Process of making olefins or alkylate by reaction of methanol and/or DME or by reaction of methanol and/or DME and butane

Abstract

Methods of simultaneously converting butanes and methanol to olefins over Ti-containing zeolite catalysts are described. The exothermicity of the alcohols to olefins reaction is matched by endothermicity of dehydrogenation reaction of butane(s) to light olefins resulting in a thermo-neutral process. The Ti-containing zeolites provide excellent selectivity to light olefins as well as exceptionally high hydrothermal stability. The coupled reaction may advantageously be conducted in a staged reactor with methanol/DME conversion zones alternating with zones for butane(s) dehydrogenation. The resulting light olefins can then be reacted with iso-butane to produce high-octane alkylate. The net result is a highly efficient and low cost method for converting methanol and butanes to alkylate.

Claims

1. A staged reactor system for the synthesis of light olefins, comprising: a reactor comprising plural alternating first and second zones in a series of at least two first zones alternating with at least two second zones; and a flow path through the series of plural alternating first and second zones; wherein the first zones comprise a first catalyst and an inlet for methanol or DME; wherein the first catalyst comprises a crystalline Si/Ti zeotype material in which tetrahedral [TiO.sub.4] and [SiO.sub.4] units are arranged in a MFI structure with a three-dimensional system of channels having a molecular dimension of 4.9 to 5.9 Å, and at least 0.5 mass % Ti; wherein the second zones comprise a dehydrogenation or cracking second catalyst; wherein the first and second catalysts are different.

2. The staged reactor system of claim 1 comprising butane in the flow path and DME in the first zones.

3. The staged reactor system of claim 2 wherein the dehydrogenation second catalyst comprises at least 1, wt % with an upper bound of 25, or 20 wt % of Pt, Pd, Nickel, Cobalt, Copper, Zinc, Iron, Ru, Rh, Sn, or combinations thereof.

4. The staged reactor system of claim 3 where the second catalyst comprises a catalyst support comprising carbon, SiO2, Al.sub.2O.sub.3, TiO2, ZrO2, CeO2, Y.sub.2O.sub.3, Silica-Alumina, Zeolite Y, Zeolite USY, Zeolite ZSM-5, TS-1, Zeolite Beta, Zeolite Mordenite or combinations thereof.

5. The staged reactor system of claim 2 comprising multiple beds of the first catalyst and wherein DME is fed into the each of the multiple beds.

6. The staged reactor system of claim 1 wherein the dehydrogenation second catalyst comprises at least 2, or at least 5 wt %, or any of these minimum amounts with an upper bound of 20, or 15, or 10 wt % of Pt, Pd, Nickel, Cobalt, Copper, Zinc, Iron, Ru, Rh, Sn, or combinations thereof.

7. The staged reactor system of claim 6 where the second catalyst comprises a catalyst support comprising carbon, SiO2, Al.sub.2O.sub.3, TiO2, ZrO2, CeO2, Y.sub.2O.sub.3, Silica-Alumina, Zeolite Y, Zeolite USY, Zeolite ZSM-5, TS-1, Zeolite Beta, Zeolite Mordenite or combinations thereof.

8. The staged reactor system of claim 1 wherein the first catalyst comprises a crystalline zeotype material in which tetrahedral [TiO.sub.4] and [SiO.sub.4] units are arranged in a MFI structure with a three-dimensional system of channels having a molecular dimension of 5.1-5.6 Å.

9. The staged reactor system of claim 8 wherein the catalyst comprises at least 1% Ti.

10. The staged reactor system of claim 9 wherein the crystalline zeotype material comprises TS-1.

11. The staged reactor system of claim 9 wherein the catalyst contains at least 20 mass % TS-1.

12. The staged reactor system of claim 9 wherein the reactor system comprises a fixed bed reactor.

13. The staged reactor system of claim 8 wherein the catalyst comprises 1 to 5 mass % Ti.

Description

BRIEF DESCRIPTION OF THE DRAWINGS

(1) FIG. 1 Three step process for the conversion of methanol and butanes to alkylate.

(2) FIG. 2. Structure of Titanium-Silicalite molecular sieve FIG. 3. Distributed feed fixed-bed reactor design for the conversion of DME and butanes to light olefins.

(3) FIG. 4. Catalyst test unit for the conversion of DME and butanes to light olefins.

(4) FIG. 5. Long term stability test for conversion of DME to light olefins over TS-1 catalyst

(5) FIG. 6. Fixed-bed recycle reactor test unit for alkylating iso-butane with light olefins

DETAILED DESCRIPTION OF THE INVENTION

(6) Methanol and butanes can be converted to alkylate using a 3-step process as illustrated below:

(7) ##STR00001##
Step 1. Conversion of Methanol to Dimethyl Ether. Methanol can be partially dehydrated to an equilibrium mixture of dimethyl ether (DME), methanol and water over a solid-acid catalyst (typically γ-alumina). This reaction is rapid, reversible and exothermic. Water can be removed to drive the reaction. In cases where the starting material is DME rather than methanol, the first step can be omitted. Although DME is the preferred reactant for Step 2, methanol or a mixture of methanol and DME can be used in place of the DME in Step 2.
Step 2: Simultaneous conversion of DME and butanes to light olefins over a fixed-bed catalyst: DME and butanes are reacted in a coupled reaction to form light olefins such as ethylene, propylene, butylene and mixtures thereof. The conversion of DME to olefins is an exothermic reaction (ΔH=−55 kJ/mole) while the dehydrogenation of butanes to olefins is an endothermic reaction (ΔH=120 kJ/mole). Coupling the two reactions simultaneously, by reacting about 2 moles of DME per mole of isobutane results in a thermo-neutral reaction, which allows one to design a simple, fixed-bed process.
Step 3: Alkylation of iso-butane with light olefins: Water, BTX and hydrogen are first removed from the effluent from the second stage and then the remaining light olefins are mixed with excess iso-butane and reacted over a solid-acid catalyst to produce high-octane alkylate. The alkylation catalyst comprises a crystalline zeolite structure with a Si/Al molar ratio of 20 or less and up to 5 wt % of Pt, Pd and or Nickel. The reaction can be run liquid phase over a fixed-bed of catalyst at temperatures 35-90° C.

(8) One possible schematic for the 3 step process is illustrated in FIG. 1. The formation of DME is a conventional process and need not be described here. The alkylation step can be a conventionally known process but is preferably an improved process developed by Exelus Inc. that is described in an earlier patent.

(9) The reactant starting compositions for the inventive processes are methanol and/or DME and butane. The butane can be n-butane, isobutane, or typically a mixture of butanes.

(10) The simultaneous conversion of DME and butanes to light olefins (Step 2) has special challenges. The reactivity of butanes is much lower compared to DME. To compensate for the lower reactivity of butanes, researchers typically use high reaction temperatures (>600° C.) and strong solid acid catalysts to achieve adequate productivity. These solid acid catalysts, however, lead to poor selectivity for the conversion of DME to olefins.

(11) An important feature in the conversion of methanol/DME is the presence of the crystalline zeotype catalyst. The catalyst typically includes binders that do not significantly degrade the catalytic activity of the zeotype catalyst. The catalyst preferably contains at least 20 mass % of the crystalline zeotype catalyst; in some embodiments at least 50 mass % or at least 75 mass % of the crystalline zeotype catalyst. With respect to the catalyst, unless specified otherwise, mass % refers to the mass as a percent of the entire mass of the catalyst including both crystalline and noncrystalline material in the catalyst. A binder need not be present; however, in cases where binder is present, the stated mass % includes binder, unless specified otherwise such as the following: In some embodiments, the Ti content of the crystalline phase comprises at least 0.5 wt % Ti, or at least 1.0 wt % Ti, or from 0.5 to 3.0 wt %.

(12) We have discovered that titanium silicalite-1 (TS-1) is an excellent catalyst for the conversion of DME to light olefins and the simultaneous conversion of DME and butanes to light olefins. Titanium silicalite was developed in 1983 by research workers of Enichem. The synthesis of this catalyst is described, for example, in U.S. Pat. No. 4,410,501 (incorporated herein by reference) and Bruno Notari, “Microporous Crystalline Titanium Silicates”, Advances in Catalysis, vol. 41 (1996), pp. 253-334. The SiO2:TiO2 molar ratio of titanium-silicalite molecular sieve can range from 10 to 200. This zeolite shows several interesting properties in heterogeneous oxidation catalysis. TS-1 shows good activity and selectivity in alcohol oxidation, alkane oxidation, alkene epoxidation and ammoximation of cyclohexanone to cyclohexanone oxime.

(13) TS-1 is a highly structured zeolite type material comprising of titania and silica arranged in a MFI structure similar to ZSM-5 without the strong acidity. The structure is shown in FIG. 2. It has a 3-D micro-porous zeolitic structure with a medium pore size (i.e., channel opening size of about 5.5 A) which provides the right shape-selectivity towards lower molecular weight products. Generally, the titanium species in TS-1 can be categorized into two types: framework Ti atoms substituted into the silica lattices and extra framework Ti atoms. Ideally, the Ti atoms are supposed to be incorporated into the lattice as isolated entities and should be surrounded by four OSiO.sub.3 tetrahedra. Studies have shown that the overall acidity as well as the number of Brönsted acid sites increase with the increase in Ti content of TS-1. We have found that TS-1 based catalysts are excellent candidates for the simultaneous conversion of DME/butanes to light olefins for reasons listed below.

(14) 1. Moderate Acid Support: The high acid strength of zeolites like ZSM-5 and H-Beta favors reactions like the aromatization and hydrogenation of olefins suggesting that moderate or weak acid strength are preferred for the methanol or DME to olefins reaction. TS-1 which is a weak acid zeolite is thus a desirable support for converting methanol and its derivatives to light olefins.
2. Optimal Pore structure: TS-1 is a highly structured zeolite type material consisted of titania and silica arranged in a MFI structure similar to ZSM-5 without the strong acidity. It has a 3-D micro-porous zeolitic structure with a medium pore size (about 5.5 A) which provides the right shape-selectivity towards lower molecular weight products (e.g. light olefins) as opposed to amorphous catalysts like silica, silica alumina and zirconia.
3. High Hydrothermal Stability: The methanol or DME to olefins reaction generates significant amounts of water which results in gradual catalyst deactivation depending on the support used. TS-1 is routinely used as a catalyst for the formation of epoxides from olefins using organic peroxides which generates significant amounts of water and still displays excellent hydrothermal stability over a large period of time.
These characteristics provide guidance for selecting appropriate catalysts for the inventive process. The use of Ti-substituted zeolite provides the desired acid strength. Preferred ratios of Si/Ti are 15 to 150, or 20 to 100, or 25 to 50. The limiting pore size is controlled to accommodate the critical sizes of methanol (4.4 {acute over (Å)}), ethylene (4.2 {acute over (Å)}), propylene (5.0 {acute over (Å)}), and butene (5.1 {acute over (Å)}) while excluding aromatics (6.7 {acute over (Å)} or greater). The requisite hydrothermal stability can be easily measured and excludes the common aluminosilicates which are unsuitable for the inventive process.

(15) The same catalyst can be used for both of the coupled reactions—DME to olefins and butanes to olefins; however, due to its cost and low acid strength, TS-1 is not an ideal catalyst for the conversion of butanes to olefins. To improve its activity for dehydrogenation of butane, metals can be added to the TS-1 catalyst. Alternatively, catalysts having different acid strengths can be disposed in separate zones within a reaction chamber to conduct the coupled reaction. The DME to olefins catalyst preferably consists of at least 95% Si, Ti, and O; more preferably at least 97%; and in some embodiments at least 99% Si, Ti, 0, and, optionally, H. The butane to olefins catalyst, on the other hand, may contain a variety of metals. For example, in some preferred embodiments, the second catalyst comprises at least 1, or at least 2, or at least 5 wt %, or any of these minimum amounts with an upper bound of 25, or 20, or 15, or 10 wt % of Pt, Pd, Nickel, Cobalt, Copper, Zinc, Iron, Ru, Rh, Sn, or combinations thereof; these metals are disposed on a catalyst support; and in some preferred embodiments, the catalyst support is selected from the group consisting of carbon, SiO2, Al2O3, TiO2, ZrO2, CeO2, Y2O3, Silica-Alumina, Zeolite Y, Zeolite USY, Zeolite ZSM-5, TS-1, Zeolite Beta, Zeolite Mordenite or combinations thereof.

(16) The reaction chamber can contain a fixed bed reactor or other reactor type. The methanol or DME can be reacted over the Ti-substituted zeolite in the same reactor volume in which the butane(s) is reacted; more preferably, the reactions are conducted in separate zones—with the methanol/DME reaction zone containing the Ti-substituted catalyst (preferably without added metal) and the dehydrogenation and cracking of butane(s) conducted over a metal-containing catalyst in a second reaction zone. The zones can be alternated within a reactor with heat from the exothermic reaction of methanol/DME carried into the zone where the dehydrogenation and cracking of butane(s) is conducted. The reactor can be insulated (for adiabatic or near adiabatic operation) or operated in conjunction with a thermally-connected heat exchanger (typically adjacent or surrounding the reactor) for adding or removing heat to/from the reactor.

(17) This invention provides methods of converting isobutane and DME to light olefins comprising: passing a feed of isobutane and DME into a reaction chamber comprising alternating beds of a crystalline zeolite catalyst with MFI structure comprised of silicon and titanium oxides and up to 15 wt % of Pt, Pd, Nickel, Cobalt, Copper, Zinc, Iron, Ru, Rh, Sn or combinations thereof and a crystalline zeolite catalyst with MFI structure comprised of silicon and titanium oxides such that the DME feed is staged over the crystalline zeolite catalyst with MFI structure comprised of silicon and titanium oxides (substantially without added elements other than Ti) while mixed butanes flow in a plug flow manner over both groups of catalyst. The dehydrogenation/cracking catalyst need not have the MFI structure and could be any suitable dehydrogenation/cracking catalyst. In this case, it is desirable for the reaction to be carried out under conditions such that the DME is at least 90% consumed, preferably at least 95% consumed in the DME-conversion zone and prior to contacting the dehydrogenation/cracking catalyst.

(18) The invention also relates to a reactor system suitable for paraffin alkylation using solid acid catalysts. Fixed-bed reactors are easier to design, scale-up and maintain and, therefore, preferred embodiments utilize a fixed bed reactor. In some preferred embodiments of the invention, the DME feed is staged over multiple beds to control temperature profile. FIG. 3 schematically illustrates a distributed feed fixed-bed reactor design for the conversion of DME and butanes to light olefins.

(19) In this type of staged reaction, the initial drop in temperature due to butane dehydrogenation or cracking reaction is compensated by a corresponding rise in temperature by the DME to olefins reaction. The extent of butane dehydrogenation or cracking and DME conversion to olefins is controlled largely the respective residence times over the catalyst beds. A secondary benefit of this arrangement is the significant boost in olefin selectivity caused by the dilution effect from butane flow. A reduction in the partial pressure of the olefin leads to a large decrease in the amount of aromatics or paraffin production which allows a high selectivity to light olefins. Another benefit of this arrangement is the ability to tailor the catalyst to the specific reaction. Butane dehydrogenation or cracking is a slower reaction than the production of olefins from DME. The ability to stage alternating types of catalyst allows one to design one set of catalyst primarily for butane dehydrogenation/cracking reaction and another set of catalyst for the DME to olefins reaction. The typical fluidized bed reactors used for the MTG, MTO or CMHC process do not allow for such an approach.

(20) In some embodiments, the inventive methods may be further characterized by specified parameters. The conversion of DME and butanes to olefins is preferably conducted at a pressure between 1 atm and 10 atm. The reactants (methanol and/or DME and butane) preferably flow into the reaction chamber at a GHSV of at least 50 or at least 100 or between 100 and 1000 l/hr. In the broadest sense, the term “butane” should be understood as either n-butane or isobutane, or a mixture of butane types. In some embodiments, the method is conducted with a butane feed having an n-butane to iso-butane ratio of 0.1 to 10 mol/mol, or between 1:1 and 2.5:1; or wherein the butane comprises at least 90 mol % isobutane or at least 95 mol % isobutane. The ratio of butane/DME in the feed stream can be controlled to balance the enthalpies of reaction; preferably, the feed butane to DME ratio is in the range of 1.0 to 20 mol/mol. In operation, the catalyst(s) can be regenerated in flowing air or oxygen at a temperature of at least 400° C. and a GHSV of at least 500.

(21) For either the DME to olefin reaction or for the combined, coupled reactions, the olefin selectivity is >50 mol %, or >70 mol %, or >80 mol % and, in some embodiments with an upper limit of about 90%; the paraffin selectivity is <50 mol %, or <20 mol %, or <15 mol %, and in some embodiments in the range of about 20 to about 9%.

(22) The inventive methods (either DME to olefins or the coupled reaction with DME and butane) can run continuously and at steady state for a period of at least 8 hours or at least 12 hours without regenerating the catalyst. Over this period of steady state operation (without regenerating the catalyst). In preferred embodiments, the conversion of methanol and/or DME and the selectivity to C2-C6 olefins remains constant or changes 10% or less, or 5% or less, for the entire period without regeneration. In some preferred embodiments, the catalyst operates for at least 4 hours without regeneration or at least 8 hours without regeneration or at least 12 hours without regeneration and the conversion of methanol and/or DME and the selectivity to C2-C6 olefins remains constant or changes 10% or less or 5% or less for at least 10 days or at least 20 consecutive days of operation. The consecutive operation typically includes intervening regeneration steps; for example, operating continuously for at least 8 hours, followed by regenerating (typically by oxidizing with air), for example by oxidizing for 0.5 to 6 hours; and then again conducting the conversion to olefins.

(23) The step of reacting the olefins with iso-butane to form alkylate in a separate reactor (Step 3) can be conducted under known conditions. In preferred embodiments the alkylation reaction is conducted over a La-exchanged sodalite-containing catalyst as described in U.S. Published patent application Ser. No. 15/190,063 incorporated herein by reference as if reproduced in full below.

(24) Mukherjee et al. in U.S. Published patent application Ser. No. 15/190,063 describe a method of alkylating isobutane, comprising: under steady state conditions, passing a feed mixture of isobutane and C2-05 olefins (which is typically conducted in a continuous fashion) into a reaction chamber such that catalyst age is 2.5 or greater and producing 5 kg of alkylate product per kg of catalyst or greater wherein the olefin conversion remains above 90%, and the Research Octane Number (RON) of the products remains above 92. Steady state means that the selectivity to C8 isomers changes by 10% or less over a time period in which the 5 kg of alkylate product is produced per kg of catalyst. For example, a change in selectivity from 80% to 72% would be a 10% change. In this method, the reaction chamber comprises a crystalline zeolite catalyst; wherein the crystalline zeolite catalyst comprises sodalite cages and supercages, a Si/Al molar ratio of 20 or less, less than 0.5 weight percent alkali metals, and rare earth elements in the range of 10 to 35 wt %. Optionally, the catalyst may comprise up to 5 wt % Pt and/or Pd; and/or Nickel.

(25) The alkylation catalyst can be made by contacting the zeolite with a solution comprising a rare earth metal; calcining said catalyst by heating said catalyst to a calcination temperature of at least 575° C. to produce a catalyst intermediate comprising the rare earth metal and second concentration of alkali metal that is less than the first concentration of alkali metal; contacting the catalyst intermediate with an ammonium solution, drying to remove excess solution, and then heating the catalyst to generate the hydrogen (active) form of the zeolite—the deammoniation step. It is believed that the deammoniation step converts the ammonium cation sites to Bronsted acid sites, especially in the supercages, while the rare earth elements remain in the sodalite cages. Because the acid, or H+, sites are located in the larger diameter supercage structure of the catalyst, pore mouth plugging is significantly reduced, allowing the catalyst to remain active for increased periods of time, while the rare earth metal cation sites, such as, for example, La.sup.+3 sites, provide enhanced stability to the sodalite structure.

(26) The invention is further elucidated in the examples below. In some preferred embodiments, the invention may be further characterized by any selected descriptions from the examples, for example, within ±20% (or within ±10%) of any of the values in any of the examples, tables or figures; however, the scope of the present invention, in its broader aspects, is not intended to be limited by these examples.

EXAMPLES

(27) Catalyst screening experiments were performed using an isothermal packed bed reactor setup as shown in FIG. 4. Heating is controlled using an Omega temperature control unit and a ceramic heating element. Feeds are sent through a preheater of ˜75 cm length prior to entering the reactor.

(28) The catalyst of interest (1-10 gms) is first loaded into a reactor shown in FIG. 4, a center thermocouple (K-type) is inserted and positioned such that the tip of the thermocouple (3.1 mm diameter) is at the bottom of the catalyst bed. 1 mm glass beads are used to fill any void space in the reactor. The catalyst is activated in dry air at atmospheric pressure at 475° C. (4 hours). Following activation the reactor is then purged with dry nitrogen for 0.5 hours and flow of DME and/or butanes initiated.
Feed and products are analyzed using a HP5890 GC equipped with a Petrocol DH column.

Example 1—Catalyst A

(29) The starting material was a commercial zeolite TS-1, obtained from ACS Materials, (Medford, Mass.) having a Si/Ti molar ratio of ˜30 and a surface area of 360-420 m.sup.2/g. 5 grams of the zeolite was crushed and sieved to 0.5-1.4 mm particles.

Example 2—Catalyst B

(30) The catalyst is the same commercial TS-1 as Example 1, which was then impregnated with Ni and Zn. The catalyst was simultaneously impregnated with a solution of Ni(NO.sub.3).sub.2*6H.sub.2O and Zn(NO.sub.3).sub.2*6H.sub.2O. The salt solution was added, using incipient wetness impregnation, in amounts such that 15 weight percent Nickel was added to the catalyst and the Ni/Zn atomic ratio was 0.57.

(31) The catalyst was then calcined in an oven. The temperature program was: 120° C. (1 hour), 230° C. (2 hours), and 500° C. (4 hours). 5 g of the catalyst was pelletized, crushed and sieved to 1.18-1.4 mm particles

Example 3—Catalyst C

(32) The catalyst was prepared from a high surface area Al.sub.2O.sub.3 catalyst, obtained from Alfa Aesar. The catalyst was impregnated with a salt solution of Nickel nitrate and Zinc nitrate and calcined in the same way as Example 2. The catalyst was crushed and sieved into 1.18-1.4 mm particles.

Example 4—Catalyst D

(33) The catalyst was zeolite H-Beta with a Silica-Alumina Ratio (SAR) of 25 and was obtained from Zeolyst. Catalyst was in ammonium form with a surface area of 680 m.sup.2/g.

Example 5—Catalyst E

(34) The catalyst was zeolite H-ZSM-5 with a Silica-Alumina Ratio (SAR) of 30 and was obtained from Zeolyst. Catalyst was in ammonium form with a surface area of 425 m.sup.2/g.

Example 6—Catalyst F

(35) The catalyst was zeolite H-ZSM-5 with a Silica-Alumina Ratio (SAR) of 280 and was obtained from Zeolyst. Catalyst was in ammonium form with a surface area of 400 m.sup.2/g.

Example 7—Catalyst G

(36) The catalyst was high surface area silica with a pore diameter of 150 Å and a surface area of 300 m.sup.2/g and was obtained from Sigma-Aldrich.

Example 8—Catalyst H

(37) The catalyst was high surface area γ-Alumina in ⅛ inch pellet form and was obtained from Alfa Aesar.

Example 9—Catalyst I

(38) The catalyst was Na-Mordenite (obtained from Zeolyst) having a SAR of 13 and surface area of 425 m.sup.2/g. The catalyst was converted to the H.sup.+ form by first exchanging it with an aqueous solution of 0.5 M Ammonium Nitrate three times for 2 hours each at 80 C. The catalyst was then dried and calcined at 400 C in air.

Example 10—Catalyst J

(39) The catalyst was Amorphous Silica Alumina having a SAR of 14 and was obtained from Sigma-Aldrich. The particle size was 149 micrometer (100 Mesh).

Example 4—Catalyst K

(40) The starting material was a commercial zeolite X having a SiO2/Al.sub.2O.sub.3 molar ratio of 2.8 (Si/Al of 1.4) and a sodium content of 15% by weight. 5 grams of the zeolite was crushed and sieved to 0.5-1.4 mm particles. They were suspended in 50 mL of deionized water and stirred for 15 minutes after which the water was decanted. This washing procedure was repeated a second time.

(41) A lanthanum ion exchange was performed immediately following the initial water wash. The zeolite was suspended in 50 mL of a 0.2 M lanthanum nitrate solution and heated to 80° C. while stirring for 2 hours. The lanthanum solution was decanted and replaced with a fresh solution. This lanthanum exchange was performed three times followed by 2 water washes of 75 mL each. The zeolite was then left to dry at room temperature.

(42) Following the lanthanum exchange, the zeolite was calcined in a burnout oven. The temperature program for calcination was 1.5° C./min ramp to 100° C. where it was held for 1 hour, 2.0° C./min ramp to 230° C. and hold for 2 hours, 10° C./min ramp to the final calcination temperature of 400° C. for 4 hours.

(43) The lanthanum exchanged zeolite was suspended in a 0.5 M ammonium nitrate solution and heated to 80° C. with stirring for 2 hours. The ammonium solution was decanted and replaced with fresh solution. This ion exchange was performed 3 times followed by 2 water washes of 75 mL each. The zeolite was then left to dry at room temperature. The zeolite was deaminated in dry air (<2 ppm) using the following temperature program: 100° C. (0.5 hours), 120° C. (1 hour), 230° C. (2 hours), 400° C. (4 hours). 400° C. is the deamination temperature required to convert the catalyst from the ammonium form to the active proton form. The lower temperatures are necessary to completely dry the catalyst.

(44) 1. Catalyst A (titanium silicalite TS1 from ACS Material, Medford, Mass.) was used. Experiments were performed using an isothermal packed bed reactor setup described above. The catalyst (5 gm) is first loaded into a reactor shown in FIG. 3, a center thermocouple (K-type) is inserted and positioned such that the tip of the thermocouple (3.1 mm diameter) is at the bottom of the catalyst bed. 1 mm glass beads are used to fill any void space in the reactor. The catalyst is activated in dry air at atmospheric pressure at 475° C. (4 hours). Following activation the reactor is then purged with dry nitrogen for 0.5 hours. Iso-butane was fed at 5 gm/hr using a mass flow controller. The reaction was then conducted at 475° C., WHSV=1.0 l/hr and atmospheric pressure. The products were analyzed using a HP5890 GC equipped with a Petrocol DH column
The isobutane conversion was less than 1%.
2. Catalyst B (Ni/Zn/TS-1) was used. Experiments were performed using an isothermal packed bed reactor setup described above. The catalyst of interest (5 g) is first loaded into a reactor and activated in dry air. Iso-butane was fed at 5 gm/hr using a mass flow controller.
The isobutane conversion was 25% and olefin selectivity >80%.
3. Catalyst B (Ni/Zn/Alumina) was used. Experiments were performed using an isothermal packed bed reactor setup described above. The catalyst of interest (5 g) is first loaded into a reactor and activated in dry air. Iso-butane was fed at 5 gm/hr using a mass flow controller.
The isobutane conversion was 7% and olefin selectivity >90%.

(45) TABLE-US-00001 TABLE 1 Performance of catalysts for iso-butane conversion to light olefins iso-Butane Conversion Product Selectivity (%) Catalyst Description (%) Olefin Paraffin BTX COx A TS-1 0.4 22 75 3 0 B Ni-Zn/TS-1 25 82 15 3 0 C Ni-Zn/y-Alumina 7 90  9 1 0
The results clearly show that a bi-functional catalyst with both metal and adequate acid functions are required to convert iso-butane to light olefins.
4—Catalyst A titanium silicalite TS1 from ACS Material, Medford, Mass.) was used. Experiments were performed using an isothermal packed bed reactor setup. Heating is controlled using an Omega temperature control unit and a ceramic heating element. The catalyst (5 gm) is first loaded into a reactor shown in FIG. 3, a center thermocouple (K-type) is inserted and positioned such that the tip of the thermocouple (3.1 mm diameter) is at the bottom of the catalyst bed. 1 mm glass beads are used to fill any void space in the reactor. The catalyst is activated in dry air at atmospheric pressure at 475° C. (4 hours). Following activation the reactor is then purged with dry nitrogen for 0.5 hours. DME with nitrogen as diluent was fed at 5 gm/hr using a mass flow controller. The reaction was then conducted at 475 C, WHSV=1.0 l/hr and atmospheric pressure. The products were analyzed using a HP5890 GC equipped with a Petrocol DH column.
The test results are shown in table 2. The DME conversion remained constant at 100% for 24 hrs while the olefin selectivity >85 wt %.
5. Catalyst D zeolite H-Beta with a Silica-Alumina Ratio (SAR) of 25 was used. The catalyst was tested as in example 4. The test results are shown in table 2. The DME conversion was 78% while the olefin selectivity was 12 wt %.
6. Catalyst E The catalyst was zeolite H-ZSM-5 with a Silica-Alumina Ratio (SAR) of 30 and was obtained from Zeolyst. Catalyst was in ammonium form with a surface area of 425 m.sup.2/g. The catalyst was tested as in example 4. The test results are shown in table 2. The DME conversion was 100% while the olefin selectivity was 52 wt %.
7. Catalyst F The catalyst was zeolite H-ZSM-5 with a Silica-Alumina Ratio (SAR) of 280 and was obtained from Zeolyst. Catalyst was in ammonium form with a surface area of 400 m.sup.2/g. The catalyst was tested as in example 4. The test results are shown in table 2. The DME conversion was 100% while the olefin selectivity was 28 wt %.
8. Catalyst G The catalyst was high surface area γ-Alumina in ⅛ inch pellet form and was obtained from Alfa Aesar. The catalyst was tested as in example 4. The test results are shown in table 2. The DME conversion was 100% while the olefin selectivity was 0 wt %.
9. Catalyst G The catalyst was high surface area silica with a pore diameter of 150 Å and a surface area of 300 m.sup.2/g and was obtained from Sigma-Aldrich. The catalyst was tested as in example 4. The test results are shown in table 2. The DME conversion was 28% while the olefin selectivity was 12 wt %.
10. Catalyst I The catalyst was Amorphous Silica Alumina having a SAR of 14 and was obtained from Sigma-Aldrich. The particle size was 149 micrometer (100 Mesh). The catalyst was tested as in example 4. The test results are shown in table 2. The DME conversion was 100% while the olefin selectivity was 19 wt %.
11. Catalyst I The catalyst was Na-Mordenite (obtained from Zeolyst) having a SAR of 13 and surface area of 425 m.sup.2/g. The catalyst was converted to the 1-1±form by first exchanging it with an aqueous solution of 0.5 M Ammonium Nitrate three times for 2 hours each at 80° C. The catalyst was then dried and calcined at 400° C. in air. The catalyst was tested as in example 4. The test results are shown in table 2. The DME conversion was 87% while the olefin selectivity was 45 wt %.

(46) TABLE-US-00002 TABLE 2 Performance of catalysts for DME conversion to light olefins (C.sub.2—C.sub.6) DME Conversion Product Selectivity (%) Catalyst Description (%) Olefin Paraffin BTX COx A TS-1 100 89  9  2  0 D H-Beta(25)  78 12 75 13  0 E ZSM-5 (30) 100 52 27 21  0 F ZSM-5 (280) 100 28 61 11  0 G y-Alumina 100  0 66  0 34 H Silica  28  5 76  0 19 I Silica-Alumina (14) 100 19 77  1  3 J Mordenite (13)  87 45 52  3  0
As shown clearly above, Catalyst A (TS-1) is the best performing catalyst for the production of light olefins from DME.

(47) TABLE-US-00003 TABLE 3 Detailed product distribution for production of hydrocarbons from DME using catalyst A (TS-1) Hydrocarbon Olefin Paraffin BTX Methane 2.0% Ethylene  4.0% 0.1% Ethane 0.0% Propylene 40.0% Propane 0.4% Butanes 1.0% Butenes 28.5% Pentanes 1.0% Pentenes 13.0% Hexanes 3.0% Hexenes  3.0% Heptanes 2.0% Benzene 0.5% Toluene 0.8% Xylene 0.8% Total   89%   9%   2%
Data from table 2 above clearly indicate the high selectivity of desired olefin using TS-1 as the catalyst which results in a feed composition suitable to alkylate production.
12 Catalyst A titanium silicalite TS1 from ACS Material, Medford, Mass.) was used for a long term stability test. Experiments were performed using an isothermal packed bed reactor setup. Heating was controlled using an Omega temperature control unit and a ceramic heating element. The catalyst (5 gm) was first loaded into a reactor shown in FIG. 3, a center thermocouple (K-type) was inserted and positioned such that the tip of the thermocouple (3.1 mm diameter) was at the bottom of the catalyst bed. 1 mm glass beads were used to fill any void space in the reactor. The catalyst was activated in dry air at atmospheric pressure at 475° C. (4 hours). Following activation the reactor was then purged with dry nitrogen for 0.5 hours. DME was fed at 5 gm/hr using a mass flow controller. The reaction was then conducted at 475° C., WHSV (weight hourly space velocity)=1.0 l/hr and atmospheric pressure. After 8 hours, the DME flow was stopped and the catalyst regenerated in flowing air at 475° C. for 4 hrs. After regeneration was complete, the DME flow was re-started. The products were analyzed using a HP5890 GC equipped with a Petrocol DH column. The results are shown below. The DME conversion remained 100% for the entire duration of the test while olefin selectivity remained constant at >85 wt %. FIG. 5. shows the results for long term stability testing for conversion of DME to light olefins over TS-1 catalyst.
13 Catalyst A titanium silicate TS1 and Catalyst B, titanium silicate impregnated with nickel and zinc, were used for the combined test. Experiments were performed using an isothermal packed bed reactor setup shown in FIG. 4. 5 gm Catalyst A mixed with glass beads were first loaded into the reactor followed by 5 gm of Catalyst B mixed with additional glass beads loaded into the same reactor. Glass beads were added to fill any void space.
The catalyst was activated in dry air at atmospheric pressure at a GHSV of approximately 1000 hr.sup.−1. The temperature program used to activate the catalyst is: 80° C. (1 hour), 120° C. (1 hr), 230° C. (2 hours), 475° C. (at least 4 hours). Following activation the reactor is purged with dry nitrogen for 0.5 hours.

(48) The reaction was conducted at 475° C. and atmospheric pressure. A mixture of iso-butane and DME were fed to the reactor at a WHSV of 1.0 l/hr. The products were analyzed using a HP5890 GC equipped with a Petrocol DH column.

(49) The results for the combined test are shown below.

(50) TABLE-US-00004 TABLE 4 Performance of coupled iso-butane dehydrogenation reaction with DME conversion to light olefins (C.sub.2-C.sub.6) over TS-1 and Ni-Zn/TS-1 catalyst bed Parameter Results DME conversion (%) 100% Olefin Selectivity (%) 81 wt % Paraffin Selectivity (%) 15 wt % BTX Selectivity (%)  3 wt % Product Iso-butane to Olefin 10 Ratio (mole/mole)
The test clearly shows the benefit of using staged catalyst beds to produce a reactor effluent that is suitable as feed for producing alkylate using solid or liquid acid catalysts.
14. Alkylation activity experiments were performed using an isothermal packed bed reactor with product recycle as shown in FIG. 6. Heating is controlled using an Omega temperature control unit and a ceramic heating element. Feeds are sent through a preheater of ˜75 cm length prior to entering the reactor.

(51) The catalyst K (10 g) is first loaded into a reactor shown in FIG. 6 (7.9 mm diameter), a center thermocouple (K-type) is inserted and positioned such that the tip of the thermocouple (3.1 mm diameter) is at the bottom of the catalyst bed. 1 mm glass beads are used to fill any void space in the reactor. The catalyst is deaminated in dry air (GHSV=1000 hr-1) at atmospheric pressure using the following temperature program: 100° C. (0.5 hour), 120° C. (1 hour), 230° C. (2 hours), 400° C. (4 hours). Following deamination the reactor is allowed to cool to reaction temperature (45° C.), then purged with dry nitrogen (GHSV=1000 hr-1) for 0.5 hours. The reactor is pressurized (300 psig) with pure isobutane to begin the experiment.

(52) The reaction feed is contained in helium-purged Hoke cylinders. Isobutane and light olefins (composition shown above) were fed to the reactor using an Eldex HPLC pump All feed and product analysis uses this GC system with the following program: 60° C. (16 min), ramp at 15° C./min to 245° C. and soak (20 min).

(53) The experiment is run using an olefin hourly space velocity equal to 0.1 hr-1 and a feed I/O ratio of ˜10 with an olefin distribution that models the feed effluent from example 13. Product samples are extracted using a high pressure sampling port and syringe (Vici Precision Sampling) and immediately injected into the HP5890 GC for analysis.

(54) TABLE-US-00005 TABLE 5 Alkylate Product Composition for alkylating iso-butane with olefins produced by coupling iso-butane dehydrogenation/ cracking with DME to olefins reaction Concentration Component (wt %) Iso-pentane 7.06% n-pentane 0.00% 2,2-dimethylbutane 0.81% 2,3-dimethylbutane 0.00% 2-methylpentane 0.00% 3-methylpentane 0.74% 2,4-dimethylpentane 3.46% 2,2,3-trimethylbutane 0.22% 2-methylhexane 0.00% 2,2-dimethylpentane 0.22% 2,3-dimethylpentane  6.4% 3-methylhexane 0.22% 2,2,4-trimethylpentane 20.22%  Unknown C8 0.25% 2,2 dimethylhexane 0.00% 2,5 dimethylhexane 0.88% 2,4 dimethylhexane 2.28% 2,2,3-trimethylpentane 4.63% 2,3,4-trimethylpentane 13.24%  2,3,3-trimethylpentane 21.4% 2,3 dimethylhexane 1.62% 2-methylheptane 0.44% 3,4-dimethylhexane 2.57% 3-methylheptane 0.22% 3,3-dimethylhexane 0.22% 2,2,5-trimethylhexane 1.25% C9+ 11.32%  Alkylate Properties: Reid Vapor Pressure: 3.82 psi Specific Gravity: 0.7 RON: 95 MON: 92 Yield (vol/vol olefin) 1.85
The product distribution is shown in table 5. The olefin conversion remains at 100% for 24 hrs while the RON and MON numbers level off to 95 and 92.
Table 6 shows the near absence of aromatics in the alkylate for the Exelus process as well as a significant increase in MON rating—a jump of over 8 points. The results obtained are compared to a conventional MTG process (Methanol to Gasoline Technology Presentation, GTL Technology Forum 2014, Houston, Tex., (July 2014).

(55) TABLE-US-00006 TABLE 6 Product Composition for Exxon MTG process compared to the Exelus methanol/butanes to alkylate process Product Distribution, wt. % MTG Exelus M2Alk LPG (C1-C4) 23 2 C5+ Distribution Alkylate 96 (iso) Naphtha (C5-C9) 56 Aromatics 2 Benzene <0.1 Toluene 1 Xylene 6 TriMethylBenzene 7 TetraMethylBenzene 7 RON 92 95 MON 82 92 Reid Vapor Pressure, psi 12.3 3.82 Density, kg/L 0.73 0.7