Process for dehydroaromatization of alkanes with in-situ hydrogen removal

Abstract

A process for conversion of natural gas to aromatic hydrocarbons in a catalytic membrane reactor is described herein. The catalytic membrane reactor comprises a dehydrogenation catalyst and a membrane that can selectively transport hydrogen under high temperature operating conditions such as 600 C. to 800 C. Aromatic hydrocarbons are produced stably for a long time by a process characterized by hydrogen co-feed with the reaction gases to the one end of the to the reaction zone while hydrogen is extracted selectively with use of the membrane as the reactive gas mix passes through the reaction zone.

Claims

1. A process for the preparation of an aromatic hydrocarbon in a catalytic membrane reactor comprising a reactor inlet, a dehydrogenation catalyst and a reactor outlet in a first zone arranged so that a feed to the reactor via the inlet must contact the dehydrogenation catalyst before passing to the reactor outlet; and a hydrogen transport membrane separating said first zone from a second zone; said process comprising: (I) feeding through said reactor inlet a feed comprising hydrogen and at least one C1-4 alkane so that the alkane and hydrogen contact said dehydrogenation catalyst in said first zone; (II) operating said reactor at a temperature and pressure sufficient to allow dehydrogenation of said alkane and formation of said aromatic hydrocarbon and hydrogen; (III) allowing hydrogen to selectively pass through the hydrogen transport membrane into said second zone; (IV) recovering through the reactor outlet in the first zone, said aromatic hydrocarbon along with any unreacted C1-4 alkane and any hydrogen not extracted through the hydrogen transport membrane; (V) separating the aromatic hydrocarbon from the hydrogen and any unreacted C1-4 alkane; and (VI) recycling at least a part of said hydrogen and any unreacted C1-4 alkane to the feed using a recycle stream, wherein the hydrogen content of the recycle stream is 3 to 35 vol %.

2. The process of claim 1, wherein hydrogen transfer through said membrane take place in the form of protons.

3. The process of claim 2, wherein a partial pressure of hydrogen in the second zone of the catalytic membrane reactor is maintained at a low level by sweeping with an oxygen containing gas to convert at least some of the hydrogen to water in the second zone.

4. The process of claim 1, wherein the ratio of hydrogen to alkanes on a molar basis is higher at a feed side than at a product side of the catalytic membrane reactor.

5. The process of claim 1, wherein the ratio of hydrogen to alkanes on a molar basis is the same at a feed side and at a product side of the catalytic membrane reactor.

6. The process of claim 1, wherein an electrochemical driving force for the hydrogen transport across the hydrogen transport membrane is the result of an external voltage being applied to the membrane.

7. The process of claim 1, wherein an electrochemical driving force for the hydrogen transport across the hydrogen transport membrane is the result of a difference in partial pressures of hydrogen on the two sides of the membrane.

8. The process of claim 1, wherein said dehydrogenation catalyst is a zeolite catalyst.

9. The process of claim 8, wherein said zeolite has the structure MWW, CHA or WI.

10. The process of claim 9, wherein said zeolite is a ZSM-5.

11. The process of claim 1, wherein said transport membrane comprises at least one mixed metal oxide of formula (II)
Ln.sub.aW.sub.b-cMo.sub.cO.sub.12-y(II) wherein Ln is Y or an element numbered 57 to 71; the molar ratio of a:b is 4.8 to 6; c is 0 to (0.5b); and y is a number such that formula (II) is uncharged.

12. The process of claim 1, wherein the feed is natural gas.

13. The process of claim 1, wherein temperature in reactor is 500 to 1000 C.

14. The process of claim 1, wherein a portion of oxygen originating from an oxygen containing gas present in the second zone is transported from the second zone to the first zone.

15. A process for the preparation of an aromatic hydrocarbon in a catalytic membrane reactor comprising a reactor inlet, a dehydrogenation catalyst and a reactor outlet in a first zone arranged so that a feed to the reactor via the inlet must contact the dehydrogenation catalyst before passing to the reactor outlet; and a hydrogen transport membrane separating said first zone from a second zone; said process comprising: (I) feeding through said reactor inlet a feed comprising hydrogen and at least one C1-4 alkane so that the alkane and hydrogen contact said dehydrogenation catalyst in said first zone; (II) operating said reactor at a temperature and pressure sufficient to allow dehydrogenation of said alkane and formation of said aromatic hydrocarbon and hydrogen; (III) allowing hydrogen to selectively pass through the hydrogen transport membrane into said second zone; (IV) removing hydrogen from said second zone in order to establish a concentration gradient for hydrogen transport between said first and second zones; (V) recovering through the reactor outlet in the first zone, said aromatic hydrocarbon along with any unreacted C1-4 alkane and any hydrogen not extracted through the hydrogen transport membrane; (VI) separating the aromatic hydrocarbon from the hydrogen and any unreacted C1-4 alkane; and (VII) recycling at least a part of said hydrogen and any unreacted C1-4 alkane to the feed using a recycle stream, wherein the hydrogen content of the recycle stream is 3 to 35 vol %.

16. The process of claim 15, wherein the step (IV) of removing hydrogen from said second zone in order to establish a concentration gradient for hydrogen transport between said first and second zones is achieved by sweeping with an oxygen containing gas to convert at least some of the hydrogen to water.

17. The process of claim 16, wherein oxygen is derived from said oxygen containing gas.

Description

BRIEF DESCRIPTION OF THE DRAWINGS

(1) FIG. 1 is a flow diagram showing an outline of a process according to the present invention. Alkanes (A) and hydrogen are fed via inlet (4) into fluidised or fixed bed reactor (1) containing a dehydrogenation catalyst bed (5). Hydrogen passes through tubular transport membrane (2) leaving formed aromatic compounds (B), unreacted hydrogen and starting material exiting at outlet (3). Components A, B and hydrogen are separated in separating apparatus (6) and component A and hydrogen recycled via line (7) back to the initial feedstock. Component (B) is taken for optional further purification and use.

(2) FIG. 2 illustrates examples of the different types of solid oxide membrane assemblies for the hydrogen transport membrane used in the present invention.

(3) FIG. 3 is a drawing of the catalytic membrane reactor used in Example 1. Fed gas passes up the outside walls of the membrane and through catalyst bed (5). Hydrogen is transported across membrane (2) into the central tube of the reactor where it is removed. Outlet gases are those that pass through catalyst bed (5).

(4) FIG. 4 shows thermodynamic equilibrium selectivity to benzene relative to other C1-C10 hydrocarbons for the dehydroaromatization of methane at different hydrogen to methane ratios at temperatures 650 C., 700 C. and 750 C. at pressures 1 and 3 bara at 700 C.

(5) FIG. 5 shows the formation rate of benzene versus time on stream at 700 C., 3000 mL/g h and 3 bara in a catalytic membrane reactor and a conventional fixed-bed reactor.

EXAMPLE 1

(6) Membrane Preparation

(7) A tubular asymmetric membrane of 65 wt. % NiBaZr.sub.0.7Ce.sub.0.2Y.sub.0.1O.sub.3- (BZCY72) support with a 30 m dense membrane was synthesized using a reactive sintering approach (Coors et al. Journal of Membrane Science 376 (2011) 50-55). Precursors of BaSO.sub.4, ZrO.sub.2, Y.sub.2O.sub.3 and CeO.sub.2 was mixed in stoichiometric amounts (metal basis) together with 1 wt. % NiO in a Nalgene bottle on a jar roller for 24 h. The material was dried in air and sieved through a 40 mesh screen. A portion of the mix was mixed additionally with 64 wt. % NiO and blended with water soluble acrylic and cellulosic ether plasticizer to prepare the extrusion batch. Green tubes were extruded using a Loomis extruder. The extruded tubes were then dried and spray coated with the mixture with a 1 wt. % NiO solution. The tubes was the co-fired by hang-firing in air at 1500 C. for 4 h. The sintered tubes were then treated in a hydrogen mixture (safe gas) at 1000 C. to reduce the NiO to Ni and give the necessary porosity in support structure. Dimensions of the half-cell tube are 25 cm long with an outer diameter of 10 mm, an inner diameter of 9.8 and a membrane thickness of 30 m.

(8) A second outer electrode was prepared by mixing 50 vol. % of Cu-metal in BCZY72 pre-calcined powder. An ink was prepared using polyvinylpyrrolidine (PVP) and 50 vol. % ethylene glycol in isopropanol and coated on the reduced tube using a brush. The electrode was fired under reducing conditions (5% H.sub.2 in Ar) at 1000 C. for 2 h. A Cu wire was used as current wires for the outer electrode and Pt wire for current wires for the inner electrode.

Catalyst Preparation

EXAMPLE 2

(9) A commercial ZSM-5 powder was acquired by Zeolyst, batch CBV 2314 with a product number 2493-39. The catalyst was supplied in ammonia form (NH.sub.4-ZSM-5) with a nominal SiO.sub.2/Al.sub.2O.sub.3 ratio of 23. The catalyst is converted to the protonated form, H-ZSM-5 by calcination in air at 550 C. for 6 hours. Mo was introduced into the catalyst by incipient wetness impregnation using a solution of ammonium heptamolybdate tetrahydrate (NH.sub.4).sub.6Mo.sub.7O.sub.24.4H.sub.2O in water, drying at 100 C. over night and then calcined at 500 C. for 18 h. Prior to catalytic membrane reactor testing the impregnated catalyst was pressed to pellets, crushed in a mortar and sieved to 250-420 m to yield a suitable particle size for the reactor dimensions.

EXAMPLE 3

(10) An MCM-22 zeolite was synthesized using hexamethylenimine as a template with a nominal SiO.sub.2/Al.sub.2O.sub.3 ratio of 30. The catalyst was converted to the protonated form, H-MCM-22, by ion-exchange with 0.1M NH.sub.4NO.sub.3 and following calcination in air at 550 C. for 6 hours. Mo/H-MCM-22 catalyst was prepared as described in Example 2.

(11) Reactor

(12) A fixed bed type reactor is utilized where the BZCY membrane is located in the centre of the fixed bed with the catalyst placed around the membrane. A sketch of the reactor is given in FIG. 4. The volume of the bed is given by the distance between the tube and the fixed bed wall. The feed flow is directed upwards through the catalyst bed along the membrane. An alumina single bore tube is placed on the inside of the membrane tube to supply sweep gas to the top of tube so that there is an even flow of sweep gas to the reactor inner tube outlet. A 3-zone furnace is used to have a uniform heat distribution of the active reactor length. It is important that the product gas is kept above 220 C. to avoid condensation of naphthalene in the gas lines.

(13) Process

(14) The reactor was operated at 700 C. using the reactor and membrane electrode assembly as described above. In total 2 g of 6 wt % Mo impregnated ZSM-5 catalyst was used in the experiment. The active area of the tube was 12.5 cm.sup.2 and the catalyst was diluted with SiC to ensure that the catalyst covered the whole active area. The Cu and Pt wires, forming the outer and inner electrodes respectively, were connected to a Gamry Reference 3000 Potentiostate/Galvanostat/ZRA utilized in galvanostatic mode. Carboratization of the catalyst was done in-situ after heating in Ar with 1% H.sub.2 to 700 C. and then switching to CH.sub.4. During operation the GHSV (gas hourly space velocity) was held at 1800 mL/h g, which yields a CH.sub.4 flow of 60 mL/min. To simulate different hydrogen content in the gas feed a mass flow controller was utilized. For the sweep side a humidified (3% H.sub.2O) 5% H.sub.2 in Ar mixture was used with a flow of 50 mL/min. Both the product and the sweep side was monitored using a Agilent 7890B GC equipped with 2 MolSieve, 2 porapak and a HP-plot-q column and using a FID and TCD detector for the product stream and a TCD detector for the sweep stream. This enables quantification of the hydrogen and hydrocarbon content in the product stream and the hydrogen content in the sweep stream.

(15) In table 1 the hydrogen concentration for the feed, sweep exhaust and product exhaust is given. In addition the benzene yield is presented.

(16) TABLE-US-00001 TABLE 1 Hydrogen concentration in feed, sweep exhaust and product exhaust and aromatics yield in an experiment with GHSV of 1800 mL/gh at 700 C. at 1 bara. Sweep Product Feed H.sub.2 exhaust H.sub.2 exhaust Aromatics Benzene mL/min mL/min H.sub.2 mL/min yield/% selectivity/% 7 1.63 9.55 2.72 68 9 1.56 10.44 2.40 69 11 1.17 11.25 1.43 70 12 1.10 10.62 1.59 70

(17) Results given in table 1 shows that with in-situ hydrogen removal it is possible to adjust the hydrogen concentration in the product exhaust. This can be made lower or the same as the feed hydrogen concentration. Coke formation decreases and catalyst lifetime increases, with increasing hydrogen partial pressure as expected from thermodynamic calculations. It is also shown that the aromatic yield decreases with increasing hydrogen feed, which implies a decrease in the overall conversion due to the shift in the equilibrium towards methane. It can also be seen that the selectivity towards benzene increases with increasing hydrogen feed, again as expected from thermodynamic calculations given in FIG. 5.

(18) The preferred operation principle is dependent on the overall process. If the goal is to achieve high yield, low hydrogen feed and/or a high level of hydrogen extraction across the hydrogen transport membrane is favoured. If the process is continuous and includes an efficient recycling loop the co-feed should be adjusted to optimize total process economics, taking into account also the cost/value balance of the recycle gas. It must be appreciated that this laboratory example is not optimised and on larger scale yields would be expected to be higher as more hydrogen can be removed via the membrane.

(19) The following examples are based on the protocols above varied as described in each example:

EXAMPLE 4

(20) A process operated at 700 C., 1500 mL/h g and 1 bara using simulated recycle feed of 5 vol % H.sub.2, Mo impregnated ZSM-5 zeolite and membrane electrode assembly as described above.

EXAMPLE 5

(21) A process operated at 700 C., 1500 mL/h g and 1 bara using simulated recycle feed of 10 vol % H.sub.2, Mo impregnated ZSM-5 zeolite and membrane electrode assembly as described above.

EXAMPLE 6

(22) A process operated at 700 C., 1500 mL/h g and 1 bara using simulated recycle feed of 15 vol % H.sub.2, Mo impregnated ZSM-5 zeolite and membrane electrode assembly as described above.

EXAMPLE 7

(23) A process operated at 700 C., 1500 mL/h g and 1 bara using simulated recycle feed of 5 vol % H.sub.2, Mo impregnated MCM-22 zeolite and membrane electrode assembly as described above.

EXAMPLE 8

(24) A process operated at 700 C., 1500 mL/h g and 1 bara using simulated recycle feed of 10 vol % H.sub.2, Mo impregnated MCM-22 zeolite and membrane electrode assembly as described above.

(25) TABLE-US-00002 TABLE 3 Feed composition and products formation rate in nmol/g s from Examples 4 to 8 after 3 h on stream. Product exhaust Feed Alkanes H.sub.2/ CH.sub.4/ H.sub.2/ (C1-C2)/ Benzene/ Coke/ Example nmol/g s nmol/g s nmol/g s nmol/g s nmol/g s nmol/g s 4 845 16380 2751 13642 900 600 5 1704 20108 3170 17497 557 1022 6 2535 20108 3604 18583 320 981 7 845 16380 2234 15556 1174 445 8 1704 20108 2241 16234 1075 357

(26) TABLE-US-00003 TABLE 4 Products formation rate in nmol/g s from Example 8 with and without hydrogen removal. Without With hydrogen removal hydrogen removal CO [nmol g.sup.1 s.sup.1] 0 63 Benzene [nmol g.sup.1 s.sup.1 ] 810 1075 Coke [nmol g.sup.1 s.sup.1] 398 357

EXAMPLE 9

(27) A process operated at 700 C., 3000 mL/h g and 3 bara using simulated recycle feed of 5 vol % H.sub.2, Mo impregnated MCM-22 zeolite and membrane electrode assembly as described above. FIG. 5 shows the benzene formation rate in nmol/s g versus time on stream for this process (catalytic membrane reactor). It is also included for comparison the obtained benzene formation in a conventional fixed-bed reactor without hydrogen removal.

EXAMPLE 10

(28) A process operated at 725 C., 3000 mL/h g and 1 bara using simulated recycle feed of 5 vol % H.sub.2, Mo impregnated MCM-22 zeolite and membrane electrode assembly as described above.

EXAMPLE 11

(29) A process operated at 725 C., 3000 mL/h g and 3 bara using simulated recycle feed of 5 vol % H.sub.2, Mo impregnated MCM-22 zeolite and membrane electrode assembly as described above.

(30) TABLE-US-00004 TABLE 5 Feed composition and products formation rate in nmol/g s from Examples 9 to 11 after 6 h on stream. Product exhaust Feed Alkanes H.sub.2/ CH.sub.4/ H.sub.2/ (C1-C2)/ Benzene/ Coke/ Ex nmol/g s nmol/g s nmol/g s nmol/g s nmol/g s nmol/g s 9 1745 32950 3287 27409 1552 1167 10 1745 32950 4406 32391 1540 924 11 1745 32950 4385 27290 1588 1025

EXAMPLE 12

(31) A process operated at 700 C., 1500 mL/h g and 1 bara with Mo impregnated MCM-22 zeolite and membrane electrode assembly as described above.

(32) TABLE-US-00005 TABLE 6 Rate of formation of benzene and hydrogen removal in an experiment with GHSV of 1500 mL/gh at 700 C. and 1 bara while using wet and dry sweep during hydrogen removal. Without With hydrogen hydrogen removal removal Dry sweep Wet sweep Benzene [nmol g.sup.1 s.sup.1] 850 1027 1174 Hydrogen removal [nmol g.sup.1 s.sup.1] 0 136 1226