, aluminum, stainless steel, copper, ceramics, other alloys or any material able to withstand the reaction conditions. Where a heat transfer structure as described herein is employed, the catalyst carrier is preferably fabricated from carbon steel, aluminum, stainless steel, copper, other alloys or any material able to withstand the reaction conditions, more preferably from steel, aluminum, copper, an alloy thereof, or a combination of any two or more thereof. For example, the catalyst carrier may include a combination of steel and aluminum, steel and copper, aluminum and copper, or steel, aluminum, and copper. Likewise, the wall of the annular container can be of any suitable thickness, such as about 0.3 mm to about 5.0 mm. The size of the perforations in the inner and outer walls of the annular container will be selected such as to allow uniform flow of the gaseous stream and/or product through the solid synthesis catalyst while maintaining the catalyst within the container. By way of a non-limiting example, when the solid synthesis catalyst includes a particulate Fischer-Tropsch catalyst, the size of the perforations will depend on the size of the particulate Fischer- Tropsch catalyst particles being used. Examples of perforations include sintered porous metals, small apertures of any suitable configuration, such as slots, formed by a wire screen, or by any other means of creating a porous or permeable surface allowing flow of reactants and products while preventing migration of solid catalyst particles. By way of illustration, openings in the inner and outer walls of the annular container may be sized such that they are larger than the weight average diameter of a particulate Fischer-Tropsch catalyst but have an additional filter mesh covering the perforations to ensure the particulate Fischer- Tropsch catalyst is maintained within the annular container. This enables larger openings to be used to facilitate the free movement of reactants without a significant loss of process pressure.

[0087] Although the top surface closing the annular container will generally be located at the upper edge of, e.g., each wall of the annular container, it may be desirable to locate the top surface below the upper edge such that a portion of the upper edge of the outer wall forms a lip. Similarly, the bottom surface may be located at the lower edge of the, or each, wall of the annular container or may be desirable to locate the bottom surface such that it is above the bottom edge of the wall of the annular container such that the wall forms a lip.

[0088] In the outward flow catalyst carriers, the bottom surface of the annular container and the surface closing the bottom of the central tube may be formed as a single unit or they may be two separate pieces connected together. The two surfaces may be coplanar or may be in different planes. For example, the surface closing the bottom of the central tube may be in a lower plane than the bottom surface of the annular container, which assists in the location of one carrier on to a carrier arranged below it when using a plurality of catalyst carriers. However, the surface closing the bottom of the central tube may be in a higher plane that the bottom surface of the annular container. The bottom surface may include one or more drain holes. Where one or more drain holes are present, such holes may be covered by a filter mesh. A drain hole, optionally covered with a filter mesh, may also be present in the surface closing the bottom of the tube. A filter mesh would be more readily filled with liquids due to capillary forces and thus enable liquid draining and minimize gaseous by-pass. In embodiments of the processes where the outward flow catalyst carrier is disposed in a non-vertical orientation, one or more drain holes may be located at the lowest points in the catalyst carrier (determined by how the catalyst carrier is disposed during the process). One or more spacers may extend downwardly from the bottom surface of the annular container and/or upwardly from the top surface of the annular container. The spacer(s) may be formed as separate components from the catalyst carrier or may be formed by depressions in the bottom surface/top surface. Spacers assist in providing a clear path for the reactants and products flowing between the bottom surface of, e.g., a first carrier and the top surface of a second lower carrier in use. As depressions, the spacer(s) may be about 3 mm to about 50 mm deep, preferably about 3 mm to about 25 mm deep; as components extending outwardly from the annular container, the spacer(s) may be about 3 mm to about 50 mm in length, more preferably about 3 mm to about 25 mm in length.

[0089] Similarly in the inward flow catalyst carriers, the top surface of the annular container and the surface closing the top of the central tube may be formed as a single unit or they may be two separate pieces connected together. The two surfaces may be coplanar or may be in different planes. For example, the surface closing the top of the central tube may be in a higher plane than the top surface of the annular container, which assists in the location of one carrier on to a carrier arranged below it when using a plurality of catalyst carriers. However, the surface closing the top of the central tube may be in a lower plane than the top surface of the annular container. The bottom surface may include one or more drain holes. Where one or more drain holes are present, such holes may be covered by a filter mesh. In embodiments of the processes where the inward flow catalyst carrier is disposed in a non- vertical orientation, one or more drain holes may be located at the lowest points in the catalyst carrier (determined by how the catalyst carrier is disposed during the process). One or more spacers may extend downwardly from the bottom surface of the annular container and/or upwardly from the top surface of the annular container. The spacer(s) may be formed as separate components from the catalyst carrier or may be formed by depressions in the bottom surface/top surface. Spacers assist in providing a clear path for the reactants and products flowing between the bottom surface of, e.g., a first carrier and the top surface of a second lower carrier in use. The spacers may also allow a closer approach of the temperature of the fluids inside the tube to the temperature of the cooling fluid outside the tube. As depressions, the spacer(s) may be about 3 mm to about 50 mm deep, preferably about 3 mm to about 25 mm deep; as components extending outwardly from the annular container, the spacer(s) may be about 3 mm to about 50 mm in length, preferably about 3 mm to about 25 mm in length. [0090] The skirt of the catalyst carrier(s) may be smooth or it may be shaped.

Suitable shapes include, but are not limited to, pleats, corrugations, and the like. Such shapes may act to ensure thermal contact between the skirt and the inner tube wall, and/or enhance heat transfer between the fluid exiting the catalyst carrier and the tube wall. Such shapes for ensuring thermal contact and/or enhanced heat transfer may include channels, indented features, and/or protruded features that increase surface area for heat transfer and/or increase the convective heat transfer coefficient to remove heat from the hot fluid exiting the catalyst carrier. The shapes may be arranged longitudinally along the length of the carrier. Such shapes include (when viewed as a cross-section as depicted in FIG. 1) square, rectangle, and triangular. Shaped skirts provide increased surface area and assist with the insertion of the catalyst carrier into the reaction tube since it will allow any surface roughness on the inner surface of the reactor tube or differences in tolerances in tubes to be accommodated. Where the skirt is shaped, it may be flattened to a smooth configuration towards the point at which it is connected to the annular container to allow a gas seal with the annular container. The skirt may be connected to the outer wall of the annular container at or near the base of the annular container for outward flow catalyst carriers. The skirt is generally fabricated from the same materials as the catalyst carrier, or may particularly include a highly conductive metal such as copper, aluminum, a combination of both, or an alloy of either.

[0091] Where the skirt is connected at a point above the bottom of the wall in the outward flow catalyst carriers, the wall may be free of perforations in the area below the point of connection. The skirt may be connected to the outer wall of the annular container at or near the top of inward flow catalyst carriers and the wall may be free of perforations in the area below the point of connection. In any embodiment herein, the skirt may be flexible. Generally, the skirt will stop at about 0.15 cm to about 2.5 cm, preferably about 1 cm, short of the top surface of the annular container in outward flow catalyst carriers, and will stop at about 0.5 cm to about 1.5 cm, preferably about 1 cm, short of the bottom surface of the annular container in outward flow catalyst carriers.

[0092] In any embodiment herein, the size of the annulus between the inner surface of the skirt and the perforated outer wall of the annular container (the "inner gap") may be selected to accommodate the gas flow rate required while maintaining high heat transfer and low pressure drop. The average inner gap between the skirt and the outer wall of the annular container may be from about 2.0 mm to about 15.0 mm. The average inner gap for each catalyst carrier may independently be about 2.0 mm, about 2.1 mm, about 2.2 mm, about 2.4 mm, about 2.6 mm, about 2.8 mm, about 3.0 mm, about 3.5 mm, about 4.0 mm, about 4.5 mm, about 5.0 mm, about 5.5 mm, about 6.0 mm, about 6.5 mm, about 7.0 mm, about 7.5 mm, about 8.0 mm, about 8.5 mm, about 9.0 mm, about 9.5 mm, about 10.0 mm, about 10.5 mm, about 11.0 mm, about 11.5 mm, about 12.0 mm, about 12.5 mm, about 13.0 mm, about 13.5 mm, about 14.0 mm, about 14.5 mm, about 15.0 mm, or any range including and between any two of these values. It should be understood that the average inner gap may be different in each catalyst carrier. Average inner gaps in this range, and especially in the preferred range, advantageously and significantly improve flow uniformity of the gaseous stream through the catalyst carrier although the flow may not be uniform with a quality factor (quality factor = [max flow - min flow]/max flow) of about 2% to about 40%. "Max flow" refers to the maximum flow rate per unit cross-sectional area of the gaseous stream as it passes out from the perforated wall, and generally occurs at a different position on the perforated wall than the minimum flow rate of the gaseous stream as it passes out from the perforated wall. Max flow may be modeled using computational fluid dynamics. Where the average inner gap is less than about 2.0 mm, severe non-uniform flow (quality factor > 50%) leads to hot spots as well as far reduced per pass conversion for the catalyst carrier. These hot spots may lead to thermal runaway in a reactor.

[0093] The skirt, fins extending radially from the perforated inner wall, and network of heat conducting surfaces may each independently include steel, aluminum, copper, an alloy thereof, or a combination of any two or more thereof. For example, the skirt, fins extending radially from the perforated inner wall, and heat transfer structure may each independently include a combination of steel and aluminum, steel and copper, aluminum and copper, or steel, aluminum, and copper.

[0094] The distance between the fins extending radially from the perforated inner wall as measured at the perforated inner wall may be from about 0.01 mm to about 10 mm (in any embodiment herein when such fins are present). Thus, this distance between fins (as measured between the fins where they contact the perforated inner wall, running along the perforated inner wall) may be each channel may be about 0.01 mm, about 0.02 mm, about 0.04 mm, about 0.06 mm, about 0.08 mm, about 0.1 mm, about 0.2 mm, about 0.4 mm, about 0.6 mm, about 0.8 mm, about 1 mm, about 1.5 mm, about 2 mm, about 2.5 mm, about 3 mm, about 3.5 mm, about 4 mm, about 5 mm, about 6 mm, about 7 mm, about 8 mm, about 9 mm, about 10 mm, or any range including and/or in between any two of these values. The thickness of each fin may independently be from about 0.1 mm to about 5 mm; thus, each fin may independently have a thickness of about 0.1 mm, about 0.2 mm, about 0.4 mm, about 0.6 mm, about 0.8 mm, about 1 mm, about 1.5 mm, about 2 mm, about 2.5 mm, about 3 mm, about 3.5 mm, about 4 mm, about 5 mm, or any range including and/or in between any two of these values.

[0095] The network of heat conducting surfaces may include a random network of heat conducting surfaces, an ordered network of heat conducting surfaces, or a combination of both. The network of heat conducting surfaces themselves and/or in combination with the perforated outer wall and/or in combination with the outer surface of the skirt, and/or in combination with the inner surface of the skirt, and/or in combination with the inner tube wall define a plurality of channels. The channels may be of any cross-sectional shape, such as circular, oval, square, rhomboid, triangular, etc. The largest measurable span of the cross- sectional shape of a channel is taken to be the "channel diameter." The channels each independently have a channel diameter from about 0.01 mm to about 10 mm; thus, each channel may independently have a channel diameter of about 0.01 mm, about 0.02 mm, about 0.04 mm, about 0.06 mm, about 0.08 mm, about 0.1 mm, about 0.2 mm, about 0.4 mm, about 0.6 mm, about 0.8 mm, about 1 mm, about 1.5 mm, about 2 mm, about 2.5 mm, about 3 mm, about 3.5 mm, about 4 mm, about 5 mm, about 6 mm, about 7 mm, about 8 mm, about 9 mm, about 10 mm, or any range including and/or in between any two of these values.

[0096] As discussed previously, the SA/V is from about 400 m ² /m ³ to about 40,000 m ² /m ³ . Thus, SA/V may be about 400 m ² /m ³ , about 450 m ² /m ³ , about 500 m ² /m ³ , about 550 m ² /m ³ , about 600 m ² /m ³ , about 700 m ² /m ³ , about 800 m ² /m ³ , about 900 m ² /m ³ , about 1000 m ² /m ³ , about 1100 m ² /m ³ , about 1200 m ² /m ³ , about 1300 m ² /m ³ , about 1400 m ² /m ³ , about 1500 m ² /m ³ , about 1600 m ² /m ³ , about 1700 m ² /m ³ , about 1800 m ² /m ³ , about 1900 m ² /m ³ , about 2000 m ² /m ³ , about 2200 m ² /m ³ , about 2400 m ² /m ³ , about 2600 m ² /m ³ , about 2800 m ² /m ³ , about 3000 m ² /m ³ , about 3200 m ² /m ³ , about 3400 m ² /m ³ , about 3600 m ² /m ³ , about 3800 m ² /m ³ , about 4000 m ² /m ³ , about 5000 m ² /m ³ , about 6000 m ² /m ³ , about 7000 m ² /m ³ , about 8000 m ² /m ³ , about 9000 m ² /m ³ , about 10,000 m ² /m ³ , about 15,000 m ² /m ³ , about 20,000 m ² /m ³ , about 25,000 m ² /m ³ , about 30,000 m ² /m ³ , about 35,000 m ² /m ³ , about 40,000 m ² /m ³ , or any range including and/or in between any two of these values. For example, the SA/V may preferably be from about 400 m ² /m ³ to about 20,000 m ² /m ³ , more preferably from about 400 m ² /m ³ to about 4000 m ² /m ³ .

[0097] In any embodiment of the process, the contacting step may include

maintaining at least about 55% carbon monoxide conversion per pass in the one or more reactor tubes, where per pass conversion of CO is defined by the difference between the inlet and outlet moles of CO divided by the inlet number of moles of CO as converted in a tube that contains multiple cans in series. The particulate Fischer-Tropsch catalyst possesses a normalized CO conversion per pass loss of less than about 0.4% after 3 weeks of continued processing. Non-limiting examples of such particulate Fischer-Tropsch catalysts are described in International Patent Pub. WO 2012/107718, incorporated herein by reference in its entirety for any and all purposes. The carbon monoxide per pass conversion may be at least about 60%, preferably at least about 65%, even more preferably at least about 70%. The carbon monoxide per pass conversion may be less than about 90%, more preferably less than about 80%), to help ensure the water partial pressure is below levels that would accelerate solid synthesis catalyst deactivation, such as Fischer-Tropsch catalyst deactivation.

[0098] The process described herein may include a catalyst gas hourly space velocity of the gaseous stream in the tubular reactor from about 3,500 hr ^"1 to about 15,000 hr ^"1 . In any embodiment herein, the catalyst gas hourly space velocity of the gaseous stream in the tubular reactor may be greater than about 5,500 hr ^-1 . It may be that the catalyst gas hourly space velocity is about 5,000 hr ^"1 , about 6,000 hr ^"1 , about 7,000 hr ^"1 , about 8,000 hr ^"1 , about 9,000

hr ^"1 , about 10,000 hr ^"1 , about 11,000 hr ^"1 , about 12,000 hr ^"1 , about 13,000 hr ^"1 , about 14,000 hr ^"1 , about 15,000 hr ^"1 , or any range including and between any two of these values. [0099] The temperature of the gaseous stream at the reactor inlet may be about 150

°C to about 240 °C. For example, the temperature of the gaseous stream at the reactor inlet may preferably be about 160 °C to about 220 °C, even more preferably about 170 °C to about 190 °C.

[0100] A person of ordinary skill in the art will be familiar with solid synthesis catalysts suitable for generating preferred synthetic products. In general, solid synthesis catalysts may be porous. Solid synthesis catalysts may be particulate synthesis catalysts, monolithic synthesis catalysts, or a mixture of the two.

[0101] For example, where the synthetic product includes hydrocarbons, the solid synthesis catalyst may include a Fischer-Tropsch catalyst. Fischer-Tropsch catalysts may include cobalt (Co) or iron (Fe), and may further include a promoter such as Cu, Mn, Pd, Pt, Rh, Ru, Re, Ir, Au, Ag, Os, or a combination of any two or more thereof. For example, the Fischer-Tropsch catalyst may include Fe, Cu, and/or Mn. The Fischer-Tropsch catalyst may also include a support material. Suitable support materials include a refractory metal oxide, carbide, carbon, nitride or a mixture of any two or more thereof. The Fischer-Tropsch catalyst may further include a surface modified support material, wherein the surface of the support has been modified by being treated with silica, titania, zirconia, magnesia, chromia, alumina, or a mixture of any two or more thereof. In any of the above embodiments, the support material may include alumina, zirconia, silica, titania, or a mixture of two or more thereof. The support material may include a Ti0 ₂ -modified silica. In any of the above embodiments, the surface of the surface-modified support material may be amorphous.

[0102] Where the synthetic product includes methanol, the solid synthesis catalyst may include a copper-based catalyst such as Cu/ZnO/Al ₂ 0 ₃ . Where the synthetic product includes DME, the solid synthesis catalyst may include a blend of a methanol synthesis catalyst, such as Cu/ZnO/Al ₂ 0 ₃ , and a dehydration catalyst, such as γ-Α1 ₂ 0 ₃ .

[0103] Where the solid synthesis catalyst is a Fischer-Tropsch catalyst, the Fischer-

Tropsch catalyst may be a particulate Fischer-Tropsch catalyst or may be a monolithic Fischer-Tropsch catalyst. [0104] Thus, for the sake of clarity, it is understood that the particulate Fischer-

Tropsch catalyst may include cobalt, and may further include a promoter such as Cu, Mn, Pd, Pt, Rh, Ru, Re, Ir, Au, Ag, Os, or a combination of any two or more thereof. The particulate Fischer-Tropsch catalyst may also include one or more support materials. Suitable support materials include a refractory metal oxide, carbide, carbon, nitride or a mixture of any two or more thereof. The particulate Fischer-Tropsch catalyst may include a surface modified support material, wherein the surface of the support has been modified by being treated with silica, titania, zirconia, magnesia, chromia, alumina, or a mixture of any two or more thereof. In any of the above embodiments, the support material may include alumina, zirconia, silica, titania, or a mixture of two or more thereof. The support material may include a Ti0 ₂ modified silica. The surface of the surface-modified support material may be amorphous.

[0105] The particulate Fischer-Tropsch catalyst may be coated on a support structure such as a carbon nanotubes, and wherein the carbon nanotubes are disposed on a support material. The support can be made of a variety of materials such as ceramic, but where rapid heat transport is preferred; the support preferably is a thermally conductive material such as a metal. The support may be stainless steel, an alloy such as monel, cordierite, silica, alumina, rutile, mullite, zirconia, silicon carbide, aluminosilicate, stabilized zironia, steel and alumina- zirconia blend. For use in the present technology, U.S. Pat. No. 6,713,519 provides examples of suitable carbon nanotube-on-support materials over which a particulate Fischer-Tropsch catalyst may be coated.

[0106] The particulate Fischer-Tropsch catalyst may have a Co surface area per unit of packed bed volume of about 15 m ² /mL to about 65 m ² /mL, or any integer in between or range including any two such integers. Thus, the Co surface area per unit of packed bed volume may be about 15 m ² /mL, about 20 m ² /mL, about 25 m ² /mL, about 30 m ² /mL, about

35 m 2 /mL, about 40 m 2 /mL, about 45 m 2 /mL, about 50 m 2 /mL, about 55 m 2 /mL, about 60 m ² /mL, about 65 m ² /mL, or any range including and/or in between any two of these values. For example, the Co surface area per unit of packed bed volume may be about 25 m ² /mL to about 60 m ² /mL, preferably from about 35 m ² /mL to about 55 m ² /mL. Such values may be determined by several methods, including hydrogen chemisorption as described in Storsaeter, S. et al. "Characterization of alumina-, silica-, and titania-supported cobalt Fischer-Tropsch catalysts" Journal of Catalysis 2005, 236, 139-152, incorporated herein by reference in its entirety for any and all purposes.

[0107] The particulate Fischer- Tropsch catalyst may possess a weight average diameter from about 100 micrometers (μιη) to about 1 millimeter (mm), including any subrange therein. For example, the particulate Fischer- Tropsch catalyst may possess a weight average diameter from about 200 μιη to about 750 μιη; the particulate Fischer-Tropsch catalyst may possess a weight average diameter from about 250 μιη to about 450 μιη. The particulate catalyst may be spherical or non- spherical. For a non-spherical particle with a spheroid shape, the weight average diameter is based on the average particle dimension of the spheroid-shaped particle's diameters. For an elliptical or rod shaped particle with two major axis, the weight average diameter is based on the average of the smaller axis dimension. As an example, for a rod shaped particle, the diameter rather than the length of the rod would be considered. For a Rachig ring shape particle, the outer diameter minus the inner ring diameter is considered as the particle diameter. For other odd shaped catalyst, including a tri-lobe and others, the weight average particle diameter is based on the circumscribed circle that would be required to have the particle fit inside.

[0108] The particulate Fischer-Tropsch catalyst may have a Co loading from about 20 wt% to about 56 wt% based on the total weight of the particulate Fischer-Tropsch catalyst, including any subrange therein. For example, the particulate Fischer-Tropsch catalyst may have a Co loading of about 20 wt%, about 22 wt%, about 24 wt%, about 26 wt%, about 28 wt%, about 30 wt%, about 32 wt%, about 34 wt%, about 36 wt%, about 38 wt%, about 40 wt%, about 42 wt%, about 44 wt%, about 46 wt%, about 48 w%, about 50 wt%, about 52 wt%, about 54 wt%, about 56 wt%, or any range including and between any two of these values.

[0109] As described above, the cooling medium is in contact with the one or more reactor tubes. The cooling medium temperature may be from about 160 °C to about 240 °C, more preferably from about 180 °C to about 210 °C, and even more preferably from about 190 °C to about 210 °C. [0110] In any embodiment described herein, the process may include a difference between the reactor inlet temperature and the cooling medium temperature that is less than about 80 °C. The difference may be less than about 40 °C, preferably less than about 30 °C, more preferably less than about 20 °C, and even more preferably less than about 10 °C.

[0111] In any aspect and embodiment herein, the process may include maintaining a temperature difference between the cooling medium and the product exiting the tube outlet between about 20 °C and about 80 °C. Thus, this temperature difference may be about 20 °C, about 25 °C, about 30 °C, about 35 °C, about 40 °C, about 45 °C, about 50 °C, about 55 °C, about 60 °C, about 65 °C, about 70 °C, about 75 °C, about 80 °C, or any range including and between any two of these values. For example, the process may preferably include maintaining a temperature difference between the cooling medium and the product exiting the tube outlet between about 30 °C and about 65 °C.

[0112] The process may include maintaining a temperature difference between the cooling medium and the solid Fischer- Tropsch synthesis catalyst between about 20 °C and about 80 °C. The temperature difference between the cooling medium and the solid synthesis catalyst may be about 20 °C, about 30 °C, about 40 °C, about 50 °C, about 60 °C, about 70 °C, about 80 °C, or any range including and between any two of these values. For example, the temperature difference between the cooling medium and a particulate Fischer-Tropsch catalyst may preferably be about 30 °C to about 60 °C. By way of a non-limiting illustration, the reactor tube may include one or more thermocouples within the reactor tube between the tube inlet and the tube outlet; and the process may further include, in response to a temperature difference between the cooling medium and a temperature detected by the thermocouple of greater than about 80 °C, decreasing the cooling medium temperature so that the temperature of the solid synthesis catalyst is subsequently reduced such that the temperature difference between the cooling medium and the solid synthesis catalyst is within the desired range. Similarly, the process may further include, in response to a temperature difference between the cooling medium and a temperature detected by the thermocouple of less than about 20 °C, increasing the cooling medium temperature. [01 13] In any aspect or embodiment herein, the process may include a temperature rise of about 20 °C to about 80 °C across the one or more catalyst carriers. Such a temperature rise may be determined by measuring the difference between the temperature of the gaseous stream prior to entering one or more catalyst carriers and the temperature of the gaseous stream exiting the one or more catalyst carriers. The temperature rise may be about 20 °C, about 25 °C, about 30 °C, about 35 °C, about 40 °C, about 45 °C, about 50 °C, about 55 °C, about 60 °C, about 65 °C, about 70 °C, about 75 °C, about 80 °C, or any range including and between any two of these values. For example, the process may preferably include maintaining a temperature rise between about 30 °C and about 65 °C.

[01 14] In any aspect or embodiment involving a Fischer- Tropsch catalyst, the processes may include a periodic catalyst rejuvenation step. In the periodic catalyst regeneration step, the contacting step is discontinued for the duration of the catalyst rejuvenation step. "Periodic" as used herein will be understood to mean occurring after the activity of the Fischer-Tropsch catalyst has decreased and/or there is a particular increase in temperature of the contacting step to maintain about a constant percent conversion of CO by the Fischer-Tropsch catalyst. The particular increase in temperature may be about 5 °C as compared to a temperature previously employed for the same percent conversion of CO. The particular increase in temperature may be about 5 °C, or about 10 °C, or about 15 °C, or about 20 °C, or an increase in temperature greater than any one of these values. The rejuvenation step involves flowing a rejuvenation gas including H ₂ over the Fischer-Tropsch catalyst, and may involve a temperature of about 200 °C to about 400 °C, or any range including and in between any two integers between these two values, preferably about 350 °C. Such a rejuvenation step strips off a portion of poisons that may become associated with the Fischer-Tropsch catalyst (e.g., l¾) during the contacting step.

[01 15] The process may include a periodic wax removal step to remove accumulated hydrocarbons from the surfaces of catalyst and the annular spaces in the reactor in order to maintain pressure drop at a manageable level, and may also be performed prior to a rejuvenation step or a regeneration step as described herein, or prior to shut down the reactor for extended periods, or prior to removal of catalyst carriers. The periodic wax removal step may include flowing a dewaxing fluid, such as hydrogen or nitrogen through the reactor tubes where the reactor tubes are at an initial temperature of about 20 °C to about 170 °C (or any range including and between any two integers thereof). Flowing the dewaxing fluid may include flowing the dewaxing fluid at a gas hourly space velocity of about 1,000 h ^"1 to about 20,000 h ^"1 . The GHSV may preferably be up to about 10,000 h ^" , more preferably about 5,000 h ^"1 , and even more preferably about 1,000 h ^"1 . The wax removal may be performed at pressures of about 1 barg to about 25 barg (or any range including and between any two integers thereof), preferably from about 5 barg to about 15 barg. Such pressures facilitate recycling of the dewaxing fluid via use of a compressor. The reactor tubes are then heating to a hold temperature of about 250 °C to about 450 °C at a rate of about 1 °C per hour to about 60 °C per hour. The temperature of the reactor tube is then maintained at the hold temperature for about 2 hours to about 72 hours, whereupon the reactor tube is subsequently brought to a final temperature of about 20 °C to about 170 °C at a rate of about 1 °C per hour to about 60 °C per hour. The final temperature used will depend upon whether the reactor will be shut down, whether conversion of synthesis gas will resume, or whether a

rejuvenation step or regeneration step will be performed. Upon reaching the final temperature, the flow of dewaxing fluid may be discontinued. Optionally, at this stage, a heated inert gas (such as nitrogen or argon) may be flowed through the reactor tube to further remove hydrocarbons, where the heated inert gas may be flowed at a gas hourly space velocity of about 1,000 h ^"1 to about 20,000 h ^"1 . The GHSV may preferably be up to about 10,000 h ^"1 , more preferably about 5,000 h ^"1 , and even more preferably about 1,000 h ^"1 . The temperature of the inert gas may be about 250 °C to about 450 °C.

[0116] The processes involving Fischer- Tropsch catalysts may include a periodic catalyst regeneration step. A person of skill in the art understands it is sometimes desirable to perform a rejuvenation step rather than a regeneration step, or vice versa, and understands when to perform one versus the other. In the periodic catalyst regeneration step, the contacting step is discontinued for the duration of the catalyst regeneration step. Such synthesis catalyst regeneration is well known in the art and as recommended by catalyst suppliers for the particular Fischer-Tropsch catalyst to be regenerated. In any of the above embodiments, the catalyst regeneration step may involve:

(1) a dewaxing step involving flowing a dewaxing gas including H ₂ over the Fischer- Tropsch catalyst, (2) an oxidation step involving flowing an oxidation gas over the Fischer-Tropsch catalyst, and

(3) a reduction step involving exposing the Fischer-Tropsch catalyst to a reducing gas that includes H ₂ .

The temperature dewaxing step, the oxidation step, and the reduction step may each independently be from about 200 °C to about 400 °C, or any range including and in between any two integers between these two values, and preferably is about 350 °C. The oxidation gas may include one or more of air and N ₂ -diluted air. The dewaxing step typically involves breaking down product associated with the Fischer-Tropsch catalyst; the oxidation step typically involves combusting residual hydrocarbons and/or oxygenated hydrocarbons and oxidizes the Fischer-Tropsch catalyst; and the reduction step typically involves reducing the oxidized Fischer-Tropsch catalyst back to its active form.

The Synthetic Product Provided by the Present Technology and Optional Further Processing

[01 17] The term "synthetic product" as used herein in regard to the presently disclosed technology includes hydrocarbons, oxygenated hydrocarbons, or combinations thereof. Oxygenated hydrocarbons include, but are not limited to, alkanes, alkenes, and alkynes that are each substituted with one or more of an epoxy, hydroxyl, or a carbonyl group. Exemplary carbonyl-containing groups include, but are not limited to an aldehyde, a ketone, a carboxylic acid, a carboxylic acid anhydride, or an ester. Thus, the synthetic product of the present technology includes one or more compounds selected from Ci to Cioo hydrocarbons, Ci to Cioo oxygenated hydrocarbons, or a combination thereof, or any range including and in between any carbon number between Ci and Cioo; for example, the synthetic product may include C10-C14 hydrocarbons. In any of the embodiments described herein, the synthetic product may predominantly include one or more compounds selected from Ci to C ₅ o hydrocarbons, Ci to C50 oxygenated hydrocarbons, or combinations thereof.

"Predominantly" as used herein means at least about 51 weight percent ("wt%") of the synthetic product. The product may include C ₅ + hydrocarbons (i.e., hydrocarbons with 5 or more carbon atoms). The synthetic product of the present technology may include one or more compounds selected from Ci to C50 hydrocarbons, Ci to C50 oxygenated hydrocarbons, or combinations thereof in an amount of about 51 wt% to about 100 wt%, or any range including and in between any integer between these two values. Thus, the synthetic product may include 40 wt% of C ₁₄ -C ₁₈ hydrocarbons; the synthetic product may include 10 wt% of C1-C4 monohydroxyalkanes (i.e., monohydric alcohols).

[0118] Hydrocarbons of the synthetic product of any aspect and embodiment described herein may further be reacted to provide a desired product.

[0119] For example, the hydrocarbons may be directed to a hydrocracking reaction to reduce the carbon number of the product to increase the content of gasoline, diesel fuel, jet fuel, or other valuable hydrocarbon-containing blends. Hydrocracking catalysts suitable for such reactions may include zeolite catalysts. Zeolite catalysts include, but are not limited to, beta zeolite, omega zeolite, L- zeolite, ZSM-5 zeolites and Y-type zeolites. The

hydrocracking catalyst may also include one or more pillared clays, MCM-41, MCM-48, HMS, or a combination of any two or more thereof. The hydrocracking catalyst may include Pt, Pd, Ni, Co, Mo, W, or a combination of any two or more thereof. The hydrocracking catalyst may further include a refractory inorganic oxide such as alumina, magnesia, silica, titania, zirconia, silica-alumina, or combinations of any two or more thereof. The

hydrocracking catalyst may further include a hydrogenation component. Examples of suitable hydrogenation components include, but are not limited to, metals of Group IVB and Group VIII of the Periodic Table and compounds of such metals. For example, molybdenum, tungsten, chromium, iron, cobalt, nickel, platinum, palladium, iridium, osmium, rhodium, ruthenium, or combinations of any two or more thereof may be used as the hydrogenation component. Exemplary catalysts are described in U.S. Patent 6,312,586, which is

incorporated herein by reference in its entirety for any and all purposes.

[0120] The hydrocarbons may be directed to a hydrotreating, where the hydrotreating involves a hydrotreating catalyst. The hydrotreating catalyst may include Ni, Mo, Co, W, or combinations of any two or more thereof. The hydrotreating catalyst may be a supported catalyst, such as a hydrotreating catalyst supported on alumina. In some embodiments, the catalyst may include Mo-W/Al ₂ 0 ₃ .

[0121] It may be that the hydrocarbons are directed to a hydrocarbon oxidation involving an oxidation catalyst. The oxidation catalyst may include a metal, metal oxide, or mixed metal oxide of Mo, W, V, Nb, Sb, Sn, Pt, Pd, Cs, Zr, Cr, Mg, Mn, Ni, Co, Ce, or a combination of any two or more thereof. These catalysts may further include one or more alkali metals or alkaline earth metals or other transition metals, rare earth metals or lanthanides. Elements such as P and Bi may be present. The catalyst may be supported and, if so, useful support materials include metal oxides (e.g. alumina, titania, zirconia), silica, mesoporous materials, zeolites, refractory materials, or combinations of two or more thereof.

[0122] In any of the aspects and embodiments described herein, it may be that the hydrocarbons are directed to a hydrocracking, a hydrotreating, or combination thereof.

EXAMPLES

[0123] The examples herein are provided to illustrate advantages of the present technology and to further assist a person of ordinary skill in the art with preparing or using the processes of the present technology. The examples herein are also presented in order to more fully illustrate the preferred aspects of the present technology. The examples should in no way be construed as limiting the scope of the present technology, as defined by the appended claims. The examples can include or incorporate any of the variations,

embodiments, or aspects of the present technology described above. The variations, embodiments, or aspects described above may also further each include or incorporate the variations of any or all other variations, embodiments, or aspects of the present technology.

[0124] The reactor model used in the following examples utilizes the following steady state balances: Species balance (all gas phase) for 10 species to track reactant usage in each a catalyst carrier (hereafter also termed "can"), energy balance for each can reactor where the catalyst zone in each can reactor operates adiabatically, energy balance for the cooling annular space where the phase change on the outside of the can reactor is used to remove heat, a momentum balance in the reactor utilizing an Ergun equation with gas-liquid flow multiplier of 2 to account for higher pressure drop in the multiphase reactor as compared to a dry bed and finally a momentum balance in the cooling annular space. The annular space of the catalyst carrier is referred to as "heat exchanger gap" or "gap." [0125] The geometry of an individual can is shown in FIG. 1 in top view. The internal feed plenum is at the center of the can reactor. The can reactor has an internal radius, RO and an external radius, Rl, both of which fix the catalyst used in the reactor while allowing reactant and products to enter and leave the can. The reactants and products leave the can in the radial direction and exit over a wall to the annular space for heat removal. The annular space (i.e., the space between the internal radius, R2 and the outer radius, R3) is shown in FIG. 1. Note the containment wall for the outer radius of the can adds thickness to the device's dimensions.

[0126] The layout of internal cross-section for a can i as used in the reactor model is illustrated in FIG. 2A. The gaseous stream containing reactants enters at the top of the can with an inlet temperature (Ti _n [i]), pressure (Pin[i]), molar flow rate (n[i]), and the ten species mole fractions (y[i,j], j = 1..10). The stream undergoes reactions in the can and exits with changes to the outlet temperature (T _out [i]), pressure (P _ou t[i]), molar flow rate (n[i+l]), and mole fractions (y[i+l j], j = 1..10). These stream properties enter the annular space and undergoes a change in pressure (Pin[i+1]) and temperature (Ti _n [i+1]). The reactor model assumes pressure drop for the flow as it exits the annular space and before it enters the internal diameter of the next can in series is negligible. The wall temperature (T _wall [i]) is set for most cases at a constant value via the use of boiling heat transfer. It is also noted that a near constant wall temperature can be achieved using non-phase change, or convective heat transfer when the liquid coolant medium (also referred to as "heat exchanger fluid") has a substantially larger flowrate than required for phase change heat transfer. One of ordinary skill in the art can calculate the required liquid heat exchanger fluid flowrate to result in a temperature rise of less than 1 °C.

[0127] The mass flow rate is set for the first can, along with individual species mole fractions of the inlet gas. The model tracks ten species: reactant and inlet diluents (H ₂ , CO, N ₂ ), light gas products and diluents (C0 ₂ , CH ₄ , C ₂ H ₆ , C ₃ H ₈ , C ₄ Hi ₀ ) and a large liquid hydrocarbon "wax" (Ci ₄ H ₃₀ ) species representative of the entire liquid product. For ease in the reactor model, all components are in the gas phase even if a species would be in the liquid phase at the operational conditions. There are six forward reactions in this model which are catalyzed in the can reactor, which are presented below along with their respective standard state heats of reaction (25°C) per gram-mole CO consumed:

1. CH ₄ synthesis: CO + 3 H ₂ CH ₄ + H ₂ 0 (ΔΗ _κη = -210 kJ/mol CO)

2. C0 ₂ synthesis: CO + H ₂ 0 C0 ₂ + H ₂ (ΔΗ^ = -41 kJ/mol CO)

3. C ₂ H ₆ synthesis: 2 CO + 5 H ₂ C ₂ H ₆ + 2 H ₂ 0 (ΔΗ _κη = -173 kJ/mol CO)

4. C ₃ H ₈ synthesis: 3 CO + 7 H ₂ ^ C ₃ H ₈ + 3 H ₂ 0 (ΔΗ _κη = -167 kJ/mol CO)

5. C ₄ Hio synthesis: 4 CO + 9 H ₂ C ₄ Hi ₀ + 4 H ₂ 0 (ΔΗ^ = -161 kJ/mol CO)

6. FTS synthesis: 14 CO + 29 H ₂ Ci ₄ H ₃₀ + 14 H ₂ 0 (ΔΗ _κη = -166 kJ/mol CO)

The conditions for reaction zone inside the can reactor section are based upon the average of Tin[i+1] and T _out [i] for temperature, the average of Pi _n [i] and P _out [i] for pressure, the average mole fractions of y[i] and y[i+l].

[0128] The rate of the chemical reactions 1 through 6 were provided by Almeida et al. "Kinetic analysis and microstructured reactors modelling for the Fischer-Tropsch synthesis over a Co-Re/ A1 ₂ 0 ₃ catalyst", Catalysis Today, 215 (2013) 103-111 (incorporated herein by reference for any and all purposes) for catalyst layer thicknesses judged to be in a kinetically controlled regime {i.e., based upon the reported ten micrometer distance in a sixty micrometer diameter particle). Reactions 1 and 3 through 6 take a Langmuir-Hinshelwood form with CO adsorption terms based upon partial pressures of hydrogen and carbon monoxide. The carbon dioxide synthesis follows a power law form with only steam partial pressure dependence to the 0.35 power. To manage the overall process pressure drop, the catalyst particle diameter used for deployment in the current can-based reactor modeling was one millimeter, almost two orders of magnitude larger than the effective dimension. The effectiveness factor for all reaction rates provided by the Almeida reference was chosen to be 0.003 with an assumed catalyst density of 1000 kg/m ³ .

[0129] The energy balance for the can assumes adiabatic operation, with all energy produced from the exothermic reactions that occur in the can going directly into the exiting gas stream. A CFD study was done and it shows negligible heat conduction between the catalyst bed to the heat exchanger fluid as each catalyst bed operates near adiabatically where the heat of reaction energy is transferred to the exiting gas before removal by convective heat transfer at the annular space of the can. The exiting gas flows down the gap made between R2 and R3 in FIG. 1 (and also illustrated in FIG. 2A). The wall temperature is set at a fixed value by highly efficient boiling heat transfer around the tubes that contain multiple can reactors in series. The stream leaving the can enters into the topmost section of the heat exchanger at T _out [i] and when heat is removed through the heat transfer wall it exits at T _in [i+1], assuming no heat exchange with the can above or below, and the stream is directed to the inner diameter of the next can. The heat exchange area for can i is the can height multiplied by π and twice the heat exchanger section's outer radius R3. The heat transfer conditions are based upon the stream exit composition y[i+l] and the average temperature of T _out [i] and T _in [i+1] for heat capacity, thermal conductivity and viscosity for calculating the Reynolds and Prandtl numbers. The Nusselt number used for this heat exchanger was for parallel plates with constant heat flux and turbulent flow from Kays and Leung, Int. J. Heat Mass Transfer, 6, 537-557, 1963 (incorporated herein by reference in its entirety), applicable for concentric cylinders as the ratio of R2 to R3 approaches unity.

[0130] The momentum balance for each can focuses on the two largest contributions to the pressure drop in the can: flow through the packed bed and flow through the annular space. The pressure drop is based upon the full Ergun equation, with viscous (Darcy) and inertial (Forschiemer) terms:

r) where ΔΡ [Pa] is the pressure drop, AL [m] is the differential path length in the direction of fluid flow, μ is the fluid's dynamic viscosity [kg/m/s], φ ₅ is the catalyst particle's sphericity,

Dp [m] is the average particle diameter, ε is the bed void fraction, G [kg/m ² /s] is the mass flux rate calculated over the channel's total cross-sectional area and p [kg/m ] is the fluid density. The can's catalyst packing void fraction is assumed to be 0.35 and the particle sphericity is 0.9. The path length of the flow is the difference between Rl and R0 (FIG. 1) and the average cross-sectional area of π times the can height time the average of Rl and R0. Because a "dry" gas flow for the reactions is used (i.e., the total gaseous flowrate is based on all reactants and products remaining in the gas phase), the higher gas-liquid pressure drop in the can (that would be expected due to the small volume fraction (0.1-1%) of liquids involved in actual reactions) is taken into account with a net multiplicative factor to the Ergun equation of two as based on typical Fischer-Tropsch experimental results for high conversion data. The pressure drop in the can heat exchanger is based upon the parallel plate losses for turbulent flow by Dean in J. Fluids Eng., 100, pp 215-223, 1978 (incorporated herein by reference in its entirety) using the same average temperatures for heat exchanger zone viscosity and density as the heat exchange calculation and the average pressure of P _ou t[i] and Pin[i+1] for density of the gas. The momentum balance ignores two-phase flow in the heat exchanger, so it is important that the heat exchanger wall temperature is above the reactor exit water dew point to avoid water condensation. The low pressure drop from the stream distributing to and from the can, the turns into and out of the annular space and the collection of flow into the inner annulus of the downstream can are ignored.

[0131] Table 1 provides initial parameters explored, while Table 2 provides the results of this initial study. "Reaction Path Length" refers to the difference between Rl and R0 of the carrier as shown in FIG. 1; "Inner Gap Diameter" refers to the difference between R2 and Rl as shown in FIG. 1, and "Annular Gap Diameter" refers to the difference between R3 an R2 as shown in FIG. 1. "Catalyst Height" refers to the vertical height of the catalyst in the carrier (see FIG. 2A). "BPD" refers to U.S. barrels per day. "Wall BC" refers to the wall temperature boundary condition. Particular advantages provided by the present technology are discussed in Examples 1 & 2 below. Table 1.

[0132] Example 1: Surprising Advantages Provided by Increasing Thermal

Gradients

[0133] FIG. 3 illustrates the effect generating a higher temperature gradient (the difference between the maximum gaseous stream temperature exiting at least one of the parallel catalyst containers and the heat exchanger wall temperature tube wall) exhibits on the CO conversion at an inlet pressure of 1,000 psig. FIG. 4 illustrates the effect of the higher temperature gradient on the methane (CH ₄ ) selectivity. While, as expected, the methane selectivity increases with larger temperature gradients, the change in methane selectivity with increasing temperature gradient is surprisingly very low - between a 20 °C gradient and a 50 °C gradient the methane selectivity only increases from about 7.8% to about 13.5%> while the CO conversion increases from 58%> to 88%>. Thus, the high conversion of CO is not at the expense of the yield for the desired higher C ₅ + carbon number product, counter to the art- appreciated expectation.

[0134] This surprising result was investigated for a lower reactor inlet pressure of 750 psig, where one would expect lower CO conversion. FIG. 5 illustrates the effect of the temperature gradient on the CO conversion at this pressure, and FIG. 6 plots the methane selectivity as the temperature gradient is increased. CO conversion surprisingly increases significantly with higher temperature gradients, from about 43% CO conversion at a 20 °C gradient to about 75% CO conversion at a 50 °C gradient, with surprisingly little loss to methane formation (about 5.8%> at the 20 °C gradient to about 11.5% at the 50 °C gradient).

[0135] The reactor inlet pressure was further decreased to 350 psig to see if this surprising and advantageous effect would be observed despite the low inlet pressure. FIG. 7 plots the temperature gradient versus CO conversion at 350 psig, and FIG 8 plots methane selectivity as the temperature gradient is increased at a reactor inlet pressure of 350 psig. Even at 350 psig, increasing the temperature gradient from 20 °C to 50 °C significantly increases CO conversion (38%> to 67%) with only small losses to methane formation (5% to 10%). [0136] While higher pressure is generally understood to reduce methane selectivity and increase productivity, it does so at significantly higher costs (due to higher costs to provide such pressure as well as and fabrication costs of the reactor (e.g., wall thickness, flange sizes, etc.). Surprisingly, generating higher temperature gradients counter to the art- appreciated understanding in the presently disclosed process enhances production of desired product at lower inlet pressures than 750 psig, thus providing significant savings. Refer to Table 1 and Table 2, which illustrates that the minimum number of tubes is achieved at 750 psig when using 200 cans per tube. The minimum number of cans per tube, required to avoid a temperature runaway, increases as the pressure is increased. A person of ordinary skill in the art, based on the present disclosure, will be able to perform an economic optimization to balance the decreasing number of tubes with the increasing pressure for a given reactor capacity and with the increasing cost per tube.

[0137] Example 2: Impact of a non-uniform catalyst profile on performance

[0138] Further explored in this reactor model was whether reactor performance can be further improved with the use of a non-uniform catalyst profile. While the previous example utilized the same catalyst activity in all cans, three models investigated non-uniform catalyst profile in the reactor. In the first of these models (Model A), the first 80 cans were modeled to exhibit 0.003 times the Almeida et al. catalyst activity ("0.003x catalyst activity") and cans 81-160 were modeled to exhibit 0.006x catalyst activity and the CO conversion versus the thermal gradient was plotted (FIG. 9) along with the methane selectivity versus the thermal gradient (FIG. 10). In the second model (Model B), the first 80 cans were modeled to exhibit 0.003x catalyst activity and cans 81-160 were modeled to exhibit 0.0045x catalyst activity. The CO conversion versus the thermal gradient for Model B is shown in FIG. 11, and the methane selectivity versus the thermal gradient illustrated in FIG. 12. The third model (Model C) was modeled so the first 80 cans exhibit 0.003x catalyst activity and cans 81-160 exhibit 0.00375x catalyst activity. The CO conversion versus the thermal gradient for Model B is shown in FIG. 13, and the methane selectivity versus the thermal gradient illustrated in FIG. 14. As shown by the results, the non-uniform catalyst profile allows for higher CO conversion with lower methane selectivity at lower catalyst thermal gradients. The reactor architecture of placing externally loaded cans/carriers in series facilitates modifying catalyst activity along the length of the reactor. In one limiting case, the catalyst activity could be modified in each individual can placed in series. As another example, one might consider four different catalyst activities along the full tube length, preferably with increasing catalyst activity towards the reactor outlet end.

[0139] Surprising Advantages of Present Technology over Gamlin's Technology

[0140] Example 3

[0141] Gamlin teaches a reactor with two meters of total flow path length as achieved by a number of cans stacked in series and a total can height on the order of twenty meters tall. Gamlin provides an example that purports to show that when processing with a gas hourly space velocity around four thousand, the reactor could operate with a Fischer-Tropsch liquid production rate of two-thousand and three-hundred barrels per day. Gamlin taught that the heat exchanger gap needed to maintain temperatures in the reactor were as large as ten millimeters and down to three millimeters.

[0142] To evaluate whether Gamlin's reactor would work as taught, the previously described reactor model was set with eighty cans in series per tube. The inner diameter R0, outer diameter Rl and height of each can was 0.635 cm, 3.175 cm and 25.4 cm, respectively (see FIG. 1). Each can's outer radius R3 of the heat exchanger was set at 3.759 cm and the inner radius R2 was set to 3.454 cm to make a heat exchanger gap just over three millimeters. These dimensions make a total radial path length of 2 meters and a can catalyst height of roughly 20 meters as taught by Gamlin.

[0143] To reach a gas hourly space velocity on the order of four thousand, a mass flow rate was set at 188 kg/hour with inlet mole fractions of hydrogen and carbon monoxide of 0.4345 and 0.2455, respectively, and the balance nitrogen. When the inlet pressure is 35.49 bar gauge, the inlet gas and heat exchanger wall surface are maintained at 172 °C.

[0144] The model results for CO conversion at the exit of the last can are plotted in

FIG. 15 with methane and C ₅₊ selectivity, with values of 68.0%, 7.4% and 87.1%>

respectively. The capacity at this condition is 4.36 barrels per day of liquid product per reactor tube of eighty cans in series and 2293 of these eighty-can reactor tubes needed to achieve ten thousand barrels per day total capacity of C ₅ + products. FIG. 16 plots the inlet and outlet temperatures for each can, with a maximum catalyst temperature of 235.5 °C leaving the hottest can. A net increase of 63.5 °C is observed for the difference between the reactor can exit temperature to heat exchanger wall temperature. This is much larger than the ten to forty degrees Celsius range that Gamlin teaches for the reactor with a three to ten millimeter annular space. The overall heat production per reactor tube was 84 kilowatts. The pressure drop for the reactor tube was 7.6 psig. The water mole fraction at the exit had a dew point temperature of 171.2 °C, very near to the heat transfer wall temperature of 172 °C and in potential danger for condensing water in the heat exchanger channel before flowing to the next catalyst can. This is especially concerning because water is known to deactivate cobalt- based Fischer-Tropsch catalysts.

[0145] Thus, the conditions required to achieve the Gamlin's reported productivity are counter to Gamlin's required conditions. Namely, Gamlin teaches a modest temperature differential in each can to achieve Gamlin's reported productivity. However, as shown by the present reactor model, providing a very high temperature gradient (about 65 °C from the maximum catalyst temperature to the heat transfer fluid temperature) would be required to achieve the productivity suggested by Gamlin.

[0146] Example 4

[0147] The reactor in Example 3 utilized process conditions to achieve a barrel per day capacity that matched the capacity described by Gamlin, however the conditions required were in contradiction with Gamlin's disclosed process conditions. To further explore this issue and compare it to the process of the present technology, exemplary reactors containing 80 catalyst carriers with average annular spaces according the present technology (1 and 2 millimeters) as well as 80 catalyst carrier reactors with Gamlin's taught gaps (3, 4, and 5 millimeters) were each modeled in an attempt to push each reactor to its highest barrel per day capacity while avoiding reactor runaway. The capacity of the reactors was increased from a gas hourly space velocity of over four thousand to over six-thousand five-hundred.

[0148] FIG. 17 plots the shared inlet and wall heat exchanger temperature

combinations that were used to maximize reactor capacity by using the catalyst kinetics from Example 3. Several gas hourly space velocities and heat exchanger gaps were evaluated while maintaining the maximum reactor can exit temperature less than 240 °C. Interestingly, with larger gaps, lower heat exchanger wall temperatures are required to maintain the heat removal. For example, for a gas hourly space velocity 4080 v/v/hour, a 208° C wall temperature is required for 1.0 mm gap while a precipitously low 151 °C is required for the 5.0 mm gap. A wall temperature near or below the dew point temperature may cause a problem of Fischer Tropsch product condensation, wax accumulation, and/or channel plugging which may reduce the reactor productivity or cease operation.

[0149] The effect of can reactor heat exchanger gap size upon the reactor capacity is plotted in FIG. 18, where the number of 80 can reactor tubes (described in Example 3) required to achieve a capacity of 10,000 barrels per day capacity are plotted versus gas hourly space velocity. While increased capacity can be obtained at higher gas hourly space velocities for all heat exchanger gap sizes, Gamlin's heat exchanger gaps limit the capacity by hundreds and up to a thousand reactor tubes, especially at lower gas hourly space velocities for a 10,000 barrel per day plant. Utilizing a process according to the present technology (e.g., with average annular space from about 1.3 mm to about 2.6 mm), the capacity improves. Such a reduction in reactor tubes translates into a significant reduction in the capital cost of facilities operating by the present technology, due to a reduction in the volume, weight, and materials required to fabricate the reactor as well as a reduction in the heat exchanger network used in concert with reactor operation.

[0150] FIG. 19 illustrates an important advantage provided by average annular spaces of the present technology: improvement of the per pass carbon monoxide conversion by increasing the temperatures over the entire reactor. FIG. 19 illustrates the difference in reactor carbon monoxide conversion at various gas hourly space velocities were four to eight percentage points higher for the two millimeter gap as compared to the three millimeter gap while maintaining carbon monoxide conversions equal to or greater than 60%. Without being bound by theory, it is believed this is at least partially due to more heat removal at each can in comparison with Gamiln's annular spaces. Annular gaps larger than three millimeters could not consistently maintain a per pass carbon monoxide conversion greater than 55% with increasing gas hourly space velocity and such reactors would require greater plant capital and operational costs due to the need to use larger recycle streams to achieve acceptable overall carbon monoxide conversion.

[0151] FIG. 20 plots the can maximum exit temperature to heat exchanger wall temperature difference for the cases described in previous figures. Notably, temperature differences for Gamlin's gap size (i.e., three millimeters and larger) are 47 °C and larger for every catalyst gas hourly space velocity, significantly greater than the 10 °C - 40 °C range taught by Gamlin. FIG. 21 plots the model pressure drop predictions for the eighty-can reactors for the different average annular spaces at differing catalyst gas hourly space velocities. FIG. 22 illustrates the per pass C ₅ + yield for the different average annular spaces at differing catalyst gas hourly space velocities.

[0152] Example 5

[0153] The reactors in Examples 3 and 4 evaluated the same number of cans per reactor tube as the Example from Gamlin, i.e., 80 cans in series. A further study was performed to examine the high end of Gamlin's range (200 cans in series). For ease of comparison, this study utilized the same overall catalyst gas hourly space velocities and catalyst volumes employed in in the previous examples, but the number of cans is increased from 80 to 200 per reactor tube. This necessitated smaller can sizes for the same reactor tube length. Potentially, this allows for more opportunities for heat removal with smaller cans.

[0154] The only change in the can designs, as compared to Examples 3 and 4, is reduction of the overall height each can from 25.4 cm to 10.2 cm. The surface area for heat transfer in each can scales proportionally to this height. The superficial velocity in each can increase two and one half times that of the cans in Examples 3 and 4. This example again sets the reasonable maximums temperature to be 240 °C, the maximum value tested by Almeida et al in their kinetics and the tendency for temperatures higher than 240 °C to cause reactor run away. It is noted that certain Fischer Tropsch catalysts may exhibit operability with a temperature up to 280 °C. One goal of the model results in this example was to determine the highest barrel per day capacity for the 200 can reactor while keeping the reactor thermally stable as the exothermic reactions take place. [0155] FIG. 23 plots the wall and inlet stream temperatures for the 200 can model results needed to maximize the barrel per day capacity. Somewhat larger temperatures with increasing gas hourly space velocity are observed as compared to the 80-can reactor (see Example 4, FIG. 17).

[0156] FIG. 24 plots the effect of gap size upon the reactor capacity, where the number of 200-can reactors needed to reach 10,000 barrels per day capacity is plotted versus catalyst gas hourly space velocity. As the wall and inlet temperatures are increased on 1 °C increments, the number of reactors needed to reach 10,000 barrels per day can vary with flow rate. However, gaps greater than 3.0 mm can increase the total number of tubes by hundreds and up to a thousand reactor tubes, especially at lower gas hourly space velocities. Moreover, the number of reactor tubes (for 200 cans in series reactors) needed for a 3 millimeter heat exchanger gap is 85 to 254 more than the total number of reactor tubes with a 2 millimeter gap. There is a 5% to 13% decrease in the number of installed 200-can reactor tubes when the annular gap is 2 mm relative to 3 mm. It is anticipated that the capital cost will reduce proportionally with a reduction in cans with an expected 0.6 to 0.7 power law used for an economy of scale. That is to say, the 5 to 13% reduction in tubes might reduce the capital cost of a plant by 3 to 10%. This savings is significant considering the high capital cost for a commercial Fischer Tropsch facility.

[0157] FIG. 25 plots the carbon monoxide conversion for each of the 200 can reactors versus gas hourly space velocity. The 200 can reactors show a different trend in carbon monoxide conversion than the 80 can reactors in Example 4 (FIG. 19) - the conversion values are more or less maintained as gas hourly space values increase while those of the 80 can reactors fall with the same increases in gas hourly space velocity. Moreover, only gaps 2 millimeters and smaller are consistently over the 70% carbon monoxide conversion, whereas use of Gamlin's gaps never achieve 70% carbon monoxide conversion at any of the catalyst gas hourly space velocities studied.

[0158] FIG. 26 plots the C ₅ + yields for the 200 can reactors versus gas hourly space velocity. For the 2-millimeter gap carriers, the change in C ₅ + yield is slightly less than for the 80 can reactors of Example 4 (FIG. 22). [0159] FIG. 27 plots the temperature difference between the maximum temperature in the reactor and the wall temperatures. The temperature differences for all but the 1 millimeter gap are on the order of 40 °C to 70 °C range. The reactor exit dew point temperature for all these cases were below the wall temperature at 500 psig.

[0160] FIG. 28 shows the change in pressure drop for the 200 can reactors due to changes in the catalyst gas hourly space velocity. For heat exchanger gaps from 2 to 5 millimeters, the transition from 4080 to 6077 gas hourly space velocity results in a pressure drop increase from roughly 70 psig to 200 psig. The one-millimeter heat exchanger gap increased the pressure drop substantially, not allowing the 6077 gas hourly space velocity condition to converge due to excessive pressure drop.

[0161] Example 6

[0162] A study was performed to determine the importance of the inner gap on distribution of flow through the annular container, assessed as the gaseous stream exited the perforated outer wall. The annular container was modeled to be 5 inches from the top surface to the bottom surface and the reactor inlet pressure was 350 psig. 3 cases were analyzed: Case 1 with an inner gap ("skirt gap") of 5 mm and an annular gap ("heat exchanger gap") of 5 mm, Case 2 with an inner gap of 5 mm and an annular gap of 1 mm; and Case 3 with an inner gap of 1 mm and an annular gap of 1 mm. The annular gap was varied to assess whether downstream constriction noticeably affected the flow distribution exiting the perforated outer wall. FIG. 29 illustrates the flow distribution of the gaseous stream exiting the perforated outer wall for each of the three cases. Notably, Case 1 and Case 2 with the 5 mm inner gap exhibited a very similar flow distribution across the length of the can. Case 3 was markedly different, showing highly problematic flow maldistribution through the can due to the 1 mm inner gap. Clearly, flow distribution in the catalyst bed was affected by the inner gap while size of the annular gap played a minimal role.

[0163] Example 7

[0164] The heat transfer between the fluid exiting the catalyst portion of the can and the tube wall is enhanced by increasing the specific area per unit volume with the use of extended heat transfer area or wall features that indent or protrude to improve heat transfer. For tube outer diameters in the range of about 40 mm to about 65 mm with a straight annular space for heat transfer, the SA/V will decrease with increasing annular space as illustrated in Tables 3 & 4 below.

Table 3.

Table 4.

[0165] For annular spaces that are larger than 3 mm and for tube outer diameters in the range of about 40 to about 65 mm, an SA/V larger than 400 m ^"1 may be provided via heat conducting surfaces. The high SA/V provides for efficient heat removal even when the overall annular gap when measured excluding the heat conducting surfaces is greater than 3 mm. It is anticipated that increasing the SA/V by (1) skirt features that indent or protrude, and/or (2) a network of heat conducting surfaces in conductive thermal contact with a portion the inner tube wall and (i) a portion of the perforated outer wall and/or (ii) a portion of the outer surface of the skirt will significantly increase heat transfer providing for significantly higher CO conversion and significantly higher C ₅ + hydrocarbon selectivity. It is also anticipated that inclusion of plurality of fins extending radially from the perforated inner wall to the perforated outer wall of the catalyst carrier (see Rl of FIG. 30) or through the perforated outer wall to the inner surface of the skirt (see R2 of FIG. 30) will also facilitate heat transfer.

[0166] The present technology is not to be limited in terms of the particular figures and examples described herein, which are intended as single illustrations of individual aspects of the present technology. Many modifications and variations of this present technology can be made without departing from its spirit and scope, as will be apparent to those skilled in the art. Functionally equivalent methods within the scope of the present technology, in addition to those enumerated herein, will be apparent to those skilled in the art from the foregoing descriptions. Such modifications and variations are intended to fall within the scope of the appended claims. It is to be understood that this present technology is not limited to particular methods, reagents, compounds, compositions, or labeled compounds, which can, of course, vary. It is also to be understood that the terminology used herein is for the purpose of describing particular aspects only, and is not intended to be limiting.

[0167] The embodiments, illustratively described herein may suitably be practiced in the absence of any element or elements, limitation or limitations, not specifically disclosed herein. Thus, for example, the terms "comprising," "including," "containing," etc. shall be read expansively and without limitation. Additionally, the terms and expressions employed herein have been used as terms of description and not of limitation, and there is no intention in the use of such terms and expressions of excluding any equivalents of the features shown and described or portions thereof, but it is recognized that various modifications are possible within the scope of the claimed technology. Additionally, the phrase "consisting essentially of will be understood to include those elements specifically recited and those additional elements that do not materially affect the basic and novel characteristics of the claimed technology. The phrase "consisting of excludes any element not specified. [0168] In addition, where features or aspects of the disclosure are described in terms of Markush groups, those skilled in the art will recognize that the disclosure is also thereby described in terms of any individual member or subgroup of members of the Markush group. Each of the narrower species and sub-generic groupings falling within the generic disclosure also form part of the invention. This includes the generic description of the invention with a proviso or negative limitation removing any subject matter from the genus, regardless of whether or not the excised material is specifically recited herein.

[0169] All publications, patent applications, issued patents, and other documents (for example, journals, articles and/or textbooks) referred to in this specification are herein incorporated by reference as if each individual publication, patent application, issued patent, or other document was specifically and individually indicated to be incorporated by reference in its entirety. Definitions that are contained in text incorporated by reference are excluded to the extent that they contradict definitions in this disclosure.

[0170] The present technology may include, but is not limited to, the features and combinations of features recited in the following lettered paragraphs, it being understood that the following paragraphs should not be interpreted as limiting the scope of the claims as appended hereto or mandating that all such features must necessarily be included in such claims:

A. A process for the conversion of synthesis gas, the process comprising contacting in a tubular reactor a gaseous stream comprising synthesis gas with a solid synthesis catalyst to produce a synthetic product,

wherein the tubular reactor comprises

a reactor inlet in fluid communication with one or more reactor tubes wherein each reactor tube comprises a tube inlet, a tube outlet located downstream of the tube inlet, an inner tube wall, an outer tube wall, and one or more catalyst carriers within the reactor tube,

a reactor outlet located downstream of the reactor inlet in fluid communication with the one or more reactor tubes, and

a cooling medium in contact with the one or more reactor tubes; wherein the catalyst carrier comprises: an annular container, the annular container comprising a perforated inner wall defining a central tube, a perforated outer wall, a top surface closing the annular container and a bottom surface closing the annular container; wherein the central tube has a top and bottom according to the top surface and the bottom surface and the annular container holds the solid synthesis catalyst;

a surface closing the bottom of the central tube;

a skirt extending upwardly from the perforated outer wall of the annular

container from a position at or near the bottom surface of the annular container to a position below the location of a seal; and the seal located at or near the top surface and extending from the container by a distance which extends beyond an outer surface of the skirt;

wherein the process comprises at least one of the following:

(1) an average annular space between the outer surface of the skirt and the inner tube wall from about 1.3 mm to about 2.6 mm; and

(2) a total combined surface area of the outer surface of the skirt and the inner tube wall, and optionally a heat transfer structure per volume of the solid synthesis catalyst in the catalyst carrier (the "SA/V") from about 400 m ² /m ³ to about 40,000 m ² /m ³ ;

wherein the heat transfer structure comprises at least one of

(i) a plurality of fins extending radially from the perforated inner wall to the perforated outer wall;

(ii) a plurality of fins extending radially from the perforated inner wall to an inner surface of the skirt; and

(iii) a network of heat conducting surfaces in conductive thermal contact with a portion the inner tube wall and a portion of the outer surface of the skirt; and

the process optionally further comprising introducing the gaseous stream through the reactor inlet at a pressure from about 250 psig to about 1,000 psig with a ratio of H ₂ /CO in the synthesis gas from about 1.6 to about 2.0.

The process of Paragraph A, wherein the contacting step further comprises maintaining at least about 55% carbon monoxide conversion per pass in the one or more reactor tubes; and

wherein the solid synthesis catalyst possesses a normalized CO conversion loss of less than about 0.4% after 3 weeks of continued processing.

C. The process of Paragraph A or Paragraph B, wherein the process comprises

introducing the gaseous stream through the reactor inlet at a pressure from about 500 psig to about 1,000 psig.

D. The process of any one of Paragraphs A-C, wherein the process comprises

introducing the gaseous stream through the reactor inlet at a pressure from about 500 psig to about 750 psig.

E. The process of any one of Paragraphs A-D, wherein a catalyst gas hourly space

velocity of the gaseous stream in the tubular reactor is from about 5,000 hr ^-1 to about 15,000 hr ^"1 .

F. The process of any one of Paragraphs A-E, wherein the catalyst gas hourly space velocity of the gaseous stream in the tubular reactor is greater than about 5,500 hr ^_1 .

G. The process of any one of Paragraphs A-F, wherein the temperature of the gaseous stream at the reactor inlet is about 160 °C to about 240 °C.

H. The process of any one of Paragraphs A-G, wherein the temperature of the gaseous stream at the reactor inlet is about 170 °C to about 190 °C.

I. The process of any one of Paragraphs A-H, wherein the flow of the gaseous steam is counter to gravity.

J. The process of any one of Paragraphs A-I, wherein the flow of the gaseous stream is with gravity.

K. The process of any one of Paragraphs A- J, wherein the synthetic product comprises hydrocarbons. L. The process of any one of Paragraphs A-K, wherein the product comprises C ₅ + hydrocarbons.

M. The process of any one of Paragraphs A-L, wherein the solid synthesis catalyst is a particulate Fischer-Tropsch catalyst.

N. The process of any one of Paragraphs A-M, wherein the solid synthesis catalyst is a particulate Fischer-Tropsch catalyst comprising Co.

O. The process of any one of Paragraphs A-N, wherein the solid synthesis catalyst is a particulate Fischer-Tropsch catalyst having a Co surface area per unit of packed bed volume of about 15 m ² /mL to about 65 m ² /mL.

P. The process of any one of Paragraphs A-O, wherein the solid synthesis catalyst is a particulate Fischer-Tropsch catalyst having a weight average diameter from about 100 micrometers (μιη) to about 1 millimeter (mm).

Q. The process of any one of Paragraphs A-P, wherein the solid synthesis catalyst is a particulate Fischer-Tropsch catalyst having a Co loading from about 20 wt% to about 56 wt%.

R. The process of any one of Paragraphs A-Q, wherein the cooling medium temperature is about 160 °C to about 240 °C.

S. The process of any one of Paragraphs A-R, wherein a difference between the reactor inlet temperature and the cooling medium temperature is less than about 40 °C.

T. The process of any one of Paragraphs A-S, wherein a temperature difference between the cooling medium and the product exiting the tube outlet is maintained between about 20 °C and about 80 °C.

U. The process of any one of Paragraphs A-T, wherein a temperature difference between the cooling medium and the solid synthesis catalyst is maintained between about 20 °C and about 80 °C. V. The process of any one of Paragraphs A-U, wherein each reactor tube comprises from about 80 to about 200 catalyst carriers within the reactor tube.

W. The process of any one of Paragraphs A-V, wherein the one or more reactor tubes each independently have a diameter of about 30 mm to about 300 mm.

X. The process of any one of Paragraphs A-W, wherein the one or more reactor tubes each independently have a diameter of about 40 mm to about 65 mm.

Y. The process of any one of Paragraphs A-X, wherein an average inner gap between the skirt and the perforated outer wall of the annular container is from about 2.0 mm to about 8.0 mm.

Z. The process of any one of Paragraphs A-Y, wherein

a first carrier group comprises at least two catalyst carriers with the average annular space from about 1.3 mm to about 2.6 mm;

a second carrier group comprises at least two catalyst carriers with a larger average annular space than the first carrier group between about 1.5 mm to about 10.0 mm.

AA. The process of Paragraph Z, wherein the second carrier group contains a solid

synthesis catalyst with higher activity than the solid synthesis catalyst of the first carrier group.

AB. The process of Paragraphs Z or Paragraph AA, wherein the solid synthesis catalyst of the second carrier group has at least 10% higher activity than the solid synthesis catalyst of the first carrier group.

AC. The process of any one of Paragraphs Z-AB, wherein the catalyst carriers of the

second carrier group comprise an average annular space between about 3.0 mm to about 10.0 mm.

AD. The process of any one of Paragraphs A- AC, wherein

a first carrier group comprises at least two catalyst carriers with the average annular space from about 1.5 mm to about 10.0 mm, and a second carrier group comprises at least two catalyst carriers with an average annular space smaller than the first carrier group between about 1.3 mm to about 2.6 mm.

AE. The process of Paragraph AD, wherein the first carrier group contains a solid

synthesis catalyst with higher activity than the solid synthesis catalyst of the second carrier group.

AF. The process of Paragraph AD or Paragraph AE, wherein the solid synthesis catalyst of the first carrier group has at least 10% higher activity than the solid synthesis catalyst of the second carrier group.

AG. The process of any one of Paragraphs AD-AF, wherein the catalyst carriers of the first carrier group comprise an average annular space between about 3.0 mm to about 10.0 mm.

AH. The process of any one of Paragraphs AD-AG, further comprising

a third carrier group comprising at least two catalyst carriers with an average annular space larger than the second carrier group between about 1.5 mm to about 10.0 mm.

AI. The process of Paragraph AH, wherein the third carrier group contains a solid

synthesis catalyst with higher activity than the solid synthesis catalyst of the second carrier group.

AJ. The process of any one of Paragraphs A-Y, wherein

a first carrier group comprises at least two catalyst carriers; and a second carrier group comprises at least two catalyst carriers; wherein the first carrier group contains a solid synthesis catalyst with higher activity than the solid synthesis catalyst of the second carrier group. AK. The process of Paragraph AJ, further comprising

a third carrier group comprising at least two catalyst carriers wherein the third carrier group contains a solid synthesis catalyst with higher activity than the solid synthesis catalyst of the second carrier group.

AL. The process of any one of Paragraphs A-AK, wherein the tubular reactor comprises at least 100 reactor tubes.

AM. The process of any one of Paragraphs A-AL, wherein the skirt comprises steel,

aluminum, copper, an alloy thereof, or a combination of any two or more thereof.

AN. The process of any one of Paragraphs A- AM, wherein the heat transfer structure comprises aluminum, copper, an alloy thereof, or a combination of any two or more thereof.

AO. The process of any one of Paragraphs A- AN, wherein the heat transfer structure

comprises a combination of steel and aluminum, steel and copper, aluminum and copper, or steel, aluminum, and copper.

AP. The process of any one of Paragraphs A- AO, wherein network of heat conducting surfaces comprises a random network of heat conducting surfaces.

AQ. The process of any one of Paragraphs A-AP, wherein the network of heat conducting surfaces comprises an ordered network of heat conducting surfaces.

AR. The process of any one of Paragraphs A-AQ, wherein the network of heat conducting surfaces define a plurality of channels.

AS. The process of Paragraph AR, wherein each channel of the plurality of channels

independently has a channel diameter from about 0.01 mm to about 8 mm.

AT. The process of any one of Paragraphs A-AS, wherein a distance between the fins extending radially from the perforated inner wall as measured at the perforated inner wall is from about 0.01 mm to about 10 mm. AU. A process for the conversion of synthesis gas, the process comprising contacting in a tubular reactor a gaseous stream comprising synthesis gas with a solid synthesis catalyst to produce a product comprising hydrocarbons,

wherein the tubular reactor comprises

a reactor inlet in fluid communication with one or more reactor tubes wherein each reactor tube comprises a tube inlet, a tube outlet located downstream of the tube inlet, an inner tube wall, an outer tube wall, and one or more catalyst carriers within the reactor tube,

a reactor outlet located downstream of the reactor inlet in fluid communication with the one or more reactor tubes, and

a cooling medium in contact with the one or more reactor tubes; wherein the catalyst carrier comprises:

an annular container, the annular container comprising a perforated inner wall defining a central tube, a perforated outer wall, a top surface closing the annular container and a bottom surface closing the annular container; wherein the central tube has a top and bottom according to the top surface and the bottom surface and the annular container holds the solid synthesis catalyst;

a surface closing the top of the central tube;

a skirt extending downwardly from the perforated outer wall of the annular container from a position at or near the top surface of the annular container to a position above the location of a seal; and

the seal located at or near the bottom surface and extending from the container by a distance which extends beyond an outer surface of the skirt.

AV. The process of Paragraph AU, wherein the process comprises at least one of the

following:

(1) an average annular space between the outer surface of the skirt and the inner tube wall from about 1.3 mm to about 2.6 mm; and

(2) a total combined surface area of the outer surface of the skirt and the inner tube wall, and optionally a heat transfer structure per volume of the solid synthesis catalyst in the catalyst carrier (the "SA/V") from about 400 m ² /m ³ to about 40,000 m ² /m ³ ;

wherein the heat transfer structure comprises at least one of

(i) a plurality of fins extending radially from the perforated inner wall to the perforated outer wall;

(ii) a plurality of fins extending radially from the perforated inner wall to an inner surface of the skirt; and

(iii) a network of heat conducting surfaces in conductive thermal contact with a portion the inner tube wall and a portion of the outer surface of the skirt.

AW. The process of Paragraph AU or Paragraph AV, the process further comprising

introducing the gaseous stream through the reactor inlet at a pressure from about 250 psig to about 1,000 psig with a ratio of H ₂ /CO in the synthesis gas from about 1.6 to about 2.0.

AX. The process of any one of Paragraphs AU-AW, wherein the contacting step further comprises

maintaining at least about 55% carbon monoxide conversion per pass in the one or more reactor tubes; and

wherein the solid synthesis catalyst possesses a normalized CO conversion loss of less than about 0.4% after 3 weeks of continued processing.

AY. The process of any one of Paragraphs AU-AX, wherein the process comprises

introducing the gaseous stream through the reactor inlet at a pressure from about 500 psig to about 1,000 psig.

AZl . The process of any one of Paragraphs AU-AY, wherein the process comprises

introducing the gaseous stream through the reactor inlet at a pressure from about 500 psig to about 750 psig. BA. The process of any one of Paragraphs AU-AZ, wherein a catalyst gas hourly space velocity of the gaseous stream in the tubular reactor is from about 5,000 hr ^-1 to about 15,000 hr ^"1 .

BB. The process of any one of Paragraphs AU-BA, wherein the catalyst gas hourly space velocity of the gaseous stream in the tubular reactor is greater than about 5,500 hr ^_1 .

BC. The process of any one of Paragraphs AU-BB, wherein the temperature of the gaseous stream at the reactor inlet is about 160 °C to about 240 °C.

BD. The process of any one of Paragraphs AU-BC, wherein the temperature of the gaseous stream at the reactor inlet is about 170 °C to about 190 °C.

BE. The process of any one of Paragraphs AU-BD, wherein the flow of the gaseous steam is counter to gravity.

BF. The process of any one of Paragraphs AU-BE, wherein the flow of the gaseous stream is with gravity.

BG. The process of any one of Paragraphs AU-BF, wherein the synthetic product

comprises hydrocarbons.

BH. The process of any one of Paragraphs AU-BG, wherein the product comprises C ₅ + hydrocarbons.

BI. The process of any one of Paragraphs AU-BH, wherein the solid synthesis catalyst is a particulate Fischer-Tropsch catalyst.

BJ. The process of any one of Paragraphs AU-BI, wherein the solid synthesis catalyst is a particulate Fischer-Tropsch catalyst comprising Co.

BK. The process of any one of Paragraphs AU-BJ, wherein the solid synthesis catalyst is a particulate Fischer-Tropsch catalyst having a Co surface area per unit of packed bed volume of about 15 m ² /mL to about 65 m ² /mL. BL. The process of any one of Paragraphs AU-BK, wherein the solid synthesis catalyst is a particulate Fischer-Tropsch catalyst having a weight average diameter from about 100 micrometers (μιη) to about 1 millimeter (mm).

BM. The process of any one of Paragraphs AU-BL, wherein the solid synthesis catalyst is a particulate Fischer-Tropsch catalyst having a Co loading from about 20 wt% to about 56 wt%.

BN. The process of any one of Paragraphs AU-BM, wherein the cooling medium

temperature is about 160 °C to about 240 °C.

BO. The process of any one of Paragraphs AU-BN, wherein a difference between the reactor inlet temperature and the cooling medium temperature is less than about 40 °C.

BP. The process of any one of Paragraphs AU-BO, wherein a temperature difference between the cooling medium and the product exiting the tube outlet is maintained between about 20 °C and about 80 °C.

BQ. The process of any one of Paragraphs AU-BP, wherein a temperature difference

between the cooling medium and the solid synthesis catalyst is maintained between about 20 °C and about 80 °C.

BR. The process of any one of Paragraphs AU-BQ, wherein each reactor tube comprises from about 80 to about 200 catalyst carriers within the reactor tube.

BS. The process of any one of Paragraphs AU-BR, wherein the one or more reactor tubes each independently have a diameter of about 30 mm to about 300 mm.

BT. The process of any one of Paragraphs AU-BS, wherein the one or more reactor tubes each independently have a diameter of about 40 mm to about 65 mm.

BU. The process of any one of Paragraphs AU-BT, wherein an average annular space between the outer surface of the skirt and the inner surface of the tube wall is from about 1.3 mm to about 10.0 mm. BV. The process of any one of Paragraphs AU-BU, wherein an average annular space between the outer surface of the skirt and the inner surface of the tube wall is from about 1.3 mm to about 2.6 mm.

BW. The process of any one of Paragraphs AU-BV, wherein an average inner gap between the skirt and the perforated outer wall of the annular container is from about 2.0 mm to about 8.0 mm.

BX. The process of any one of Paragraphs AU-BW, wherein

a first carrier group comprises at least two catalyst carriers with the average annular space from about 1.3 mm to about 2.6 mm;

a second carrier group comprises at least two catalyst carriers with a larger average annular space than the first carrier group between about 1.5 mm to about 10.0 mm.

BY. The process of Paragraph BX, wherein the second carrier group contains a solid

synthesis catalyst with higher activity than the solid synthesis catalyst of the first carrier group.

BZ. The process of Paragraph BX or Paragraph BY, wherein the solid synthesis catalyst of the second carrier group has at least 10% higher activity than the solid synthesis catalyst of the first carrier group.

CA. The process of any one of Paragraphs AU-BZ, wherein the catalyst carriers of the second carrier group comprise an average annular space between about 3.0 mm to about 10.0 mm.

CB. The process of any one of Paragraphs AU-CA, wherein

a first carrier group comprises at least two catalyst carriers with the average annular space from about 1.5 mm to about 10.0 mm, and

a second carrier group comprises at least two catalyst carriers with an average annular space smaller than the first carrier group between about 1.3 mm to about 2.6 mm. CC. The process of Paragraph CB, wherein the first carrier group contains a solid synthesis catalyst with higher activity than the solid synthesis catalyst of the second carrier group.

CD. The process of Paragraph CB or Paragraph CC, wherein the solid synthesis catalyst of the first carrier group has at least 10% higher activity than the solid synthesis catalyst of the second carrier group.

CE. The process of any one of Paragraphs CB-CD, wherein the catalyst carriers of the first carrier group comprise an average annular space between about 3.0 mm to about 10.0 mm.

CF. The process of any one of Paragraphs CB-CE, further comprising

a third carrier group comprising at least two catalyst carriers with an average annular space larger than the second carrier group between about 1.5 mm to about 10.0 mm.

CG. The process of claim CF, wherein the third carrier group contains a solid synthesis catalyst with higher activity than the solid synthesis catalyst of the second carrier group.

CH. The process of any one of Paragraphs AU-CG, wherein

a first carrier group comprises at least two catalyst carriers; and a second carrier group comprises at least two catalyst carriers; wherein the first carrier group contains a solid synthesis catalyst with higher activity than the solid synthesis catalyst of the second carrier group.

CI. The process of Paragraph CH, further comprising

a third carrier group comprising at least two catalyst carriers wherein the third carrier group contains a solid synthesis catalyst with higher activity than the solid synthesis catalyst of the second carrier group.

The process of any one of Paragraphs AU-CI, wherein the tubular reactor compri at least 100 reactor tubes. CK. The process of any one of Paragraphs AU-CJ, wherein the skirt comprises steel, aluminum, copper, an alloy thereof, or a combination of any two or more thereof.

CL. The process of any one of Paragraphs AU-CK, wherein the heat transfer structure comprises aluminum, copper, an alloy thereof, or a combination of any two or more thereof.

CM. The process of any one of Paragraphs AU-CL, wherein the heat transfer structure comprises a combination of steel and aluminum, steel and copper, aluminum and copper, or steel, aluminum, and copper.

CN. The process of any one of Paragraphs AU-CM, wherein network of heat conducting surfaces comprises a random network of heat conducting surfaces.

CO. The process of any one of Paragraphs AU-CN, wherein the network of heat

conducting surfaces comprises an ordered network of heat conducting surfaces.

CP. The process of any one of Paragraphs AU-CO, wherein the network of heat

conducting surfaces define a plurality of channels.

CQ. The process of Paragraph CP, wherein each channel of the plurality of channels independently has a channel diameter from about 0.01 mm to about 8 mm.

CR. The process of any one of Paragraphs AU-CQ, wherein a distance between the fins extending radially from the perforated inner wall as measured at the perforated inner wall is from about 0.01 mm to about 10 mm.

[0171] Other embodiments are set forth in the following claims, along with the full scope of equivalents to which such claims are entitled.