POLYFUNCTIONAL CATALYSTS
20200031735 ยท 2020-01-30
Inventors
- Anurag Kumar (Minneapolis, MN, US)
- Neil Razdan (Minneapolis, MN, US)
- Aditya Bhan (Minneapolis, MN, US)
Cpc classification
B01J29/48
PERFORMING OPERATIONS; TRANSPORTING
C07C15/00
CHEMISTRY; METALLURGY
C07C2529/48
CHEMISTRY; METALLURGY
B01J2229/18
PERFORMING OPERATIONS; TRANSPORTING
C07C15/00
CHEMISTRY; METALLURGY
C07C2/76
CHEMISTRY; METALLURGY
C07C2/76
CHEMISTRY; METALLURGY
International classification
B01J29/48
PERFORMING OPERATIONS; TRANSPORTING
Abstract
Embodiments of the present disclosure describe catalysts, methods of preparing catalysts, methods of forming hydrocarbons using the catalysts, and the like.
Claims
1. A method of producing aromatics, comprising: contacting a methane-containing feed stream with a pre-carburized catalyst to form aromatics and hydrogen, wherein the pre-carburized catalyst is disposed within a reactor and comprises a carbidic form of an active metal supported on a proton form of a zeolite, wherein the reactor further comprises a hydrogen-accepting component that removes at least a portion of said hydrogen.
2. The method of claim 1, wherein the contacting proceeds at or to a temperature of at least about 950 K.
3. The method of claim 1, wherein the contacting proceeds under about atmospheric pressure.
4. The method of claim 1, wherein single-pass methane conversion is greater than equilibrium methane conversion under the same reaction conditions.
5. The method of claim 1, wherein an average particle size of the pre-carburized catalyst is in the range of about 180 m to about 425 m.
6. The method of claim 1, wherein an average particle size of the hydrogen-accepting component is in the range of about 180 m to about 425 m.
7. The method of claim 1, wherein the reactor is a staged and stratified reactor comprising alternating layers of the pre-carburized catalyst and hydrogen-accepting component.
8. The method of claim 1, wherein the hydrogen-accepting component is positioned upstream from the pre-carburized catalyst.
9. The method of claim 1, wherein the hydrogen-accepting component is positioned downstream from the pre-carburized catalyst.
10. The method of claim 1, wherein the hydrogen-accepting component is positioned both downstream and upstream from the pre-carburized catalyst.
11. The method of claim 1, wherein the pre-carburized catalyst and hydrogen-accepting component are disposed within the reactor as an interparticle mixture.
12. The method of claim 1, wherein the pre-carburized catalyst comprises an active metal of the formula MC.sub.x, where M is the active metal, C is carbon, and x is at least 0.01.
13. The method of claim 1, wherein the active metal is selected from the group consisting of molybdenum, vanadium, chromium manganese, zinc, iron, cobalt, nickel, copper, gallium, germanium, niobium, molybdenum, ruthenium, rhodium, silver, tantalum, tungsten, rhenium, platinum, or lead.
14. The method of claim 1, wherein the hydrogen-accepting component is selected from the group consisting of zirconium, titanium, niobium, tantalum, hafnium, vanadium, or zinc.
15. The method of claim 1, wherein the aromatics are formed without an induction period.
16. The method of claim 1, wherein the aromatics include one or more of benenze, naphthalene, toluene, and xylene.
17. The method of claim 1, further comprising regenerating the hydrogen-accepting component by thermal treatment under inert flow.
18. The method of claim 1, further comprising exposing, prior to the contacting, a precatalyst comprising an active metal supported on a proton form of a zeolite to methane to obtain the pre-carburized catalyst.
19. A staged and stratified catalyst for non-oxidative methane dehydroaromatization to form aromatics and hydrogen from methane, the catalyst comprising: alternating layers of a pre-carburized catalyst and a hydrogen-accepting component, wherein the pre-carburized catalyst comprises a carbidic form of an active metal supported on a proton form of a zeolite, wherein the hydrogen-accepting component absorbs at least a portion of said hydrogen.
20. An interparticle catalyst mixture for non-oxidative methane dehydroaromatization to form aromatics and hydrogen from methane, the catalyst comprising: a mixture comprising particles of a pre-carburized catalyst and particles of a hydrogen-accepting component, wherein the pre-carburized catalyst comprises a carbidic form of an active metal supported on a proton form of a zeolite, wherein the hydrogen-accepting component absorbs at least a portion of said hydrogen.
Description
BRIEF DESCRIPTION OF DRAWINGS
[0010] This written disclosure describes illustrative embodiments that are non-limiting and non-exhaustive. In the drawings, which are not necessarily drawn to scale, like numerals describe substantially similar components throughout the several views. Like numerals having different letter suffixes represent different instances of substantially similar components. The drawings illustrate generally, by way of example, but not by way of limitation, various embodiments discussed in the present document.
[0011] Reference is made to illustrative embodiments that are depicted in the figures, in which:
[0012]
[0013]
[0014]
[0015]
[0016]
[0017]
[0018]
[0019]
[0020]
[0021]
[0022]
[0023]
[0024]
[0025]
[0026]
[0027]
[0028]
[0029]
[0030]
[0031]
[0032]
[0033]
[0034]
[0035]
[0036]
[0037]
[0038]
[0039]
[0040]
[0041]
[0042]
DETAILED DESCRIPTION
[0043] The invention of the present disclosure relates to catalysts and/or pre-catalysts for, among other things, methane dehydroaromatization with transient absorptive removal of hydrogen. In particular, the invention of the present disclosure relates to catalysts and/or pre-catalysts that increase yields of desired products in non-oxidative methane dehydrogenation on polyfunctional formulations (e.g., catalysts and/or pre-catalysts) by introducing an additional function that scavenges byproduct hydrogen thereby lifting thermodynamic constraints on attainable yields and increasing rates of products formation. The use of polyfunctional formulations (e.g., catalysts and/or pre-catalysts) enhance yield of aromatic products in methane conversion over, for example, carbidic forms of metal (e.g., Mo and W, among others) catalysts supported on porous supports (e.g., such as proton-form zeolites). Staged and stratified catalyst beds, and interparticle mixtures, among others, may be used. See, for example,
[0044] For example, non-oxidative conversion of methane by pyrolysis reactions to produce aromatics may occur with near equilibrium yield on carbidic forms of molybdenum encapsulated in proton form zeolites at temperatures 700 C. The dehydroaromatization of methane to benzene (with 70% selectivity on a carbon basis) at these temperatures may be restricted by equilibrium considerations (6CH.sub.4.Math.C.sub.6H.sub.6+9H.sub.2) to 10% single pass conversion due to abundant hydrogen formed in the catalyst bed. In-situ removal of hydrogen by addition of zirconium reported herein may be used as an effective strategy to overcome these thermodynamic restrictions on single pass conversion. The improvement in single pass conversion of methane and yield of aromatics may be attained without any deleterious effects on selectivity to desirable products. The absorbent function can be regenerated by thermal treatment in inert flow and the regenerated absorbent can be used to attain higher than 10% conversion in successive reaction-regeneration cycles.
[0045] Embodiments of the present disclosure describe a catalyst comprising a first component and a second component. The first component may include a hydrogen-accepting component. The second component may include a porous support and an active metal on a surface and/or pores of the porous support. In other embodiments, the first component may include a porous support and an active metal on a surface and/or pores of the porous support; and the second component may include a hydrogen-accepting component.
[0046] The first component and the second component may be arranged in any of a variety of configurations (e.g., staged and/or stratified catalyst beds, interparticle mixtures, intraparticle mixtures, etc.). In many embodiments, the first component and the second component are arranged such that they are in physical contact, or immediate or close proximity. For example, in an embodiment, the polyfunctional catalyst may be provided as alternating stratified layers of a first component and a second component. In an embodiment, the polyfunctional catalyst may be provided as interparticle physical mixtures of a first component and a second component. In an embodiment, the polyfunctional catalyst may be provided as intraparticle physical mixtures of a first component and a second component. In an embodiment, the polyfunctional catalyst may be provided in which the first component or second component encapsulates the other component (e.g., the second component or first component, respectively).
[0047] The hydrogen-accepting component may include any element, compound, and/or complex capable of accepting a hydrogen. For example, the hydrogen-accepting component may include any element, compound, and/or complex capable of accepting and/or bonding to a hydrogen to form a hydride. The hydrogen-accepting component may include at least one or more of alkali metals, alkaline earth metals, transition metals, and metalloids. In many embodiments, the hydrogen-accepting component includes at least transition metals. The transition metals may include, but are not limited to, one or more of zirconium, titanium, niobium, tantalum, hafnium, vanadium, and zinc. In preferred embodiments, the hydrogen-accepting component includes zirconium. The hydrogen-accepting component may be provided in any form, such as particles and/or powders. For example, a particle size of the hydrogen-accepting component may range from about 180 m to about 425 m. In other embodiments, a particle size of the hydrogen-accepting component may be less than about 180 m and/or greater than about 425 m.
[0048] The active metal may include any metal that exhibits catalytic activity upon being contacting with a fluid composition containing methane (e.g., under methane dehydroaromatization conditions, non-oxidative conversion of methane, etc.). For example, the active metal may include one or more of molybdenum, vanadium, chromium manganese, zinc, iron, cobalt, nickel, copper, gallium, germanium, niobium, molybdenum, ruthenium, rhodium, silver, tantalum, tungsten, rhenium, platinum, and lead. In many embodiments, the active metal includes a carbidic form of the active metal. For example, in a preferred embodiment, the active metal includes a carbidic form of molybdenum. In another preferred embodiment, the active metal includes a carbidic form of tungsten. In an embodiment, the pre-carburized catalyst comprises an active metal of the formula MC.sub.x, where M is the active metal, C is carbon, and x is at least 0.01.
[0049] The porous support may include any porous material. In many embodiments, the porous support includes an inorganic oxide support. For example, the porous support may include one or more of zeolites, non-zeolitic molecular sieves, silica, alumina, zirconia, titania, yttria, ceria, and rare earth metal oxides. In many embodiments, the porous support includes zeolites, preferably a proton form of zeolites. The zeolites may include one or more of ZSM-5, ZSM-8, ZSM-11, ZSM-12, ZSM-22, and ZSM-35. In a preferred embodiment, the porous support includes ZSM-5. In a more preferred embodiment, the porous support includes a proton form of ZSM-5.
[0050] In an embodiment, the polyfunctional catalyst comprises a first component and a second component, wherein the first component includes a zirconium particle and the second component includes a proton form of a zeolite and a carbidic form of Mo or W on a surface and/or pores of the proton form of the zeolite.
[0051] In an embodiment, the polyfunctional catalyst is provided as a staged and stratified catalyst for non-oxidative methane dehydroaromatization to form aromatics and hydrogen from methane. For example, in certain embodiments, the staged and stratified catalyst comprises alternating layers of a pre-carburized catalyst and a hydrogen-accepting component, wherein the pre-carburized catalyst comprises a carbidic form of an active metal supported on a proton form of a zeolite, wherein the hydrogen-accepting component absorbs at least a portion of said hydrogen.
[0052] In an embodiment, the hydrogen-accepting component is positioned upstream from the pre-carburized catalyst. In an embodiment, the hydrogen-accepting component is positioned downstream from the pre-carburized catalyst. In an embodiment, the hydrogen-accepting component is positioned both downstream and upstream from the pre-carburized catalyst.
[0053] In an embodiment, a polyfunctional catalyst is provided in a continuous flow configuration. For example, in certain embodiments, the polyfunctional catalyst comprises alternating layers of a pre-carburized catalyst and a hydrogen-accepting component, wherein the hydrogen-accepting component continuously flows (e.g., as a moving bed of particles) to a regeneration unit, wherein accepted or absorbed hydrogen is removed or released from the hydrogen-accepting component and the hydrogen-accepting component is regenerated to full hydrogen-storage capacity. The hydrogen-accepting component can be thus be regenerated and reused in one or more reaction cycles.
[0054] In an embodiment, a polyfunctional catalyst is provided in a fluidized bed reactor, wherein the discrepancy in density of the two components is leveraged to separate and regenerate each component cyclically. For example, in certain embodiments, the pre-carburized catalyst and hydrogen-accepting component can be separated based on differences or discrepancies in density.
[0055] In an embodiment, the polyfunctional catalyst is provided as an interparticle mixture for non-oxidative methane dehydroaromatization to form aromatics and hydrogen from methane. For example, in certain embodiments, the interparticle mixture or interparticle catalyst mixture comprises a mixture comprising particles of a pre-carburized catalyst and particles of a hydrogen-accepting component, wherein the pre-carburized catalyst comprises a carbidic form of an active metal supported on a proton form of a zeolite, wherein the hydrogen-accepting component absorbs at least a portion of said hydrogen.
[0056]
[0057] The step 201 includes exposing a pre-catalyst to a fluid composition containing at least methane to convert an active metal precursor of the pre-catalyst to a carbidic form of the active metal. The exposing may include bringing the pre-catalyst and the fluid composition containing at least methane into physical contact, or immediate or close proximity. The exposing may proceed under heating and/or thermal treatment. For example, in an embodiment, the exposing may proceed at a temperature greater than about 600 K. In many embodiments, the exposing may proceed at a temperature greater than about 900 K. In preferred embodiments, the exposing may proceed at about 973 K.
[0058] The pre-catalyst may include any of the porous supports, active metals, and/or precursors thereof described herein. For example, the pre-catalyst may include an active metal precursor on a surface of the pre-catalyst and/or in pores of the pre-catalyst. The active metal precursor may include any active metal compound. In an embodiment, the active metal precursor may include metal oxides, such as MoO.sub.3, and/or dimers, such as (Mo.sub.2O.sub.5).sup.+2 on a surface and/or pores of the pre-catalyst. Upon the exposing of the pre-catalyst to the fluid composition containing at least methane, the active metal precursor may be carburized to form a carbidic form of the active metal. For example, in embodiments in which the active metal precursor includes one or more of MoO.sub.3 and (Mo.sub.2O.sub.5).sup.+2 on a surface and/or pores of the pre-catalyst, a carbidic form of the active metal may be characterized by the chemical formula MoC.sub.x. The carbidic form of the active metal may be catalytically active. While the discussion above relates to Mo, any of the active metals of the present disclosure may be used herein, such as W, among others.
[0059] The step 202 includes contacting the pre-catalyst with a hydrogen-accepting component to form a catalyst. The pre-catalyst and hydrogen-accepting component may be contacted in a manner sufficient to form a catalyst in any of a variety of configurations (e.g., staged and/or stratified catalyst beds, interparticle mixtures, intraparticle mixtures, etc.). For example, the contacting may include bringing the pre-catalyst and the hydrogen-accepting component into physical contact, or immediate or close proximity. In an embodiment, the contacting may include mixing to form one or more of interpellet physical mixtures, interparticle physical mixtures, and intraparticle physical mixtures of the pre-catalyst and hydrogen-accepting component. In an embodiment, the contacting may include arranging the pre-catalyst and hydrogen-accepting component to provide a catalyst with alternating stratified layers of the pre-catalyst and hydrogen-accepting component. In an embodiment, the contacting may include encapsulating the pre-catalyst with the hydrogen-accepting component and/or encapsulating the hydrogen-accepting component with the pre-catalyst. These shall not be limiting as the contacting may proceed to form catalysts in other configurations.
[0060] The pre-catalyst may include any of the pre-catalysts of the present disclosure. For example, in many embodiments, the pre-catalyst may include a carbidic form of the active metal on a surface and/or pores of a porous support. For example, the pre-catalyst may include MoC.sub.x on a surface and pores of a porous support, such as H-ZSM-5, where x is at least 0.01. In other embodiments, the pre-catalyst may include an active metal precursor on a surface of the pre-catalyst and/or in pores of the pre-catalyst. The pre-catalyst may be characterized by a particle size ranging from about 180 m to about 425 m. In certain embodiments, the pre-carburized catalyst has an average particle size in the range of about 180 m to about 425 m. The hydrogen-accepting component may include any of the hydrogen-accepting components described herein. In certain embodiments, the hydrogen-accepting component has an average particle size in the range of about 180 m to about 425 m. In many embodiments, the hydrogen-accepting component includes zirconium. In preferred embodiments, the hydrogen-accepting component includes zirconium particles with a particle size ranging from about 180 m to about 425 m.
[0061] In an embodiment, the method may comprise exposing a Mo/H-ZSM-5 pre-catalyst to a fluid composition containing methane and argon to convert molybdenum oxides of the Mo/H-ZSM-5 pre-catalyst to a carbidic form of molybdenum thereby forming MoC.sub.x/H-ZSM-5; and mixing the MoC.sub.x/H-ZSM-5 with zirconium particles to form a catalyst, such as an interpellet physical mixture of Zr and MoC.sub.x/H-ZSM-5.
[0062] In an embodiment, the method of preparing a catalyst and/or pre-catalyst may comprise one or more of the following steps: contacting a first active metal precursor and a porous support sufficient to deposit the first active metal precursor on a surface of the porous support and form a first pre-catalyst; treating the first pre-catalyst sufficient to convert the first active metal precursor to a second active metal precursor on a surface and/or pores of the porous support and form a second pre-catalyst; contacting the second pre-catalyst with a fluid composition containing at least methane sufficient to convert the second active metal precursor to a carbidic form and form a third pre-catalyst, and contacting the third pre-catalyst with a hydrogen accepting component to form a catalyst.
[0063] In an embodiment, the method of preparing a catalyst and/or pre-catalyst may comprise one or more of the following steps: heating a zeolite precursor to form a proton form of a zeolite; contacting (grinding and heating) the proton form of the zeolite with a Mo precursor and/or W precursor sufficient for the Mo precursor and/or W precursor to disperse on a surface of the proton form of the zeolite and form a first intermediate; heating the first intermediate sufficient for the Mo precursor and/or W precursor to migrate into pores of the proton form of the zeolite and form a second intermediate (e.g., Mo/H-ZSM-5); optionally reducing a particle size of the second intermediate; exposing the second intermediate to at least methane sufficient for carburization of the Mo precursor and/or W precursor thereby forming a third intermediate; and contacting the third intermediate with a metal to form a catalyst.
[0064]
[0065] The step 301 includes contacting a catalyst with a fluid composition containing at least methane to form hydrocarbons. The contacting may include bringing the catalyst and the fluid composition into physical contact, or immediate or close proximity. Examples of the contacting may include, but are not limited to, one or more of feeding, flowing, passing, and pumping. The contacting may proceed under conditions suitable for dehydroaromatization of methane and/or non-oxidative methane aromatization to yield hydrocarbons, such as aromatics. For example, the contacting may proceed at or to temperatures ranging from about 500 K to about 1500 K. In many embodiments, the contacting may proceed at or to temperatures ranging from about 900 K to about 1200 K. In preferred embodiments, the contacting may proceed at or to temperatures ranging from about 973 K to about 1193 K. In other embodiments, the contacting may proceed at or to temperatures of less than about 500 K and/or greater than about 1500 K. The contacting may proceed at or to a pressure ranging from about greater than 0 kPa to about 150 kPA. In many embodiments, the contacting may proceed at or to a pressure ranging from about 3 kPa to about 100 kPA, such as about 3.28 kPa to about 95.13 kPa. In some embodiments, the contacting proceeds at or to about atmospheric pressure. In other embodiments, the contacting may proceed at or to a pressure greater than about 150 kPa. In certain embodiments, prior to the contacting, the method includes exposing a pre-catalyst comprising an active metal supported on a proton form of a zeolite to methane to obtain the pre-carburized catalyst.
[0066] The catalyst may include any of the catalysts and/or pre-catalysts described herein. For example, in an embodiment, the catalyst may include a pre-catalyst, such as the Mo/H-ZSM-5 pre-catalyst. In an embodiment, the catalyst may include a catalyst including Zr+MoC.sub.x/H-ZSM-5. In an embodiment, the catalyst includes a zirconium mixed with a porous support, wherein the porous support includes a proton form of a zeolite and a carbidic form of Mo or W deposited on a surface and pores of the proton-form zeolite. The fluid composition may include at least methane. In some embodiments, the fluid composition may include one or more other chemical species. For example, in an embodiment, the fluid composition may include natural gas feedstock.
[0067] The contacting may form hydrocarbons, such as aromatics. For example, in many embodiments, the major products may include one or more of benzene, toluene, naphthalene (e.g., with greater than about 92% carbon selectivity), with the minor products including xylenes and C.sub.10.sup.+ (e.g., with less than about 6% carbon selectivity). C.sub.2H.sub.x products, such as ethane and ethylene, may also be produced (e.g., accounting for less than about 3% of total products on a carbon basis). In certain embodiments, the hydrocarbons, or aromatics, are formed without an induction period.
[0068] While not wishing to be bound to a theory, it is believed that the CH bonds in methane may be activated on MoC.sub.x clusters formed during carburization, which may also catalyze dehydrogenation to form C.sub.2H.sub.x species which subsequently undergo oligomerization/hydrogen transfer on residual Bronsted acid sites within zeolite channels to form aromatics. In addition, in-situ removal of hydrogen may be achieved using the hydrogen-acceptor component, such as zirconium, resulting in the formation of hydride, for example, ZrH.sub.x species. The in-situ removal of hydrogen overcomes thermodynamic equilibrium constraints to enhance methane conversion and aromatic product formation. In certain embodiments, the single-pass methane conversion can be greater than the equilibrium methane conversion under the same reaction conditions. For example, single-pass methane conversion may be as high as 27%, as compared to 10% equilibrium conversion.
[0069] The step 302 includes regenerating the catalyst. The regenerating may include desorption of species deposited on the catalyst during methane dehydroaromatization. In many embodiments, the catalyst may be regenerated by contacting the catalyst with an inert flow under thermal treatment. For example, the catalyst may be contacted with a helium flow at temperatures ranging from about 900 K to about 1200 K, or preferably about 1193 K. The regenerating may release hydrogen absorbed by the hydrogen-acceptor during methane dehydroaromatization and that resulted in the formation of hydrides, thereby regenerating the hydrogen-acceptor component. In addition or in the alternative, the released hydrogen may hydrogenolyze carbonaceous species deposited during methane dehydroaromatization.
[0070] The following Examples are intended to illustrate the above invention and should not be construed as to narrow its scope. One skilled in the art will readily recognize that the Examiners suggest many other ways in which the invention could be practiced. It should be understand that numerous variations and modifications may be made while remaining within the scope of the invention.
Example 1
[0071] The following Example relates to the absorptive hydrogen scavenging for enhanced aromatics yield during non-oxidative methane aromatization on Mo/H-ZSM-5 catalysts. More specifically, this Example describes an increase in yields of desired products in non-oxidative methane dehydrogenation on Mo/HZSM-5 formulations by introducing an additional function that scavenges byproduct hydrogen thereby lifting thermodynamic constraints on attainable yields and increasing rates of products formation. The increase in aromatics yield by addition of a hydrogen-adsorbent was rationalized. It was surmised that an optimum in efficacy of hydrogen removal existed because excess hydrogen removal should result in multi-ring aromatic and coke formation as the reversibilities for these reactions depend more sensitively on H.sub.2 concentration than the reversibilities for ethylene and benzene. The inquiry focused on, among other things, the effects of proximity between Mo/ZSM-5 and the hydrogen adsorptive function and what material characteristics of the additive determined its effectiveness as a hydrogen adsorbent.
[0072] It was proposed to examine the effects of proximity between Mo/ZSM-5 domains and Zr domains on rates and selectivity in polyfunctional CH.sub.4 dehydroaromatization. Molecular transport events conveyed hydrogen byproduct formed at Mo/ZSM-5 domains to Zr domains which stoichiometrically scavenge hydrogen forming a metal (oxy)hydride. The consequences of the kinetic relevance of this molecular transport event on rates and selectivity was assessed by studying mixtures of Mo/ZSM-5 and Zr in three configurations spanning seven decades of length scales: alternating stratified layers of Mo/ZSM-5 and Zr (10.sup.2 m) (
[0073] Inert quartz wool was utilized to separate stratum of Mo/ZSM-5 from stratum of Zr-adsorbents, and homogeneous interparticle and intraparticle physical mixtures were prepared. Zr- and Mo-exchanged ZSM-5 formulations were prepared following protocols described herein. This portfolio of materials, systematically variate in proximity between Mo/ZSM-5 and Zr, was subjected to a suite of structural and chemical characterization protocols to determine bulk structure and morphology (X-ray diffraction and electron microscopy), chemical composition and speciation, (ICP-OES, X-ray absorption and photoelectron spectroscopy), and active site identification, enumeration, and accessibility (transient sorption measurements and temperature programmed surface reactions) of both as-prepared and post-reaction catalysts.
[0074] It was expected that the kinetic relevance of hydrogen transport events from the Mo/ZSM-5 domain to the Zr domain would increase with decreasing proximity, and consequently, it was expected that the efficacy of hydrogen removal would decrease with decreasing proximity. It was posited that the selectivity to desired products (e.g., ethylene and benzene) was maximized at an intermediate value of hydrogen removal efficacy. It was anticipated that too great of hydrogen removal efficacy would increase net rates of highly unsaturated species more strongly than ethylene and benzene because the hydrogen concentration dependencies of net rates increased with increasing product unsaturation. Tuning length scales to tune kinetic relevance of hydrogen transport provided one handle for optimizing desired selectivity in polyfunctional catalysis for CH.sub.4 dehydroaromatization. The identity, structure, and composition of the hydrogen removal function was another handle.
[0075] Chemical transient studies and Raman spectroscopy evinced formation of (Mo.sub.2O.sub.5).sup.2+ dimers on air treatment of MoO.sub.3/H-ZSM-5 physical mixtures at 973 K and removal of 2.440.1 O:Mo during initial CH.sub.4 reactions with MoO.sub.x precursors in a stoichiometric reaction that led to MoC.sub.x moieties. The resulting MoC.sub.x/H-ZSM-5 formulation exhibited a steady-state forward benzene synthesis rate of (5.050.09)10.sup.4 mol mol.sub.Mo.sup.1 s.sup.1 during methane dehydroaromatization at 973 K. Addition of Zr metal particles to MoC.sub.x/ZSM-5 in interpellet mixtures resulted in a 2.7 increase in methane converted and a concurrent 1.4, 2.1, 2.6, and 5.6-fold increase in benzene, naphthalene, toluene, and C.sub.10+ yields, respectively, in reference to a conventional MoC.sub.x/ZSM-5 catalyst at equivalent time-on-stream (8.7 ks). The maximum methane conversion on these interpellet mixtures exceeded 27% demonstrating that equilibrium limitations encountered in methane dehydroaromatization on MoC.sub.x/ZSM-5 were circumvented by addition of Zr as a hydrogen absorbent. Subsequent thermal treatment of the catalyst in helium flow at 973 K resulted in desorption of absorbed hydrogen and in regeneration of the Zr absorbent leading to partial regeneration of the polyfunctional catalyst formulation yielding above equilibrium methane conversions. Hydrogen uptake experiments on Zr metal at 973 K demonstrated that zirconium metal forms ZrH.sub.1.75 on hydrogen exposure (3.28-95.13 kPa) as reconciled by X-ray diffraction.
Materials and Methods
[0076] Catalyst Synthesis.
[0077] NH.sub.4-ZSM-5 (Zeolyst International, Si/Al=11.5, CBV 2314) was converted to H-ZSM-5 by treating in dry air (1.67 cm.sup.3 s.sup.1) to thermally decompose NH.sub.4.sup.+ to H.sup.+ and NH.sub.3 (g) by increasing the temperature from room temperature to 773 K at 0.0165 K s.sup.1 and holding at 773 K for 36 h. Intimate mixtures containing Mo: Al.sub.f0.25 were prepared by grinding together MoO.sub.3 (Sigma-Aldrich, 99.9%) and H-ZSM-5 powders in an agate mortar and pestle for 0.25 h. The mixture was heated from room temperature to 623 K at 0.0167 K s.sup.1 and held at this temperature for 15 h in dry air (0.67 cm.sup.3 s.sup.1) resulting in water removal and dispersion of MoO.sub.3 on the external surface of the zeolite. Finally, the mixture was heated to 973 K at 0.167 K s.sup.1 and held at this temperature for 10 h to facilitate molybdenum oxide migration into the zeolite pores. The catalyst prepared via the above-mentioned protocol is referred to as Mo/H-ZSM-5 in this Example. The Mo/H-ZSM-5 powder was pressed to form pellets that were then crushed and sieved to obtain particle sizes between 180 and 425 m (mesh 40-80) for use in catalytic reactions. Pure zirconium (Zr) metal was obtained from American Elements (PN # ZR-M-0251M-GR.1T2MM, 99.5+% purity, metal basis) as granule-shaped particles (1-2 mm granule size) which were crushed and sieved to obtain particle sizes between 180 and 425 m (mesh 40-80) for use in the catalytic reactions.
[0078] Catalyst Characterization.
[0079] The bulk structure of the zirconium (American Elements PN # ZR-M-0251M-GR.1T2MM, 99.5+% purity, metal basis) and zirconium hydride particles (prepared as discussed in Section 2.4) was determined by X-ray diffraction (XRD) using a Bruker D8 Discover 2D X-ray diffractometer with a 2-D VNTEC-500 detector, Co K.sub. X-ray radiation with a graphite monochromator, and a 0.8 mm point collimator. Scans were measured in three measurement frames at 20 (2), 45 (2), and 70 (2) at 900 s frame.sup.1 while rotating the sample. Area detector images were finally converted to one-dimensional intensity vs. 2 data sets by using an averaging integration algorithm and the radiation wavelength was recalibrated for Cu K.sub. wavelength (=1.541 ).
[0080] Raman spectra of three samplesa physical mixture of MoO.sub.3 and H-ZSM-5 with Mo:Al.sub.f0.25, Mo/H-ZSM-5 with Mo:Al.sub.f0.25, and H-ZSM-5 air treated at 973 K for 5 hwere collected on a WITec alpha 300R confocal Raman microscope (WITec Instrument Corp., Germany) equipped with a frequency doubled 532 nm Nd:YAG laser, a 100 Nikon air objective with a numeric aperture of 0.90, a UTS300 spectrometer with 1800 groves/mm grating, and a DV401 CCD detector. The lateral spatial resolution was 0.5 m and the data were processed using the WITec Project 4.05 software.
[0081] Brnsted acid site density of the H-ZSM-5 sample employed was measured using NH.sub.3 temperature-programmed-desorption. H-ZSM-5 (0.160 g) was treated in flowing NH.sub.3 (1.67 cm.sup.3 s.sup.1, 1.01% NH.sub.3 in He, Praxair, certified standard) at 423 K for 1.5 ks, purged in flowing He (Minneapolis Oxygen, 99.997%; 1.67 cm.sup.3 s.sup.1) for 28.8 ks, and ramped to 823 K at 0.167 K s.sup.1 in flowing He and Ar (Matheson, UHP/Zero; used as an internal standard) while continuously monitoring the effluent via an online mass spectrometer (MKS Cirrus, m/z=16, 17, 18, 40) to quantify desorbed NH.sub.3. The H.sup.+ site density was calculated by assuming unit stoichiometry between H.sup.+ and NH.sub.3 desorbed from 423 to 823 K.
[0082] Methane Dehydroaromatization Catalytic Reactions.
[0083] Methane dehydroaromatization reactions were performed in a fixed bed tubular quartz reactor (I.D. 10 mm) with an outer thermowell above the porous quartz frit (for holding the catalyst bed stationary) to hold a thermocouple to monitor the reaction temperature. The catalyst sample was heated in helium flow (0.33 cm.sup.3 s.sup.1, UHP, Minneapolis Oxygen) from room temperature to the reaction temperature, 973 K, at 0.18 K s.sup.1 in a resistively heated furnace (National Element FA120). Axial temperature gradients in the reactor were minimized by using an annular inconel cylinder to fill the vacant space in the furnace. A feed gas mixture of CH.sub.4/Ar (90 vol % CH.sub.4 and 10 vol % Ar, total feed flow rate 0.27 cm.sup.3 s.sup.1, UHP, Matheson Tri-Gas) was introduced to the reactor to perform methane reactions at 973 K and atmospheric pressure, with Ar serving as an internal standard. All flow lines were heated to temperatures in excess of 473 K via resistive heating to prevent condensation of effluents. The composition of the reactor effluent was analyzed using a mass spectrometer (MKS Cirrus 200 Quadrupole MS system) and a gas chromatograph (Agilent 7890) equipped with a methyl-siloxane capillary column (HP-1, 50 m320 m0.52 m) connected to a flame ionization detector (FID) for detection of hydrocarbons and a GS-GasPro column (60 m0.320 mm) connected to a thermal conductivity detector (TCD) for detection of permanent gases (H.sub.2, Ar, and CH.sub.4). Transient product evolution throughout the course of the reaction was measured with an online mass spectrometer (MS) (MKS Cirrus 200 Quadrupole MS system) connected to the outlet of the GC. The number of removed O atoms during carburization of molybdenum oxide was determined by the cumulative amount of CO and CO.sub.2 eluted as calculated from the transient MS signal with Ar as an internal standard.
[0084] The samples designated as MoC.sub.x/H-ZSM-5 in this Example were prepared by exposing the Mo/H-ZSM-5 formulation to a CH.sub.4/Ar mixture (90 vol % CH.sub.4 and 10 vol % Ar, total feed flow rate 0.27 cm.sup.3 s.sup.1) for 15.5 ks (for catalyst loading 1.2 g) leading to complete carburization of the molybdenum oxide and subsequently, switching the feed to helium (0.83 cm.sup.3 s.sup.1). The sample was flushed in helium flow (0.83 cm.sup.3 s.sup.1) at 973 K for 0.9 ks and then cooled to room temperature in the same helium flow. For the formulation designated as Zr+MoC.sub.x/H-ZSM-5 in this Example, the MoC.sub.x/H-ZSM-5 sample (1.2 g) was unloaded from the reactor and mixed with Zr particles (180-425 m, 2.4 g) in a glove bag (Sigma Aldrich, Z530212, AtmosBag, two-hand, non-sterile, size M, Zipper-lock closure type) under inert (helium) atmosphere to prevent oxidation of the carburized catalyst. This procedure gave interpellet physical mixtures of Zr and MoC.sub.x/H-ZSM-5. The interpellet mixture was subsequently loaded in the same quartz reactor under an inert atmosphere and transferred to the furnace without exposing the mixture to ambient conditions. The mixture was heated to the reaction temperature, 973 K, in helium flow (0.33 cm.sup.3 s.sup.1) at 0.18 K s.sup.1 before performing methane dehydroaromatization reactions.
[0085] Experiments to demonstrate in-situ regeneration of the absorbent function involved running methane dehydroaromatization on Zr+MoC.sub.x/H-ZSM-5 formulations for 3.6 ks. Subsequently, the flow was switched to helium (0.83 cm.sup.3 s.sup.1) for 61.2-82.8 ks and hydrogen in the reactor effluent was monitored using a mass spectrometer before performing methane dehydroaromatization reactions following the procedure described above.
[0086] Hydrogen Uptake on Zirconium.
[0087] The hydrogen uptake capacity of Zr was examined by loading 2.4 g of Zr particles (180-425 m) in a tubular quartz reactor (I.D. 10 mm) and heating the sample in helium flow (0.83 cm.sup.3 s.sup.1) to 973 K from ambient temperature at 0.18 K s.sup.1. Subsequently, the flow was switched to a H.sub.2/Ar gas mixture (3.28-95.13 kPa H.sub.2/balance Ar, total flow rate 1.7 cm.sup.3 s.sup.1) and the reactor effluent transient was monitored via a mass spectrometer. Once the hydrogen (m/z=2) mass spectrometer signal reached a steady value in 0.4-12 ks depending on the feed hydrogen pressure, the feed was switched to helium flow (0.83 cm.sup.3 s.sup.1).
[0088] Temperature-Programmed Desorption (TPD).
[0089] Temperature-programmed desorption in helium flow was performed either post methane dehydroaromatization on Zr+MoC.sub.x/H-ZSM-5 or after hydrogen uptake on Zr particles. The feed was switched from a CH.sub.4/Ar or a H.sub.2/Ar mixture to helium (0.83 cm.sup.3 s.sup.1) at 973 K. The sample was flushed in helium flow at 973 K for 0.9 ks. Subsequently, the sample was heated to 1193 K at 0.061 K s.sup.1 and held at 1193 K for 25-35 ks in helium flow. The reactor effluent was monitored using a mass spectrometer to quantify hydrogen and methane removed from the catalyst sample during TPD with helium as an internal standard. Finally, the catalyst (Zr+MoC.sub.x/H-ZSM-5) was cooled to room temperature whereas Zr particles were cooled to 973 K. Subsequent hydrogen uptake/TPD cycles at different hydrogen pressures (3.28-95.13 kPa H.sub.2/balance Ar, total flow rate 1.7 cm.sup.3 s.sup.1) were performed on a single Zr sample (2.4 g loading) as discussed in Section 3.4.
Results and Discussion
[0090] The addition of Zr to pre-carburized Mo/H-ZSM-5 formulations (denoted as MoC.sub.x/H-ZSM-5 in this Example) resulted in formation of ZrH.sub.x species during methane dehydroaromatization reactions at 973 K resulting in 27% maximum single-pass methane conversion and an increase in yield of methane-derived hydrocarbon products at equivalent time-on-stream because equilibrium limitations were transiently circumvented by the hydrogen-absorptive Zr function. Thermal treatment of the Zr+MoC.sub.x/H-ZSM-5 formulation in helium flow resulted in regeneration of the Zr-absorbent leading to above-equilibrium methane conversions in successive reaction-regeneration cycles.
[0091] Chemical and Structural Characterization of Mo/H-ZSM-5 Formulations.
[0092] The transient evolution of products during the induction period of methane dehydroaromatization (DHA) on Mo/H-ZSM-5 catalyst at 973 K is shown in
[0093]
[0094] Formation of (Mo.sub.2O.sub.5).sup.2+ dimers upon ion exchange of MoO.sub.3 with two Al proximate centers within H-ZSM-5 channels was evidenced using X-ray absorption spectroscopy and supported by computational chemistry studies. Mo/H-ZSM-5 showed a pre-edge feature in near-edge X-ray absorption spectra similar to MgMo.sub.2O.sub.7 which contained ditetrahedral Mo centers and also a post-edge energy (20 keV) absorbance similar to MgMo.sub.2O.sub.7 as opposed to MoO.sub.3 which contained distorted octahedral Mo centers. In order for (Mo.sub.2O.sub.5).sup.2+ moieties to stabilize, the framework oxygen connected to Al atoms should reside within 0.42-0.55 nm of each other. Computational studies in other reports show that the fraction of Al atoms that reside within 0.55 nm of another framework Al for a Si:Al.sub.f15 is 0.42 assuming random Al occupancy of T sites in the sample. It has also been demonstrated that at Mo loadings in excess of 2 wt % for H-ZSM-5 with Si:Al13, irreversible damage to the zeolite framework occurred due to MoO.sub.x reaction with framework Al leading to aluminum molybdate formation which was inactive for methane DHA. An upper limit to Mo loading in H-ZSM-5 catalysts for methane DHA applications was attributed to the unavailability of enough Al pairs to accommodate all Mo dimers which led to the need for migration of Al atoms during thermal treatment leading to dealumination. Consequently, a Mo:Al.sub.f0.25 for a H-ZSM-5 catalyst with Si:Al11.5 was used to ensure the required number of Al atom pairs were available for (Mo.sub.2O.sub.5).sup.2+ dimer formation and stabilization within zeolite channels. The concentration of free Brnsted acid sites in H-ZSM-5 (Si/Al=11.5) was determined via NH.sub.3 uptake to be 1.2110.sup.3 mol g.sub.cat.sup.1, which was similar to the concentration of Al in the zeolite as calculated from the Si-to-Al ratio (1.3310.sup.3 mol g.sub.cat.sup.1).
[0095] CH.sub.4 Dehydroaromatization on Mo/H-ZSM-5 and MoC.sub.x/H-ZSM-5 Catalysts.
[0096]
[0097] Catalyst deactivation with time-on-stream was attributed to the continuous buildup of coke within the zeolite channels which presumably led to reduction of effective zeolite diameter, thus inhibiting the formation of bulky aromatics like naphthalene. DME titrations, XPS, ion-scattering spectroscopy, and FT-IR measurements by others demonstrated a reduction in the number of acid sites as well as formation of carbonaceous deposits in zeolite channels at long times-on-stream during methane DHA on Mo/H-ZSM-5. It was observed that benzene and naphthalene selectivity to be 67% and 21% respectively at 15 ks time-on-stream which were typical for methane DHA reactions on Mo/H-ZSM-5. Aromatic selectivity (naphthalene and benzene) decreased at longer times-on-stream with a concurrent increase in C.sub.2H.sub.x selectivity, consistent with deactivation of acidic sites in zeolites by coke deposition which limited access to these oligomerization/dehydroaromatization sites.
[0098] The induction (carburization) period and the initial increase and subsequent reduction in aromatic formation rates with time-on-stream was suggestive of a bifunctional catalytic mechanism wherein CH bonds in CH.sub.4 were activated on MoC.sub.x clusters formed during carburization which also catalyzed dehydrogenation to form C.sub.2H.sub.x species which subsequently underwent oligomerization/hydrogen transfer on residual Brensted acid sites within zeolite channels to form aromatics (benzene and naphthalene). Others have demonstrated that bulk molybdenum carbide catalyzed methane conversion to C.sub.2 products but did not yield any aromatics. Furthermore, others also showed that unsupported Mo.sub.2C catalyzed ethane dehydrogenation to ethylene but not aromatics whereas MoC.sub.x/H-ZSM-5 catalyzed ethane conversion to benzene. Formation of MoC.sub.x species in zeolite micropores and their subsequent function as catalytic centers for methane DHA has been established through numerous characterization studies including Raman spectroscopy, X-ray absorption, X-ray photoelectron spectroscopy, ion sputtering spectroscopy, UV-visible near infrared spectroscopy, and .sup.95Mo nuclear magnetic resonance (NMR). One report demonstrated, using DME chemical titrations during methane DHA on Mo/H-ZSM-5 at 950 K, that the concentration of free Brnsted acid sites increased with carburization time due to MoC.sub.x cluster formation from (Mo.sub.2O.sub.5).sup.2+ species which had exchanged with the Brnsted acid sites during Mo/H-ZSM-5 preparation. Subsequently, the number of DME accessible protons decreased with time-on-stream which was attributed to adsorption of aromatics formed on regenerated acid sites, as suggested by a concurrent increase in benzene formation rate. These characterization studies were in line with the bifunctional mechanism for methane DHA on Mo/H-ZSM-5 catalyst at 973 K.
[0099] The exact stoichiometry and coordination of MoC.sub.x clusters are still debated in the literature, with emphasis on the role of Mo speciation on methane DHA reactions, but the reduced carbidic nature of active Mo-centers and their proficiency during methane DHA reactions is well established. All rates reported in this Example were normalized to the total number of Mo atoms to account for all available Mo atoms resulting in the lowest possible rate.
[0100] The forward rate of benzene formation for the stoichiometric reaction of CH.sub.4 to H.sub.2 and C.sub.6H.sub.6 (equation 1) on Mo/H-ZSM-5 at 973 K was calculated by analyzing the net rate of benzene formation and effluent hydrogen, benzene, and methane pressures in the regime (12-22 ks time-on-stream) in which net benzene formation rate was invariant.
CH.sub.4.Math.C.sub.6H.sub.6+ 3/2H.sub.2(1)
[0101] The approach to equilibrium for the reaction in equation 1, , was calculated via equation 2 using outlet pressures of benzene, hydrogen, and methane and the equilibrium constant determined from thermodynamic values at 973 K (K.sub.eq=0.0302). The forward rate of benzene formation (R.sub.for) was related to the net benzene rate (R.sub.net) and as shown in equation 3. Using the calculated values of R.sub.net (2.710.sup.4 mol s.sup.1 mol.sub.Mo.sup.1) and (0.5), the forward rate of benzene formation (R.sub.for) was (5.050.09)10.sup.4 mol s.sup.1 mol.sub.Mo.sup.1. This forward rate of benzene formation was similar to rates reported by others (4.70.810.sup.4 mol s.sup.1 mol.sub.Mo.sup.1) and (5.0310.sup.4 mol s.sup.1 mol.sub.Mo.sup.1) for methane DHA on Mo/H-ZSM-5 at 950 K.
[0102] Another report showed that the net rate of benzene formation systematically decreased with increasing H.sub.2 co-feed concentration (H.sub.2:CH.sub.4=0.09-0.27 molar ratio) and that benzene synthesis rates recover to their pre-H.sub.2 co-feed values upon removal of H.sub.2 from the reactor influent demonstrating that hydrogen did not cause any irreversible structural or chemical modification to the MoC.sub.x moieties present in the catalyst during methane DHA at 950 K. Moreover, an invariant forward rate of benzene formation ((3.80.5)10.sup.4 mol s.sup.1 mol.sub.Mo.sup.1) was obtained for different catalyst loadings (0.1-1.0 g) and H.sub.2:CH.sub.4 co-feeds (0.03-0.11 molar ratio) at 950 K demonstrating that hydrogen did not have any kinetic effect on the rate-limiting step of CH.sub.4 DHA and that hydrogen reduced the aromatic formation rate by increasing the thermodynamic reversibility of methane to benzene reaction. These results showed that methane DHA on Mo/H-ZSM-5 was limited by thermodynamic equilibrium due to inhibition by abundant amounts of hydrogen produced in the catalyst bed.
[0103] CH.sub.4 reaction on Mo/H-ZSM-5 was performed at 973 K for 15.5 ks, during which (Mo.sub.2O.sub.5).sup.2+ dimers were carburized to MoC.sub.x species with O.sub.removed:Mo2.440.1 as CO and CO.sub.2 as noted in Table 1, and then switched the reactor influent to an inert (helium, 0.83 cm.sup.3 s.sup.1). Methane conversion and product formation rates exhibited a time-on-stream profile similar to Mo/H-ZSM-5 (
[0104] CH.sub.4 Dehydroaromatization on Physical Mixtures of Zr and MoCJH-ZSM-5.
[0105] Methane dehydroaromatization was investigated at 973 K on an interpellet physical mixture of zirconium metal and MoC.sub.x/H-ZSM-5. The physical mixing of Zr particles with the MoC.sub.x/ZSM-5 formulation was performed in an inert environment to avoid oxidation of MoC.sub.x species on atmospheric exposure due to their oxophilic nature. Methane flow over Zr particles at 973 K did not result in any product (aromatic or C.sub.2H.sub.x) formation as shown in
[0106] Benzene and naphthalene instantaneous selectivity calculated on a carbon basis were 60% and 30% at 0.3 ks whereas toluene and C.sub.2H.sub.x accounted for 3% and 2%, respectively, of the total carbon in the products observed. The identity and sequence of appearance of products remain unchanged for MoC.sub.x/H-ZSM-5 and Zr+MoC.sub.x/H-ZSM-5 suggesting that the bifunctional reaction pathways of methane dehydroaromatization were unperturbed upon Zr addition. The observed enhancement in methane conversion and aromatic product rates can be explained as a consequence of in-situ hydrogen removal resulting in alleviation of thermodynamic equilibrium constraints and was evident when comparing the net formation rate of products and product yields during methane DHA on MoC.sub.x/H-ZSM-5 and Zr+MoC.sub.x/H-ZSM-5 at 973 K.
[0107]
[0108] Cumulative product yield for methane DHA on MoC.sub.x/H-ZSM-5 and Zr+MoC.sub.x/H-ZSM-5 at time-on-stream, t, was defined as the total number of moles of product formed per total Mo atom in the catalyst at the end of time, t, and was calculated from the area under the curve of the plot for net rate as a function time-on-stream shown in
C.sub.product=(2C.sub.2H.sub.x+6C.sub.6H.sub.6+7C.sub.7H.sub.8+8C.sub.8H.sub.10+10C.sub.10H.sub.8+11C.sub.10.sup.+)(4)
[0109] Hydrogen Uptake and Structural Characterization of Zirconium Metal (Zr) Hydrogen Uptake of Zirconium Metal.
[0110] Hydrogen uptake experiments were performed on zirconium metal particles to determine the hydrogen absorption capacity of pure Zr metal at 973 K.
[0111] As noted in Table 2, the hydrogen breakthrough times decreased with increasing hydrogen pressures and the uptake curves at hydrogen pressures exceeding 10 kPa resembled heavyside functions as shown in
[0112] Structural Characterization of Zirconium and Zirconium Hydride.
[0113] Following the hydrogen uptake of zirconium metal particles, temperature-programmed-desorption (TPD) in helium flow was performed to investigate the stoichiometry of zirconium hydride formed and the proficiency of zirconium metal for regeneration after hydrogen absorption. Hydrogen was observed to evolve from zirconium hydride particles as the temperature of the sample was increased from 973 K to 1193 K at 0.061 K s.sup.1 and the corresponding time-on-stream evolution for different hydrogen uptake pressures (3.28-95.13 kPa), monitored using an online mass spectrometer, is shown in
[0114] X-ray diffraction patterns of zirconium metal prior to hydrogen uptake and after hydrogen uptake at 33.34 kPa hydrogen pressure are shown in
[0115] Regeneration of CH.sub.4 Dehydroaromatization CatalystsRegeneration of Physical Mixture of Zr and MoC.sub.x/H-ZSM-5 for CH.sub.4 DHA.
[0116] Regeneration of the polyfunctional Zr+MoC.sub.x/H-ZSM-5 catalyst formulation was investigated by treating the catalyst in helium flow at 973 K after performing methane DHA for 3.6 ks. During the helium flush at 973 K, hydrogen was observed to elute from the catalyst, presumably from the hydrogen absorbed by zirconium during reaction, as monitored using an online mass spectrometer (shown in
H.sub.missing=(2H.sub.2+4C.sub.2H.sub.x+6C.sub.6H.sub.6+8C.sub.7H.sub.8+10C.sub.8H.sub.10+8C.sub.10H.sub.8+10C.sub.10.sup.+)(4CH.sub.4 reacted)(5)
[0117] Temperature-Programmed-Desorption after CH.sub.4 Dehydroaromatization on Zr+MoC.sub.x/H-ZSM-5.
[0118] A temperature-programmed-desorption (TPD) in helium flow (0.83 cm.sup.3 s.sup.1) at 1193 K was performed following methane dehydroaromatization reaction at 973 K on an interpellet physical mixture of Zr and MoC.sub.x/H-ZSM-5 for 9 ks (shown in
[0119] The total number of hydrogen atoms absent in the reactor effluent at the end of 8.7 ks during methane reaction on Zr+MoC.sub.x/H-ZSM-5 was estimated using equation 5 and normalized by total Zr atoms present in the catalyst bed. The post-reaction TPD yielded H:Zr1.60 which was within 10% of the hydrogen missing from hydrogen balance (H.sub.missing:Zr1.48 as noted in Table 1) suggesting that all the hydrogen absorbed by zirconium particles during methane DHA could be removed by high temperature desorption in line with hydrogen uptake experiments discussed in Section 3.4. Similar calculations for methane removed during TPD resulted in C:Mo1.69 which was less than the carbon deposited calculated from carbon balance (C:Mo3.46). This difference can be attributed to the insufficient availability of hydrogen to hydrogenolyze all the carbon deposited.
[0120] In summary, methane reactions on Mo/H-ZSM-5 catalyst at 973 K resulted in an initial carburization period during which O:Mo2.440.10 was removed as CO and CO.sub.2 due to the formation of (Mo.sub.2O.sub.5).sup.2+ dimers on thermal treatment of MoO.sub.3 and H-ZSM-5 physical mixtures at 973 K as reconciled by Raman spectroscopy. Steady state methane dehydroaromatization on MoC.sub.x/H-ZSM-5 catalysts at 973 K resulted in a forward rate of benzene synthesis (5.050.09)10.sup.4 mol s.sup.1 mol.sub.Mo.sup.1 with 67% benzene and 21% naphthalene product selectivity. Maximum methane conversion (27%), aromatic synthesis rates, and benzene, naphthalene, toluene, xylene, and C.sub.10.sup.+ product yields (1.4-5.4 times) were enhanced on interpellet physical mixtures of Zr and MoC.sub.x/H-ZSM-5 as compared to MoC.sub.x/H-ZSM-5 due to in-situ absorptive-hydrogen removal by zirconium metal thereby lifting the thermodynamic constraints for methane dehydrogenation. A post-reaction helium temperature-programmed-desorption at 1193 K resulted in absorbed hydrogen removal (H:Zr1.60) from zirconium hydride particles (H:Zr1.48 from hydrogen balance) formed during methane reactions. Hydrogen uptake experiments on zirconium metal particles at 973 K demonstrated that Zr forms ZrH.sub.1.75 across a large hydrogen pressure range (3.28-95.13 kPa) which can be regenerated to Zr metal by high temperature (1193 K) helium TPD as also supported by XRD analysis. A post-reaction thermal treatment in helium at 973 K of the polyfunctional Zr+MoC.sub.x/H-ZSM-5 catalyst formulation led to regeneration of the absorptive-hydrogen removal function resulting in above equilibrium methane conversions (22-15%) in successive reaction-regeneration cycles.
TABLE-US-00001 TABLE 1 Oxygen removed during carburization, methane converted, product yields, and hydrogen balance for MoC.sub.x/H-ZSM-5 and an interpellet physical mixture of MoC.sub.x/H-ZSM-5 and Zr metal. MoC.sub.x/H-ZSM-5 loading ~1.2 g with Mo/A1.sub.f ~0.25, Zr metal loading ~2.4 g, total flow rate ~0.21 cm.sup.3 s.sup.1 (90 vol % CH.sub.4/balance Ar), reaction at ~973 K. Corresponding time-on-stream data shown in FIGS. 5-8. MoC.sub.x/H-ZSM-5 Zr + MoC.sub.x/H-ZSM-5 O.sub.removed:MO 2.44 0.1 2.51 0.1 (mo1.sub.o mo1.sub.Mo.sup.1).sup.a CH.sub.4 converted (mo1.sub.o 13.68 35.85 mo1.sub.Mo.sup.1).sub.b Product Yield (mo1.sub.o mo1.sub.Mo.sup.1).sub.b C.sub.2H.sub.x 0.49 0.70 Benzene 2.07 3.39 Toluene 0.099 0.199 Xylenes 0.014 0.200 Naphthalene 0.284 0.602 C.sub.10.sup.+ 0.028 0.150 H.sub.2.sup.c 9.02 7.57 H.sub.missing:Zr (mo1.sub.H mo1.sub.Zr.sup.1).sup.d NA 1.48 H:Zr (helium TPD) (mo1.sub.H NA 1.60 mo1.sub.zr.sup.1).sup.e .sup.anumber of oxygen atoms removed per Mo atom calculated considering oxygen-containing products (CO and CO.sub.2) eluted from the reactor during the initial 15.5 ks of methane reaction on Mo/H-ZSM-5 as shown in FIG. 2A. .sup.bcumulative methane converted or product yields calculated after 8.7 ks time-on-stream as shown in FIG. 8. .sup.chydrogen (H.sub.2) yield per Mo atom calculated from time-on-stream data shown in FIG. 7(f) .sup.dnumber of hydrogen atoms missing in the reactor effluent per Z.sub.r atom at the end of 8.7 ks time-on-stream, calculated as: H.sub.missing = (2 H.sub.2 + 4 C.sub.2H.sub.x + 6 C.sub.6H.sub.6 + 8 C.sub.7H.sub.8 + 10 C.sub.8H.sub.10 + 8 C.sub.10H.sub.8 + 10 C.sup.+.sub.10) (4 CH4 reacted) where, effluent flow rates are used for all products and CH.sub.4 converted is calculated as (CH.sub.4 in CH.sub.4 out) during reaction .sup.enumber of hydrogen atoms per Zr atom eluted during temperature-programmed-desorption (TPD) in helium flow (~0.83 cm.sup.3 s.sup.1) at ~1193 K, following the methane reaction on interpellet physical mixture of MoC.sub.x/H-ZSM-5 and Zr metal, as shown in FIG. 12A-12C.
TABLE-US-00002 TABLE 2 Experimental conditions and measured hydrogen uptake during hydrogen uptake experiments for Zr metal shown in FIG. 9A-9B. Zr metal loading ~2.4273 g (~0.0267 moles), Zr particle diameter ~3 10.sup.4 m. H.sub.2 flow rate/ H.sub.2 uptake Breakthrough H.sub.2 absorbed/ H.sub.2 desorbed/ H absorbed: H desorbed: (10.sup.6 mol s.sup.1).sup.a pressure (kPa) time (ks).sup.b (.sup.102 moles).sup.c (.sup.102 moles).sup.d Zr.sup.e Zr.sup.f 2.80 3.28 6.4 2.27 2.34 1.71 1.75 8.14 11.64 1.9 2.32 2.38 1.74 1.79 33.34 51.74 0.6 2.42 2.41 1.81 1.81 61.92 95.13 0.3 2.29 2.28 1.72 1.71 .sup.ahydrogen flow rate during hydrogen uptake experiment for Zr metal shown in FIG. 9A, total feed flow rate ~1.7 cm.sup.3 s.sup.1 (balance Ar), temperature ~973 K. .sup.btime delay for hydrogen signal in mass spectrometer with respect to Ar signal (internal standard and tracer) as shown in FIG. 9A. .sup.cnumber of moles of hydrogen absorbed by Zr metal, calculated using the breakthrough time and the transient during hydrogen uptake experiment shown in FIG. 9A. .sup.dnumber of moles of hydrogen eluted from the reactor during temperature-programmed-desorption (TPD) in helium flow (~0.83 cm.sup.3 s.sup.1) at ~1193 K as shown in FIG. 9B, following hydrogen uptake shown in FIG. 9A). .sup.enumber of hydrogen atoms per Zr atom absorbed during hydrogen uptake (FIG. 9A). .sup.fnumber of hydrogen atoms eluted per Zr atom during temperature-programmed-desorption (TPD) in helium flow shown in FIG. 9B.
TABLE-US-00003 TABLE 3 Methane converted for MoC.sub.x/H-ZSM-5 and an interpellet physical mixture of MoC.sub.x/H-ZSM-5 and Zr metal before and after regeneration in He flow at ~973 K. MoC.sub.x/H-ZSM-5 loading ~1.2 g with Mo/A1.sub.f ~0.25, Zr metal loading ~2.4 g, total flow rate ~0.21 cm.sup.3 s.sup.1 (90 vol % CH.sub.4/balance Ar), reaction at ~973 K. Corresponding time- on-stream data shown in FIG. 11A-G. Zr + MoC.sub.x/H-ZSM-5 MoC.sub.x/H- After After After ZSM-5 Fresh Regeneration 1.sup.b Regeneration 2.sup.b Regeneration 3.sup.b CH4 6.14 19.8 13.75 11.2 10.73 converted (mo1 mo1.sub.mo.sup.1).sup.a .sup.acumulative methane converted calculated after 3.6 ks time-on-stream as shown in FIG. 11A-11G. .sup.bregenerations 1, 2, and 3 of Zr + MoC.sub.x/H-ZSM-5 were performed by flushing the catalyst in helium flow (~0.83 cm.sup.3 s.sup.1) at 973 K for 61.2 ks, 84.6 ks, and 34.2 ks, respectively.
TABLE-US-00004 TABLE S1 Methane converted, product yields, hydrogen balance for MoC.sub.x/H-ZSM-5 and interpellet physical mixture of MoC.sub.x/H-ZSM-5 and Zr metal before and after regeneration in He flow at ~973 K. MoC.sub.x/H-ZSM-5 loading ~1.2 g with Mo/A1.sub.f ~0.25, Zr metal loading ~2.4 g, total flow rate ~0.21 cm.sup.3 s.sup.1 (90 vol % CH.sub.4/balance Ar), reaction at ~973 K. Corresponding time-on-stream methane conversion data shown in FIG. 11A-11G. Zr + MoC.sub.x/H- ZSM-5 After After After MoC.sub.x/H-ZSM-5 Fresh Regeneration 1.sup.b Regeneration 2.sup.b Regeneration 3.sup.b CH4 converted 6.14 19.8 13.75 11.2 10.73 (mo1 mo1.sub.mo.sup.1).sup.a Product Yield (mo1 mo1.sub.mo.sup.1)c C.sub.2H.sub.x 0.19 0.214 0.213 0.216 0.210 Benzene 0.864 1.545 1.266 0.887 0.640 Toluene 0.04 0.103 0.105 0.087 0.071 Xylenes 0.005 0.106 0.104 0.067 0.045 Naphthalene 0.126 0.356 0.171 0.091 0.053 C.sub.10.sup.+ 0.013 0.085 0.051 0.017 0.008 H.sub.2.sup.d 9.02 3.5 3.17 2.18 1.64 H.sub.missing:Zr NA 0.83 0.53 0.47 0.49 (mo1.sub.H mo1.sub.Zr.sup.1).sup.e H:Zr (helium NA 0.84.sup.f 0.88.sup.f 0.71.sup.f 0.67.sup.g flush/TPD) (mo1.sub.H mo1.sub.Zr.sup.1) .sup.acumulative methane converted calculated after 3.6 ks time-on-stream as shown in FIG. 11A-11G. .sup.bregeneration 1, 2, and 3 of Zr + MoC.sub.x/H-ZSM-5 were performed by flushing the catalyst in helium flow (~0.83 cm.sup.3 s.sup.1) at 973 K for 61.2 ks, 84.6 ks, and 34.2 ks, respectively. .sup.ccumulative product yields calculated after 3.6 ks time-on-stream. .sup.dhydrogen (H.sub.2) yield per Mo atom calculated from time-on-stream mass spectrometer data. .sup.enumber of hydrogen atoms missing in the reactor effluent per Zr atom at the end of 3.6 ks time-on-stream, calculated as: H.sub.missing = (2 H.sub.2 + 4 C.sub.2H.sub.x + 6 C.sub.6H.sub.6 + 8 C.sub.7H.sub.s + 10 C.sub.sH.sub.10 + 8 C.sub.10H.sub.s + 10 C.sup.+.sub.10) (4 CH.sub.4 reacted) where, effluent flow rates are used for all products and CH.sub.4 converted is calculated as (CH.sub.4in CH.sub.4out) during reaction .sup.fnumber of hydrogen atoms per Zr atom eluted during helium flush in helium flow (~0.83 cm.sup.3 s.sup.1) at ~973 K, following the methane reaction on interpellet physical mixture of MoC.sub.x/H-ZSM-5 and Zr metal. .sup.gnumber of hydrogen atoms per Zr atom eluted during temperature-programmed-desorption (TPD) in helium flow (~0.83 cm3 s1) at ~1193 K, following the methane reaction on interpellet physical mixture of MoC.sub.x/H-ZSM-5 and Zr metal.
Example 2
[0121] The following Example relates to controlling kinetic and diffusive length-scales during absorptive hydrogen removal in methane dehydroaromatization on MoC.sub.x/H-ZSM-5 catalysts.
[0122] Addition of Zr metal absorbent to MoC.sub.x/H-ZSM-5 in the form of staged-bed, stratified-bed, and interpellet physical mixtures effectively scavenged H.sub.2 from catalyst proximity, enhancing maximum single-pass benzene and naphthalene yield during methane dehydroaromatization (DHA) reactions to 14-16% compared to 8% in formulations without zirconium. The coupling of spatially-distinct catalytic and absorptive functions was achieved by dispersive/diffusive transport which conveyed H.sub.2 to staged Zr both co- and counter-current to bulk advection, thereby suppressing axial H.sub.2 partial pressure profiles along the catalyst bed and enhancing net aromatization rates. The significance of dispersive hydrogen transport during methane DHA was shown for the first time by measurement of Pclet number, Pe=1.32, in H.sub.2 tracer studies with step-change or impulse input to inert catalyst proxies. Kinetic limits to methane pyrolysis were quantified by Damkhler number, Da, for synthesis of benzene, Da.sub.B=0.15, and naphthalene, Da.sub.N=0.03, determined from kinetic studies which rigorously accounted for reversibility of DHA reactions. Detailed reaction-transport models synthesized interplay of kinetic, diffusive, and convective length-scales captured by Pclet and Damkhler numbers to predict influence of catalyst-absorbent proximity and process flow-conditions on aromatization rates. Systematic control of catalyst bed-length, L, or linear flow velocity, u, predictably altered Pe and Da to effect improvements in methane conversion with and without Zr metal, corroborating results from simulation of the reaction-transport model.
INTRODUCTION
[0123] Dehydroaromatization (DHA) of methane to benzene (6CH.sub.4.Math.C.sub.6H.sub.6+9H.sub.2) is highly endothermic and requires reaction temperatures950 K to achieve about 10% equilibrium conversion at ambient pressure. Carbidic forms of Mo encapsulated in medium-pore MFI zeolites catalyze methane DHA with high benzene selectivity (about 70%) at conversions near the thermodynamic limit. Oxidic Mo precursors, deposited either by aqueous impregnation or stoichiometric exchange with zeolitic H.sup.+ pairs, reduce and carburize upon exposure to CH.sub.4 at high temperatures (950 K) to form molecular-size MoC.sub.x clusters (0.6-1.5 nm) which activate CH bonds in methane to initiate CC coupling reactions that lead to aromatics. Oxygen removed (2.440.1 O:Mo) in the form of H.sub.2O, CO, and CO.sub.2 during reduction/carburization procedures corroborated formation of putative (Mo.sub.2O.sub.5).sup.2+ dimers which are anchored to proximal Al site pairs on MoO.sub.x/H-ZSM-5 materials prepared by solid-state stoichiometric exchange. Carbon deposition of about 9 C:Mo concurrent with oxygen evolution during catalyst reduction/carburization is the predominant source of coke formation on materials with about 3 wt % Mo.
[0124] H.sub.2 formed during dehydrogenation and cyclization events on Mo/H-ZSM-5 inhibits net synthesis rates during methane pyrolysis; in-situ removal of hydrogen from catalyst proximity is virtuous in overcoming thermodynamic barriers of methane DHA set by reaction endothermicity. It has been demonstrated that permselective membrane walls integrated into CH.sub.4 DHA flow reactors effectively scavenged liberated hydrogen, modestly enhancing methane conversion from about 10% to about 12%. However, improvement in methane pyrolysis rates was principally owed to increased formation of entrained C.sub.12+ polynuclear aromatics as product distribution shifts to higher molecular weight unsaturated hydrocarbons which benefit most from hydrogen removal. The accumulation of large polycyclic aromatics congests zeolite pores and leads to a more rapid decrease in instantaneous conversionfirst-order deactivation rate constants increase from 0.044 h.sup.1 to 0.059 h.sup.1 upon integration of permselective membrane walls.
[0125] Example 1 demonstrates that in-situ hydrogen removal can be achieved by introduction of a selective hydrogen absorbent, such as zirconium metal. Addition of Zr metal particles to MoC.sub.x/ZSM-5 in interpellet mixtures resulted in a 2.7 increase in methane converted and concurrent 1.4, 2.1, 2.6, and 5.6-fold increase in effluent benzene, naphthalene, toluene, and C.sub.10+ yields, respectively, in reference to a conventional MoC.sub.x/ZSM-5 catalyst at equivalent time-on-stream (8.7 ks). The maximum methane conversion on interpellet mixtures exceeded 27%in significant excess of about 10% equilibrium-prescribed conversion at these reaction conditionsdemonstrating that thermodynamic limitations encountered in methane dehydroaromatization on MoC.sub.x/ZSM-5 are circumvented by addition of Zr as a hydrogen absorbent. Characterization of Zr metal via hydrogen uptake experiments and X-ray diffraction evinced formation of ZrH.sub.1.75 on hydrogen exposure (3.28-95.13 kPa) at 973 K.
[0126] In this Example 2, the critical role of catalyst-absorbent proximity and reactor hydrodynamics in determining efficacy of polyfunctional mixtures to achieve supra-equilibrium aromatic yields in methane DHA reactions is disclosed. H.sub.2 tracer studies with impulse and step-change inputs to flow vessels charged with inert catalyst proxies revealed fluid flow was highly dispersed under reaction flow conditions. Measurement of axial dispersion coefficients, kinetic orders, and forward rate constants formed the foundation of detailed reaction-transport models used to simulate DHA reactions with and without Zr addition. Pclet number, Pe, and Damkhler number, Da, arose naturally from non-dimensionalization of governing transport equations and provided compendious metrics to evaluate significance of kinetic, diffusive, and convective length-scales. Interplay of kinetic and transport length-scales was studied by systematic change of catalyst-absorbent intimacy and flow conditions in both experimental investigations and model simulations. Reaction-transport models corroborate observed proximity effects and provided a well-founded mathematical framework to rationalize influence of reactor hydrodynamics captured by Pclet and Damkhler numbers.
Materials and Methods
[0127] 2.1. Catalyst Synthesis and Preparation.
[0128] Mo/H-ZSM-5, MoC.sub.x/H-ZSM-5, and polyfunctional configurations of MoC.sub.x/H-ZSM-5+Zr were prepared by methods detailed in Example 1 and summarized herein. Briefly, NH.sub.4-ZSM-5 (Zeolyst International, Si/Al=11.5, CBV 2314) was converted to H-ZSM-5 by treating in dry air (about 1.67 cm.sup.3 s.sup.1) to thermally decompose NH.sub.4.sup.+ to H.sup.+ and NH.sub.3 (g) by increasing the temperature from ambient temperature to 773 K at 0.0165 K s.sup.1 and holding at 773 K for 36 h. Intimate mixtures containing a nominal 3 wt % Mo loading (Mo/Al0.25) were prepared by grinding together MoO.sub.3 (Sigma-Aldrich, 99.9%) and H-ZSM-5 powders in an agate mortar and pestle for 0.25 h. The mixture was heated from ambient temperature to 623 K at 0.0167 K s.sup.1 and held at this temperature for 15 h in dry air (ca. 0.67 cm.sup.3 s.sup.1) resulting in water removal and dispersion of MoO.sub.3 on the external surface of the zeolite. Finally, the mixture was heated to 973 K at 0.167 K s.sup.1 and held at this temperature for 10 h to facilitate molybdenum oxide migration into the zeolite pores. The catalyst prepared via the above-mentioned protocol is referred to as Mo/H-ZSM-5 in this Example. The Mo/H-ZSM-5 powder was pressed to form pellets that were then crushed and sieved to obtain particle sizes between 180 and 425 mm (mesh 40-80) for use in catalytic reactions.
[0129] The samples designated as MoC.sub.x/H-ZSM-5 in this Example were prepared by exposing the Mo/H-ZSM-5 formulation to a CH.sub.4/Ar mixture (90 vol % CH.sub.4 and 10 vol % Ar, total feed flow rate ca.0.27 cm.sup.3 s.sup.1) for 15.5 ks (for catalyst loading 1.2 g) leading to complete carburization of the molybdenum oxide as described in Example 1. Subsequently, the feed was switched to helium (ca. 0.83 cm.sup.3 s.sup.1). The sample was flushed in helium flow (ca. 0.83 cm.sup.3 s.sup.1) at 973 K for 0.9 ks and then cooled to ambient temperature in the same helium flow.
[0130] Zirconium (Zr) metal was obtained from American Elements (PN # ZR-M-0251 M-GR.1T2MM, 99.5+% purity, metal basis) as granule-shaped particles (1-2 mm granule size) which were crushed and sieved to obtain particle sizes between 180 and 425 mm (mesh 40-80) for use in the catalytic reactions.
[0131] All polyfunctional configurations of MoC.sub.x/H-ZSM-5 and Zr were constituted in an inert environment using a glove bag (Sigma Aldrich, Z530212, AtmosBag, two-hand, non-sterile, size M, Zipper-lock closure type) to eliminate oxidation of MoC.sub.x species under ambient conditions. This procedure gave five fixed-bed configurations: (i) MoC.sub.x/H-ZSM-5 only, (ii) Zr packed upstream of the MoC.sub.x/H-ZSM-5 catalyst bed, (iii) Zr packed downstream of the MoC.sub.x/H-ZSM-5 catalyst bed, (iv) Zr packed both upstream and downstream of the MoC.sub.x/H-ZSM-5 catalyst bed (referred to as sandwich configuration), and (v) an interpellet mixture of MoC.sub.x/H-ZSM-5 and Zr metal. Reconfigured catalyst absorbent mixtures were subsequently loaded in the same quartz reactor under an inert atmosphere and transferred to the furnace without exposure to ambient conditions. The reactor was heated to the reaction temperature, 973 K, in helium flow (ca. 0.33 cm.sup.3s.sup.1) at 0.18 K s.sup.1 before performing methane dehydroaromatization reactions.
[0132] 2.2. Methane Dehydroaromatization Catalytic Reactions.
[0133] Methane dehydroaromatization reactions were performed in a fixed bed tubular quartz reactor (I.D. 10.5 mm). The catalyst sample was heated in helium flow (ca. 0.33 cm.sup.3 s.sup.1, UHP, Minneapolis Oxygen) from ambient temperature to the reaction temperature, 973 K, at 0.18 K s.sup.1 in a resistively heated furnace (National Element FA120). Axial temperature gradients in the reactor were minimized by using an annular inconel cylinder to ensure conduction of heat between walls of the furnace and the tubular reactor. The temperature of the catalyst sample was monitored by two thermocouples stationed diametrically on the exterior wall of the quartz reactor. A feed gas mixture of CH.sub.4/Ar (90 vol % CH.sub.4 and 10 vol % Ar, total feed flow rate ca. 0.27 cm.sup.3 s.sup.1, UHP, Matheson Tri-Gas) was introduced to the reactor to perform methane reactions at 973 K and atmospheric pressure, with Ar serving as an internal standard. All flow lines were heated to temperatures in excess of 450 K via resistive heating to prevent condensation of effluents. The composition of the reactor effluent was analyzed using a mass spectrometer (MKS Cirrus 200 Quadrupole MS system) and a gas chromatograph (Agilent 7890) equipped with a methyl-siloxane capillary column (HP-1, 50 m320 mm0.52 mm) connected to a flame ionization detector (FID) for detection of hydrocarbons and a GS-GasPro column (60 m0.320 mm) connected to a thermal conductivity detector (TCD) for detection of permanent gases (H.sub.2, Ar, and CH.sub.4). Transient product evolution throughout the course of the reaction was measured with an online mass spectrometer (MS) (MKS Cirrus 200 Quadrupole MS system) connected to the outlet of the GC. The number of O atoms removed during carburization of molybdenum oxide was determined by the cumulative amount of H.sub.2O, CO, and CO.sub.2 eluted as calculated from the transient MS signal with Ar as an internal standard.
[0134] Evaluation of methane DHA kinetic orders and rate constants was performed on MoC.sub.x/H-ZSM-5 in the regime when benzene net rate was nearly invariant with time-on-stream (15-25 ks time-on-stream for 1.2 g catalyst loading). H.sub.2 (0-0.23 kPa at 97.6 kPa CH.sub.4 and balance argon, total flow rate ca. 0.27 cm.sup.3 s.sup.1 and catalyst loading=1.2 g) and methane (2.7-97.6 kPa, balance argon, total flow rate ca. 0.27 cm.sup.3 s.sup.1 and catalyst loading=1.2 g) partial pressures were varied at constant contact time by altering H.sub.2, inert (Ar), and methane flow rates.
[0135] 2.3. Measurement of Axial H.sub.2 Diffusion Coefficient in Packed-Bed Flow Reactors.
[0136] Inert H.sub.2 tracer experiments were performed on beds of 180-425 m quartz sand aggregates packed to 2.7 cm, an identical length as 1.2 g of 180-425 m MoC.sub.x/H-ZSM-5 aggregates in a 1.05 cm ID tubular quartz reactor. Quartz sand was heated to 973 K in inert (He or Ar) flow (ca. 1 cm.sup.3 s.sup.1). For pulse tracer experiments, a sample of H.sub.2 (ca. 1 cm.sup.3) was introduced to a stable reactor feed flow of CH.sub.4 using an electronic six-way valve (ED66UWE, VICI Valco 6-port valve). In step-change experiments, reactor feed flow was switched from Ar to an Ar and H.sub.2 gas mixture (5 vol % H.sub.2/balance Ar). Reactor effluent was monitored using an online mass spectrometer (MS) (MKS Cirrus 200 Quadrupole MS system).
Results and Discussion
[0137] Rates of non-oxidative methane conversion to benzene, naphthalene, methylarenes, and C2 hydrocarbons were measured over 1.2 g of MoC.sub.x/H-ZSM-5 catalysts (Mo/Al.sub.f=0.25) with and without 2.4 g of Zr metal at a temperature of 973 K and 13.0 cm.sup.3 s.sup.1 flow of 90%/10% CH.sub.4/Ar in five reactor configurations listed in Section 2.1.
[0138] 3.1. Methane DHA on MoC.sub.x/H-ZSM-5 in Absence and Presence of Zr Metal.
[0139]
[0140] Cumulative product yield at time-on-stream, t, is defined as the total number of moles of product formed per total moles of Mo in the catalyst bed at the end of time, t, and is obtained from the area under the curve of the net product formation rate vs. time-onstream plots shown in
[0141] These results demonstrated that irrespective of spatial intimacy between catalyst and Zr metal, the polyfunctional MoC.sub.x/H-ZSM-5 and Zr absorbent formulation maintains efficacy to enhance methane conversion without detriment to aromatic selectivity. Ability of staged/stratified zirconium to continuously abstract hydrogen and effect improvement to net rates demonstrates gas phase H.sub.2 is conveyed dispersively to Zr absorbent both upstream and downstream of catalyst under reaction conditions. Significance of dispersive transport considerations and influence of catalyst absorbent proximity can be rigorously described by differential mole balance on catalyst and absorbent beds. The resultant reaction-transport model, presented in Section 3.3, requires determination of reaction rate equations and measurement of relevant kinetic and transport parameters, as discussed next in Section 3.2.
[0142] 3.2. Evaluation of Methane DHA Kinetic and Transport Parameters on MoCx/H-ZSM-53.2.1. Kinetics of Methane DHA on MoCx/H-ZSM-5.
[0143] Kinetic measurements for methane DHA were performed on pre-carburized MoCx/H-ZSM-5 formulations after exposing Mo/HZSM-5 (1.2 g catalyst loading) to CH.sub.4/Ar mixtures (90 vol % CH.sub.4/balance Ar) for 15.5 ks. The forward rate of benzene and naphthalene formation for the stoichiometric reactions of CH.sub.4 to H.sub.2 and C.sub.6H.sub.6(Eq. (1)) and CH.sub.4 to H.sub.2 and C.sub.10H.sub.8 (Eq. (2)) on MoCx/HZSM-5 at 973 K were obtained using a methodology reported previously.
CH.sub.4.Math.C.sub.6H.sub.6+ 3/2H.sub.2(1)
CH.sub.4.Math. 1/10C.sub.10H.sub.8+ 8/5H.sub.2(2)
[0144] Briefly, the net rate of benzene and naphthalene formation and approach to equilibrium of benzene (B) and naphthalene (N) was utilized,
calculated from gas-phase total pressure, P, partial pressures, P.sub.j, and equilibrium constants K.sub.B and K.sub.N, to quantify forward rates of benzene formation, {right arrow over (r.sub.B)}=(5.050.09)10.sup.4 mol s.sup.1 mol.sub.Mo.sup.1, and naphthalene formation, {right arrow over (r.sub.N)}=(0.990.01)10.sup.4 mol s.sup.1 mol.sub.Mo.sup.1, at 973 K, 90 vol % CH.sub.4/balance Ar, and 1 atm total pressure. The measured forward rate of benzene formation was similar to rates reported by others (4.70.810.sup.4 mol s.sup.1 mol.sub.Mo.sup.1) and (5.0310.sup.4 mol s.sup.1 mol.sub.Mo.sup.1) for methane DHA on Mo/H-ZSM-5 at 950 K.
[0145]
[0146] Benzene similarly has no kinetic effect on the stoichiometric reaction of CH.sub.4 to H.sub.2 and C.sub.6H.sub.6(Eq. (1)) on Mo/H-ZSM-5 at 950 K as evident from an invariant benzene forward rate ((3.80.5)10.sup.4 mol s.sup.1 mol.sub.Mo.sup.1) at varying outlet benzene partial pressure (0.1-0.6 kPa) achieved via different catalyst loadings (0.05-1.0 g Mo/H-ZSM-5). It was presumed that naphthalene also had no kinetic effect on aromatization rates. Methane kinetic orders for benzene and naphthalene formation are 0.72 and 0.74 (
{right arrow over (r.sub.B)}+k.sub.BP.sub.CH.sub.
{right arrow over (r.sub.N)}+k.sub.NP.sub.CH.sub.
[0147] 3.2.2. Measurement and Calculation of Pclet Number, Pe.
[0148] The observed effects of catalyst-absorbent proximity on aromatic synthesis rates during methane DHA reactions, particularly those noted upon introduction of Zr upstream and/or downstream of the catalyst bed, highlight the critical role of dispersive H.sub.2 transport in polyfunctional configurations. Efficacy of staged Zr beds to enhance aromatic yields to a similar degree as in interpellet mixtures demonstrates gas-phase H.sub.2 generated over Mo/H-ZSM-5 catalysts is capable of bed-scale transport both co- and counter-current to bulk advection. Relative contributions of diffusive/dispersive and convective transport in tubular flow reactors can be assessed based on a Pclet number
which is the non-dimensional ratio of characteristic rates of diffusive and convective gas-phase transport in tubular flow reactors. Pe provides a useful metric to evaluate significance of axial dispersion considerations in terms of catalyst bed-length, L, effective diffusion coefficient, D.sub.eff, and total linear flow velocity, u, or superficial linear flow velocity, u.sub.s, and catalyst bed void fraction, E. Ideal plug flow reactors (PFRs) operate in the limit of Pe.fwdarw. wherein unidirectional bulk advection is the sole mechanism of mass transport, thereby rendering neighboring axial fluid-elements incapable of exchanging matter. Axial dispersion becomes relevant for Pe which permits bed-scale molecular transport in response to concentration gradients arising from chemical reaction and/or absorption rate processes either (pseudo)-homogenously or at reactor bounds.
[0149] Introduction of catalyst or absorbent packed-beds in tubular flow reactors obstruct molecular movement, thereby shortening mean-free path and dampening diffusional motion. Consequently, effective bulk gas-phase diffusion coefficients, D.sub.eff, decrease compared to the molecular limit, D, which defines proportionality between diffusive flux and species concentration gradients in an unobstructed fluid medium. Substitution of D for D.sub.efft into Eq. (7) provides an absolute lower-bound on Pe. D.sub.eff, and Pe were evaluated by theoretical and experimental means to corroborate observations of bed-scale hydrogen diffusion during methane dehydroaromatization reactions in polyfunctional catalyst+absorbent formulations.
[0150] Levenspiel details systematic procedures for approximation of Pclet number for conditions which deviate from the ideal plug flow reactor. Calculation of Bodenstein number
Bo=ud/D,(8)
and catalyst-bed aspect-ratio,
A=L/d,(9)
where d is reactor diameter, give Bo=0.12 and A=2.57 and suggest reactor configurations discussed in this investigation are neither ideally-dispersed nor convectively-controlled and thus prohibit calculation of Pe by asymptotic, closed-form solution of the Navier-Stokes equation]. Application of Darcy's law and Poiseuille's law to a theoretical treatment of viscous fluid flow presented by Marshall gives a simple relation between D, D.sub.eff, and packed-bed porosity,
[0151] Others report mutual diffusion coefficients for H.sub.2 in a variety of gaseous mixtures measured by the Taylor dispersion method and fit these data to power law models with coefficients similar to those predicted by the kinetic theory of gases. From reported correlations, the following was calculated, D (973 K)=6.30.2 cm.sup.2 s.sup.1 giving, as an absolute lower-bound, Pe0.3 using Eq. (7). Taking =0.35, a typical value for packed bed columns, Pe was found to be 1.300.2 under reaction conditions conforming with the balance of convective and diffusive time scales which the Bodenstein number and aspect ratio suggest.
[0152] To corroborate approximations of Pe from theoretical calculation, Pe was measured by both step-change and pulse input of H.sub.2 through an inert bed of quartz sand particles sieved to identical size and packed to identical bed-length compared to 1.2 g MoCx/H-ZSM-5 catalyst beds used in methane DHA reactions. Pulse and step-change tracer experiments were performed changing reactor diameter (0.68-1.05 cm), temperature (303973 K), and total flow-rate (17.6-40 cm.sup.3 s.sup.1) to note dependence of D.sub.eff solely on temperature. Following a formulism presented by Levenspiel, residence-time (exit-age) distribution, E, and mean residence time, t.sub.avg, was calculated from effluent H.sub.2 pressure histories for both pulse and step-change inputs. E is the distribution of exit ages of fluid elements which elute from the flow vessel (packed-bed reactor) and is defined as a normalized quantity such that .sub.0.sup.Edt=1. Profiles of non-dimensional exit-age distribution
E.sub.0+t.sub.avgE,(11)
as a function of non-dimensional time elapsed from tracer entry
are compared to analytical solutions of differential mole balances for a perfect pulse (i.e. Dirac-delta function) or perfect step-change (i.e. Heaviside function) of inert gas passing through an axially-dispersed tubular reactor (see Supporting Information Section S3) with Pe fit by mean-square error minimization, as shown in
[0153] Table 4 collates D.sub.eff and Pe measured by H.sub.2 tracer studies and calculated using Eq. (7). Pclet numbers found from fit to experimental tracer response curves range from 1.8 to 10 depending on reactor diameter, total flow-rate, and bed temperature. D.sub.eff calculated from Pe per Eq. (7), 1.4-1.5 cm.sup.2 s.sup.1 at 973 K, agree well with theoretical value of 1.5 cm.sup.2 s.sup.1 from Eq. (10) using =0.35. Taking an average experimental value D.sub.eff=1.43 cm.sup.2 s.sup.1 at 973 K gives Pe=1.32 at temperature and flow-rate relevant to methane DHA reactions. Pclet numbers near unity dictate a balance of convective and diffusive mass transport in tubular flow reactors and demonstrate the presence of axial dispersion phenomena in studied polyfunctional reaction configurations.
[0154] 3.3. Reaction-Transport Model.
[0155] A 1-dimensional (1D) reaction-transport model is presented for a steady-state axially-dispersed packed-bed reactor inclusive to all MoCx/H-ZSM-5+Zr formulations. Methane DHA reactions and H.sub.2 absorption are treated as pseudo-homogenous processes in catalyst and absorbent beds, respectively. Methane conversion solely to benzene and naphthalene was considered, which together constitute 92% carbon selectivity. Benzene+naphthalene (B+N) yield in reaction-transport simulations is thus identical to methane conversion and permits comparison to observed rates without complications conferred by methane conversion to entrained carbon and/or coke. Differential mole balance in the catalyst bed yields non-dimensional equations:
where y.sub.j is dimensionless gas-phase partial pressure, normalized by 1 bar, of component j, x is dimensionless length, and subscripts M, H.sub.2, B, and N, refer to methane, hydrogen, benzene, and naphthalene respectively. Mole balance in the absorbent beds staged upstream and/or downstream of catalyst gives
[0156] The final term in Eq. (14) is only included in interpellet mixtures wherein Eqs. (17)-(20) are unnecessary. Eqs. (13)-(20) collapse to those of an ideal PFR in the limit Pe.fwdarw. in which bulk convection is the dominant mode of mass transport (see Supporting Information Section S4,
[0157] Non-dimensionalization of constitutive mass balance equations for a packed-bed reactor gave a rise to dimensionless parameters:
where u is linear velocity, L is the length of the catalyst bed, D.sub.eff is effective diffusivity, and k.sub.B, k.sub.N, and k.sub.Zr are 1st order rate constants for the benzene synthesis reaction, naphthalene synthesis reaction, and H.sub.2 absorption, respectively.
[0158] Rate equations for forward rates of benzene and naphthalene synthesis are determined by kinetic measurements which rigorously account for influence of H.sub.2, C.sub.6H.sub.6, and C.sub.10H.sub.8 partial pressures on reversibility of methane pyrolysis reactions (Section 3.2.1).
[0159] Axial changes in total molar flow-rate, arising from stoichiometric imbalance of aromatization reactions and in-situ removal of H.sub.2, was assessed by alteration of total pressure, P; Da and Pe remain axially invariant. Axial changes in total pressure are, at maximum, 15% and are typically 5-10% during methane DHA reactions as predicted by simulation and confirmed by GC analysis of reactor effluent.
[0160] Da and Pe.sup.1 reflect ratios of the characteristic convective length-scale to kinetic and diffusive length-scales, respectively. Pe and Pe.sub.Zr are Pclet numbers in the catalyst and Zr bed, calculated using D.sub.eff measured by H.sub.2 tracer pulse and step-change experiments (Section 3.2.2). Da.sub.B and Da.sub.N, both order unity, are Damkhler numbers for forward synthesis rates of benzene and naphthalene. H.sub.2 removal by Zr is transport-limited under reaction-relevant process flow conditions, evidenced by breakthrough curves in H.sub.2 uptake experiments which resemble Heaviside functions. Thus, kinetic measurement of intrinsic first-order absorption rate constants, k.sub.Zr, is inaccessible in current reactor configurations. Da.sub.Zr>10.sup.4 give essentially invariant axial methane conversion and H.sub.2 pressure profiles and are in quantitative agreement with net synthesis rates (Section 3.4). Others have reported activation energies and pre-exponential factors for first-order hydrogen absorption determined by dehydriding pressure-buildup experiments and thermogravimetric studies which give Da.sub.Zr10.sup.6, conforming with Da.sub.Zr>10.sup.4 per results of reaction-transport simulations.
[0161] Danckwerts boundary conditions are applied at both bounds of the catalyst and absorbent beds.
where x.sub.entrance refers to inlet of the catalyst or absorbent bed furthest upstream, x.sub.exit refers to outlet of the catalyst or absorbent bed furthest downstream, and x.sub.cat/abs refers to bounds between catalyst and absorbent. Eqs. (26) and (27), generic closed Danckwerts boundary conditions, account for discontinuous change in dispersion length-scales between entrance and reaction/absorption sections and reaction/absorption and exit sections, respectively. It was presumed flow lines upstream and downstream of the reactor were dominated by convective flow such that Pclet numbers in entrance and exit sections far exceed those in reaction and absorption sections. Eqs. (28) and (29), generic open Danckwerts boundary conditions, demand continuity of mass and of diffusive flux between catalyst and absorbent beds which are of comparable length and thus are characterized by similar Pclet numbers.
[0162] Numerical simulation of the reaction-transport model, Eqs. (13)-(20), with appropriate boundary conditions, Eqs. (26)-(29), provides a useful tool to understand observed proximity effects in polyfunctional reaction configurations by inspection of axial product profiles otherwise unavailable by experimental means.
[0163] 3.4. Measurement and Simulation of CH.sub.4 Aromatization with In-Situ Hydrogen Abstraction.
[0164]
[0165] Maximum possible single-pass B+N yield over MoC.sub.x/H-ZSM-5 catalysts, henceforth referred to as the kinetic limit, is achieved in the complete absence of H.sub.2 (i.e. .fwdarw.) and is limited solely by reaction kinetics and contact-time. The kinetic limit, 17% for all configurations in
[0166] In all formulations, addition of Zr enhances methane conversion to benzene+naphthalene to near the kinetic limit, 17%. Interpellet physical mixtures provide moderate improvements to maximum single-pass B+N yield compared to layered formulations, which yield similar aromatic production in either downstream or upstream packing of Zr. Similarity of effluent aromatic yield in staged configurations is reflected in symmetry of axial H.sub.2 profiles; equivalent removal of H.sub.2 at low or high contact times manifests equivalent enhancement to aromatic synthesis. Partitioning zirconium in sandwiched configurations macroscopically mimics interpellet mixtures, improving B+N yield compared to singularly layered formulations; interpellet physical mixtures correspond to the limit of an infinite number of partitioned/stratified catalystabsorbent layers.
[0167] Non-zero B+N yield and H.sub.2 pressure at reactor inlet in certain configurations arises from closed Danckwerts boundary conditions, Eqs. (26) and (27), which demand discontinuous change in species concentration from x.sup.+.sub.entrance to x.sup..sub.entrance as dispersion considerations become relevant upon transition from upstream flow lines, wherein Pe.fwdarw., to the finite reaction and/or absorption section. Interpellet mixtures of zirconium and MoCx/H-ZSM-5 beget near-complete removal of hydrogen owing to proximity of catalyst and absorbent. Response of H.sub.2 axial profiles throughout the catalyst bed in staged configurations results from balance of convective and diffusive length scales (Pe1) which permits motility of gas phase hydrogen in response to chemical potential gradients set by favorable thermodynamics of metal hydride formation. Axial H.sub.2 pressures rise with distance from staged Zr beds as diffusional requirements rapidly increase (L.sub.diffusion(x).sup.2), preventing complete removal of hydrogen seen in interpellet mixtures. Staging zirconium before and after MoCx/H-ZSM-5, at identical mass loading, shortens average separation between catalyst and absorbent and results in suppressed H.sub.2 pressures throughout the catalyst bed. Ability of Zr upstream of MoC.sub.x/H-ZSM-5 to scavenge H.sub.2 is conferred by open Danckwerts boundary conditions, Eqs. (28) and (29), which enables back-mixing into the absorbent bed.
[0168] Observed aromatization rates and simulated axial B+N yield and H.sub.2 pressure profiles demonstrate synergistic interaction between MoCx/H-ZSM-5 and Zr succeeds in influencing reversibility of methane DHA reactions despite bed-scale separation of catalyst and absorbent in layered formulations, suggesting catalyst absorbent proximity is crucial only to the extent that physical rate processes convey intermediates (e.g. H.sub.2) between the absorbent and catalytic functions at time scales similar in magnitude to bulk advection and characteristic kinetic rates (i.e. PeDa1).
[0169] 3.5. Control of Diffusive and Convective Length-Scales.
[0170] The impact of convective and diffusive length scales on methane conversion was investigated in the steady-state catalytic regime by (i) change of catalyst bed-length (equivalently catalyst mass loading) and (ii) change of linear velocity (equivalently total flow rate) in reactor configurations with Zr packed downstream of MoC.sub.x/H-ZSM-5 at identical catalyst loading, reaction temperature, and reactor diameter. Systematic modification of Da and Pe.sup.1, by change of L (L/L.sub.o=0.2-1.6) and u (u/u.sub.o=1-8), provides a stringent test of the presented reaction-transport model and elucidates the key role of diffusive transport in polyfunctional staged-bed formulations. Direct influence of Da and Pe was assessed by use of the reaction-transport model;
[0171]
[0172] Decrease of u/u.sub.0 from 8 to 1, data II to data I in
[0173] Data I-III in
[0174] 3.6. Impacts of Catalyst-Absorbent Proximity on MoCx/H-ZSM-5 Deactivation.
[0175] Apparent loss of methane DHA activity in polyfunctional mixtures is complicated by conflation of dynamics attributable to catalytic CH.sub.4 activation and stoichiometric absorption of H.sub.2 into bulk Zr. As a metric for deactivation, decay in benzene yield, concomitant with decrease in total methane conversion, was considered, which resulted from (i) a deactivation of the MoCx/H-ZSM-5 catalyst and (ii) gradual saturation of Zr hydrogen absorption capacity. Others have demonstrated the kinetics of deactivation on conventional Mo(Cx)/H-ZSM-5 catalysts are well-described by a first-order deactivation rate constant, which assumes decrease in activity arises solely from a loss of catalytic active sites, S, at a rate
[0176]
[0177]
[0178]
[0179] In summary, introduction of a continuous hydrogen scavenging function via interpellet and staged fixed-bed configurations of MoCx/H-ZSM-5 and Zr metal circumvents intrinsic thermodynamic limitations of methane DHA at 973 K, thereby, increasing single-pass methane conversion to near the kinetic limit (14% conversion to benzene and 17% conversion to benzene+naphthalene), as dictated by forward synthesis rates. Zr metal beds staggered downstream and/or upstream of the MoCx/H-ZSM-5 catalyst bed results in 1.5-1.7 fold enhancement in cumulative product yield due to transient hydrogen absorption by Zr accelerating methane conversion to aromatics. A reaction-transport model accounting for axial dispersion in packed bed reactors elucidates the underlying role of reactor hydrodynamics in lab-scale investigations. Simulated axial profiles of hydrogen and hydrocarbon partial pressure reveal Pclet number values near unity are required for efficacious removal of H.sub.2 from the proximity of MoCx/H-ZSM-5 catalyst by staged Zr. Inert H.sub.2 tracer studies with impulse and step-change input measure effective H.sub.2 dispersion coefficients, D.sub.eff=1.43 cm.sup.2 s.sup.1, and Pclet number, Pe=1.32, to confirm the significant role of axial dispersion at process conditions relevant to catalytic studies. Both reaction-transport simulations and kinetic measurements demonstrate control of diffusive and convective length scales by alteration of catalyst bed-length and total flow-rate systematically changes methane to benzene+naphthalene yield (3%-17%), thereby explicating the critical influence of H.sub.2 motility on efficacy of polyfunctional interaction between catalyst and absorbent. Impacts of in-situ H.sub.2 removal on catalyst deactivation are observed to be most severe and sometimes irreversible in configurations with high concentrations of H.sub.2 absorptive capacity at low contact times; interpellet mixtures or configurations with Zr packed solely downstream of MoCx/H-ZSM-5 have mild and reversible effects on catalyst deactivation.
[0180] Supplemental InformationS1. Hydrogen Balance for CH.sub.4 Dehydroaromatization on MoC.sub.x/H-ZSM-5 and Zr Metal in Various Fixed-Bed Configurations.
[0181] The quantity of hydrogen absorbed by zirconium metal during methane DHA reactions on all on MoC.sub.x/H-ZSM-5 and Zr fixed-beds configurations was determined by (i) quantifying hydrogen missing in the effluent and (ii) performing He TPD after reaction. Corresponding results are shown in
H.sub.missing=(2H.sub.2+4C.sub.2H.sub.x+6C.sub.6H.sub.6+8C.sub.7H.sub.8+10C.sub.8H.sub.10+8C.sub.10H.sub.8+10C.sub.10.sup.+)(4CH.sub.4 reacted)(S1)
[0182] S2. CH.sub.4 Dehydroaromatization on MoC.sub.x/H-ZSM-5 and Zr Metal in Various Fixed-Bed Configurations.
[0183] Cumulative methane conversion and product yields during methane DHA reactions were compared at 973K on five fixed-bed configurations of MoC.sub.x/H-ZSM-5 and Zr.
[0184] S3. Elaboration on Measurement of Pclet Number by Tracer Response Curves.
[0185] Section 3.1.2 describes measurement of Pclet number by step-change and impulse of H.sub.2 tracer through an inert quartz sand bed. In this supplemental section, equations necessary for calculation of residence-time distribution are detailed and provide an example of tracer response curves. Formalism summarized in Levenspiel finds residence time distribution, E, from concentration-time curves of H.sub.2 elution. Integrating the concentration of H.sub.2 resultant from impulse, C.sub.pulse, over duration, t, gives exit-age distribution
Residence-time distribution is non-dimensionalized by mean duration of the C.sub.pulse curve
as stated in Section 3.2.2
E.sub.=t.sub.avgE(11)
Non-dimensional duration is
=t/t.sub.avg.(12)
Considering flow vessels with closed boundary conditions, as was done in for the reaction system, tracer concentration-time curves resultant from step-input, C.sub.step, are transformed to E curves by Eq. (S4)
where C.sub.max is the tracer maximum concentration following step-change input.
E.sub. curves are fit to predictions of an unsteady differential mole balance of an axially dispersed tubular reactor
where C is non-dimensional concentration of tracer at a position x along the inert bed and time . Applying closed-closed boundary conditions with initial condition C(x, =0)=0 gives
and eigenvalues, .sub.i, are found by numerical solution to transcendental equation
Solution to Eqs. S5-S7 are used for fits to measured E curves shown in
[0186] S4. Simulation of CH.sub.4 Aromatization with In-Situ Hydrogen Abstraction for Ideal Plug Flow Reactor.
[0187] The reaction-transport model presented in Section 3.3 was used to predict the hydrogen partial pressure profile and methane single-pass conversion along the axial coordinate of the catalyst bed for an ideal plug flow reactor with Pe=100. The results of the simulation are shown in
[0188] S5. Calculation of H.sub.2 Mutual Diffusion Coefficients from Correlative Relations.
[0189] Others reported mutual diffusion coefficients of dilute H.sub.2 in air, N.sub.2, and O.sub.2 at atmospheric pressure over a wide range of temperature (303 K-503 K). Data are fit to correlations resembling power-law dependences predicted by the kinetic theory of gases
D=AT.sup.B(S9)
where A and B are fitting parameters listed below.
[0190] S6. Reaction-Transport Simulations: Effects of Damkhler and Pclet Number.
[0191] Section 3.5 examines the key interplay of kinetic, convective, and diffusive length-scales in catalyst-absorbent mixtures for methane DHA by systematic change of catalyst bed-length, L, and linear velocity, u.
TABLE-US-00005 TABLE 4 Peclet number, Pe, and effective diffusion coefficients, D.sub.eff, determined from fit to H.sub.2 tracer responses using Eqs. 11 and 12 and calculated from correlation 28 and Eqs. 7 and 10. Reactor Temperature Flow Rate Diameter D.sub.eff.sup.b Method [K] [cm.sup.3 s.sup.1] [cm] Pe.sup.b [cm.sup.2 s.sup.1] H.sub.2 Pulse in 973 20.0 0.68 5.0 1.4 CH.sub.4 (FIG. 24A) H.sub.2 Pulse in 973 40.0 0.68 10 1.4 CH.sub.4 (FIG. 24A) H.sub.2 Step- 303 17.6 1.05 5.0 0.5 (2.3.sup.e) Change in Argon (FIG. 24A) H.sub.2 Pulse in 973 17.6 1.05 1.8 1.5 CH.sub.4 (FIG. 24A) Eqs. (7) and .sup.973.sup.d 13.0.sup.d 1.05.sup.d 1.3.sup.d 1.5.sup.d (10).sup.c .sup.aAll experiments performed at 973 K .sup.bCalculated taking = 0.35 .sup.cD taken from correlations given by others. .sup.dReaction conditions relevant to methane DHA reactions performed in this Example .sup.eAdjusted to 973 K assuming D.sub.eff ~T.sup.3/2
TABLE-US-00006 TABLE 5 Summary of deactivation constants for benzene calculated using Eq. 30 for MoCx-H-ZSM-5 and Zr in different fixed bed configurations. k.sub.d before H Fixed-bed configuration breakthrough/s.sup.1 K.sub.d after H breakthrough/s.sup.1 MoC.sub.x/H-ZSM-5 only NA 0.0094 Interpellet mixture.sup.a 0.1058 0.0113 Zr downstream.sup.b 0.0418 0.0109 Zr upstream.sup.c NA 0.0243 Sandwich.sup.d 0.0733 0.0171 .sup.aCorresponding data shown in FIG. 31A .sup.bCorresponding data shown in FIG. 31B .sup.cCorresponding data shown in FIG. 31C .sup.dCorresponding data shown in FIG. 31D
TABLE-US-00007 TABLE 6 Diffusion-coefficient correlation parameters from literature for intermolecular bulk counter diffusion of H.sub.2 and various gases or gaseous mixtures. System A/[cm.sup.2K.sup.Bs.sup.1] B H.sub.2-air 4.19 .Math. 10.sup.5 1.73 H.sub.2N.sub.2 4.68 .Math. 10.sup.5 1.71 H.sub.2O.sub.2 3.68 .Math. 10.sup.5 1.76 Using the reported values gives an average diffusion coefficient D (973 K) = 6.3 0.2 cm.sup.2 s.sup.1 across the three studied systems.
[0192] Other embodiments of the present disclosure are possible. Although the description above contains much specificity, these should not be construed as limiting the scope of the disclosure, but as merely providing illustrations of some of the presently preferred embodiments of this disclosure. It is also contemplated that various combinations or sub-combinations of the specific features and aspects of the embodiments may be made and still fall within the scope of this disclosure. It should be understood that various features and aspects of the disclosed embodiments can be combined with or substituted for one another in order to form various embodiments. Thus, it is intended that the scope of at least some of the present disclosure should not be limited by the particular disclosed embodiments described above.
[0193] Thus the scope of this disclosure should be determined by the appended claims and their legal equivalents. Therefore, it will be appreciated that the scope of the present disclosure fully encompasses other embodiments which may become obvious to those skilled in the art, and that the scope of the present disclosure is accordingly to be limited by nothing other than the appended claims, in which reference to an element in the singular is not intended to mean one and only one unless explicitly so stated, but rather one or more. All structural, chemical, and functional equivalents to the elements of the above-described preferred embodiment that are known to those of ordinary skill in the art are expressly incorporated herein by reference and are intended to be encompassed by the present claims. Moreover, it is not necessary for a device or method to address each and every problem sought to be solved by the present disclosure, for it to be encompassed by the present claims. Furthermore, no element, component, or method step in the present disclosure is intended to be dedicated to the public regardless of whether the element, component, or method step is explicitly recited in the claims.
[0194] The foregoing description of various preferred embodiments of the disclosure have been presented for purposes of illustration and description. It is not intended to be exhaustive or to limit the disclosure to the precise embodiments, and obviously many modifications and variations are possible in light of the above teaching. The example embodiments, as described above, were chosen and described in order to best explain the principles of the disclosure and its practical application to thereby enable others skilled in the art to best utilize the disclosure in various embodiments and with various modifications as are suited to the particular use contemplated. It is intended that the scope of the disclosure be defined by the claims appended hereto
[0195] Various examples have been described. These and other examples are within the scope of the following claims.