METHOD FOR LOW HYDROGEN CONTENT SEPARATION FROM A NATURAL GAS MIXTURE

Abstract

A method for low hydrogen content separation from a natural gas mixture includes the following steps: a) providing a stream having hydrogen; b) transferring the stream having hydrogen of a) as an inlet stream to a first membrane unit for obtaining a retentate and a permeate, wherein the molar fraction of hydrogen in the permeate is higher that the molar fraction of hydrogen in the retentate, c) transferring the retentate to an electrochemical hydrogen compressor (EHC) for further hydrogen separation and purification.

Claims

1. A method for low hydrogen content separation from a natural gas mixture, the method including the following steps: a) providing a stream comprising hydrogen, b) transferring the stream comprising hydrogen of a) as an inlet stream to a first membrane unit for obtaining a retentate and a permeate, wherein the molar fraction of hydrogen in the permeate is higher than the molar fraction of hydrogen in the retentate, and c) transferring the retentate to an electrochemical hydrogen compressor (EHC) for further hydrogen separation and purification.

2. The method according to claim 1, wherein the method further includes step b1), wherein the permeate of step b) is transferred as an inlet stream to a second membrane unit, in which second membrane unit a second retentate and a second permeate is produced, wherein the molar fraction of hydrogen in the second permeate is higher than the molar fraction of hydrogen in the second retentate, the second retentate is sent back as an inlet stream to the membrane unit of step b).

3. The method according to claim 1, wherein the inlet stream comprising hydrogen is heated in a heat exchanger to the operation temperature of the first or second membrane unit before transferring the inlet stream comprising hydrogen to the first or second membrane unit.

4. The method according to claim 1, wherein the retentate obtained in step b) is cooled down in a heat exchanger to the operation temperature of the electrochemical hydrogen compressor (EHC) before transferring the retentate to the electrochemical hydrogen compressor (EHC).

5. The method according to claim 1, wherein a vacuum unit is used for increasing the driving force via the first and/or second membrane unit.

6. The method according to claim 2, wherein the second retentate stream originating from the second membrane unit is heated in a heat exchanger before transferring the second retentate stream to the inlet of the first membrane unit.

7. The method according to claim 1, wherein the first membrane unit is chosen from the group of Pd-based ceramic supported membrane and Pd-based metallic supported membrane.

8. The method according to claim 1, wherein the inlet pressure of the stream comprising hydrogen of a) is at least 5 bara.

9. The method according to claim 1, wherein the permeate pressure of the first membrane unit is lower than 130 mbar.

10. The method according to claim 1, wherein the hydrogen concentration of the stream comprising hydrogen of a) is at least 10 vol. %.

11. The method according to claim 1, wherein the permeate pressure of the first membrane unit is lower than 5 bar.

12. The method according to claim 1, wherein the retentate pressure of the second membrane unit is lower than 10 bar.

13. An apparatus for low hydrogen content separation from a natural gas mixture, the apparatus comprising: a first membrane unit having an inlet for a stream comprising hydrogen, an outlet for retentate and an outlet for permeate, wherein the molar fraction of hydrogen in the permeate is higher than the molar fraction of hydrogen in the retentate; and an electrochemical hydrogen compressor (EHC) having an inlet for the retentate, an outlet cathode site and an outlet anode site, wherein the molar fraction of hydrogen in the outlet cathode site is higher than the molar fraction of hydrogen in the outlet anode site.

14. The apparatus according to claim 13, the apparatus further comprising a second membrane unit, the second membrane unit having an inlet for a stream comprising hydrogen, an outlet for second retentate and an outlet for second permeate, wherein the molar fraction of hydrogen in the second permeate is higher than the molar fraction of hydrogen in the second retentate, wherein the outlet for retentate of the first membrane unit is connected to the inlet of the second membrane unit.

Description

BRIEF DESCRIPTION OF THE DRAWINGS

[0033] FIG. 1 is a process flow diagram for Embodiment 1 of the present disclosure;

[0034] FIG. 2 is a process flow diagram for Embodiment 2 of the present disclosure; and

[0035] FIG. 3 is a process flow diagram for Embodiment 3 of the present disclosure.

DETAILED DESCRIPTION OF THE DRAWINGS

[0036] In order to make the technicians of this field better understand the present disclosure, the technical schemes in the embodiments of the present disclosure will be clearly and completely described by combining with the drawings in the embodiments of the present disclosure below. The term module is used to describe a complete unit composed of the membranes, the pressure support structure, the feed inlet, the outlet permeate and retentate streams, and an overall support structure. The flow that passes the membrane is called permeate. The materials rejected by the membrane are called retentate. In the present description the molar fraction of hydrogen in the permeate is higher than the molar fraction of hydrogen in the retentate.

[0037] Three different embodiments are proposed for hydrogen separation and purification from a 10% H.sub.2 and 90% CH.sub.4 mixture by combining palladium membrane, vacuum pump, mechanical compressor and electrochemical hydrogen compressor technologies. The reference for all the different embodiments is a production of 25 kgH.sub.2/day; all the membrane surface area and feed rate were fixed based on this production.

[0038] FIG. 1 shows a process flow diagram 10 for Embodiment 1 of the present disclosure. An incoming stream 1 with a total flow rate of 6246.1 mol/h (coming from the grid) is initially pre-heated in a heat exchanger 2 using the outlet retentate stream 6 (while it is being cooled down), and then a heater 3 supplies the additional heat needed to reach the membrane working temperature of 400° C. After heater 3, stream 4 is sent to a membrane module 5, which has a surface area of 1.2 m.sup.2. The selected membrane is Pd-based ceramic supported with a hydrogen permeance of 2.2*10.sup.6 mol/s/m.sup.2/Pa at 400° C. and an ideal perm-selectivity, defined as the ratio between H.sub.2 and CH4 permeance at 1 bar pressure difference, of 20000 (obtained experimentally). The retentate side of the membrane is at 8 bara (same as the stream coming from the grid), while the permeate side was kept at 100 mbara by using vacuum pump 9 (to increase the driving force via the membrane). The permeate stream 7 is cooled down with a cooler 8 before entering a vacuum pump 9 (the maximum inlet temperature of the vacuum pump is 75° C.). The retentate stream 6 of the membrane module 5 is then sent to heat exchanger 2 and the thus cooled downstream 12 is sent to an electrochemical hydrogen compressor 13 (EHC) with a protonic membrane resistance of 6 mΩ and 350 cells in parallel with working temperature of 65° C. for further hydrogen separation and purification. The retentate outlet stream 6 is cooled down in the heat exchanger 2 previously mentioned while heating stream 1 coming from the grid. In addition, hydrogen 11 coming from the outlet cathode side of electrochemical hydrogen compressor 13 is kept at 8 bara to make sure the proper protonic membrane humidification which otherwise would be difficult to obtain at atmospheric pressure. Purified hydrogen 15 from the permeate side of membrane module 5 and hydrogen 11 from the cathode side of the electrochemical hydrogen compressor 13 are then mixed as stream 16, after depressurization of stream 11 from electrochemical hydrogen compressor 13. The outlet anode side stream 14 of the electrochemical hydrogen compressor 13 is sent back to the grid with no or low hydrogen concentration in stream 14 (depends on which configuration is chosen).

[0039] FIG. 2 shows a process flow diagram 20 for Embodiment 2 of the present disclosure, which combines two membrane modules, carbon molecular sieve membrane (CMSM) and Pd-based ceramic supported membrane, vacuum pump and EHC. This configuration is proposed for the high-pressure gas grid (approx. 40-80 bar) which allows the connections between the two membrane modules without any mechanical compressor in between. The considered grid pressure is 40 bar with a total feed rate similar to Embodiment 1. Feed 1 is initially heated in heat exchanger 2 to reach the operating temperature of 70° C. Then after heat exchanger 2 and heater 3, a heated stream 4 is sent to a first membrane module 5 with a surface area of 5.02 m.sup.2, which is a large surface area because the hydrogen permeance of CMSM is 7-10.sup.8 mol/s/m.sup.2/Pa at 70° C. with an ideal selectivity of 550 at 40 bar. The permeate side is kept at 3 bar to give enough driving force to membrane module 5 and keep a pressure difference for second membrane module 17. The permeate side 7 is heated in heat exchanger 8 and further heated in heater 16 and the temperature of stream 18 thus heated is about 400° C., which is the operating temperature of second membrane module 17 comprising a Pd-based membrane. The selected membrane is a ceramic supported Pd-based with a hydrogen permeance of 2.2-1 O.sup.6 mol/s/m.sup.2/Pa at 400° C. and 1 bar pressure difference and an ideal selectivity of 20000. A membrane area of 0.62 m.sup.2 is required to achieve a final separation of 25 kg/day. Permeate side 20 of second membrane module 17 (using vacuum pump 22) was kept same as Embodiment 1 by using a heater 21. Outlet retentate 19 of second membrane module 17 is recycled and used as an inlet stream for first membrane module 5. Outlet retentate 6 of first membrane module 5 is heated in heat exchanger 2 and then sent to electrochemical hydrogen compressor 13 (EHC) to further purify the hydrogen. After mixing a stream of hydrogen 11 separated from electrochemical hydrogen compressor 13 (EHC) and a stream of hydrogen 23 coming from second membrane module 17 comprising a Pd-based membrane, mixed stream 24 is sent to the end user while retentate stream 14 is fed back to the grid.

[0040] FIG. 3 shows a process flow diagram 30 for Embodiment 3 of the present disclosure in which process flow diagram two membrane modules with a mechanical compressor in between, a vacuum pump and an EHC have been combined. The first membrane module 5 is a Pd-based ceramic supported membrane with an ultra-thin (1-3 pm) palladium layer which allows high hydrogen permeance with a lower selectivity compared to the membranes adopted for Embodiment 1 and 2. The hydrogen permeance and the ideal perm-selectivity were 4*1 O.sup.6 mol/s/m.sup.2/Pa and 4000 respectively. The second membrane module 17, which is connected in series to the previous one, is a Pd-based double-skin membrane with a hydrogen permeance of 2*1 O.sup.6 molls/m.sup.2/Pa and an ideal perm-selectivity of 60000. The second membrane module 17 is mainly responsible for increasing the hydrogen purity. Stream 1 which comes from the grid has the same feed rate and composition as Embodiment 1 and 2 and is also initially heated in heat exchanger 2 and heater 3 where outlet retentate stream 12 of the first membrane module 5 is cooled down to 65° C., which is the working temperature of electrochemical hydrogen compressor 13 (EHC). Outlet anode side 14 of electrochemical hydrogen compressor 13 (EHC) is sent back to the natural gas grid. After heater 3, the stream is sent to first membrane module 5 with a surface area of 1.2 m2, while the permeate side was kept at a pressure of 100 mbara which was achieved using vacuum pump 20. Permeate stream 7 from first membrane module 7 is sent to heat exchanger 30 and stream 25 is sent to vacuum pump 20. Outlet stream 27 of vacuum pump 20 was then compressed to 8 bara in compressor 28 and then sent as stream 29 to second membrane module 17 with a surface area of 0.15 m.sup.2 after two cascaded heat exchangers 30, 31. The thus obtained stream 32 is a feed stream for second membrane module 17. Retentate stream 33 from second membrane module 17 (mainly some impurities and remaining inextricable H.sub.2) is sent back to first membrane module 5 for further purification. A stream of purified hydrogen 36 from second membrane module 17 and a stream 11 from the cathode side of electrochemical hydrogen compressor 13 (EHC) is then mixed as stream 37 and sent to the end users.

[0041] The inventors calculated the performance of different configurations. In the present description configuration A refers to Embodiment 1, configuration B to Embodiment 2 and configuration C to Embodiment 3. The results are shown in Table 1.

[0042] From configuration A, which includes a ceramic supported Pd-based membrane connected to a vacuum pump and an electrochemical hydrogen compressor, it was possible to recover 83.39% with a purity of 99.93%. The Pd-based membrane, with a surface area of 1.62 m.sup.2, recovers 328.9 mol/h of H.sub.2 and 0.3 mol/h of CH.sub.4. The concentration polarization in the retentate side plays a role in terms of hydrogen driving force, i.e. the higher the retentate pressure, the higher the mass transfer limitation between the bulk and the palladium surface. The retentate side was then sent to the electrochemical hydrogen compressor (EHC), where an extra 191.9 mol/h of hydrogen with a purity of 100% is separated. By varying the applied voltage, it is possible to change the hydrogen recovery from the EHC (the efficiency of the EHP was considered to be 60% (optimal value for the energy consumption), while the voltage was changed consequently).

[0043] Configuration B guarantees very high hydrogen purity (99.99%) compared to configuration A thanks to the further purification achieved with the second membrane module. The membrane is responsible for separating 52.66% of hydrogen with a surface area of 6.32 m.sup.2 and a purity of 91.61%. The final purity reached is higher than configuration A because it is easier to further purify the stream when contains 91.61% of hydrogen. The electric consumption required is associated only to the heat required to reach the working temperature of the system and was 5.62 kWh/kgH.sub.2. The total hydrogen production separated in this configuration was 25 kg/day with only 1.81% of hydrogen is sent back to the grid, Therefore, with configuration B, it is even possible to produce high purity (99.99) hydrogen with power consumptions lower than 6 kWh/kg H.sub.2.

[0044] On the other hand, configuration C gives higher hydrogen purity compared to the configurations A and B, but the energy consumption (7.95 kWh/kgH.sub.2) required was the highest.

[0045] In configuration “A1”, the type of membrane selected is a Pd based metallic supported membrane, which has lower hydrogen permeance but higher perm selectivity compared to the ceramic supported membrane. For this reason, the membrane area required to separate 25 kg{circumflex over ( )}/day increases to 2.92 m.sup.2 in the current configuration while the purity raises to 99.99%. The energy consumption is similar to configuration “A”, lower than “B” and “C”. Furthermore, the retentate pressure was varied for a proper understanding of the HRF and purity.

[0046] Configuration “A2” is based on an inlet pressure of 15 bara coming from the natural gas grid instead of 8 bara. From the results it is possible to notice that by increasing the retentate pressure, lower surface are is needed (1.43 m.sup.2), compared to the master configuration “A” (1.62 m.sup.2) but the purity drops from 99.93% to 99.88% due to higher driving force for the contaminant gas to permeate through. Lower membrane surface area is required because of the larger driving force along the membrane thanks to higher pressure difference; the energy consumption is very similar to master configuration “A”.

[0047] In configuration “A3” the permeate pressure is changed from 100 mbar to 70 mbar to verify the influence of a different vacuum on the performance of the overall system in terms of purity. To reach the same HRF the membrane surface area is reduced from 1.62 m.sup.2 for the master configuration “A” to 1.54 m.sup.2 with a vacuum of 70 mbar.

[0048] Configuration “A4” considers a H.sub.2 concentration from the natural gas grid of 15% instead of 10% like the previous cases, which results in a reduction of the membrane area from 1.62 to 1.56 m.sup.2 and a slightly higher final purity (99.96%) compared to the master configuration (99.93%). According to the inventors this effect is related to the higher contaminant driving force (higher methane concentration at the inlet).

[0049] Configuration “A5” is based on a lower total flow rate from the grid: 1784.6 mol/h which 10% is H.sub.2 and 90% CH.sub.4. It is possible to increase the HRF of the membrane from 48.80% to 79.67% and from a total HRF of 79.52% to 91.86% with a purity of 99.77%. The hydrogen purity decreases for a combination of two different reasons. The first one is related to the lower hydrogen separated from the EHP, which could guarantee a purity of 100%, while the second reason is the higher mass transfer limitation occurring at lower flow rate. The main advantage of configuration “A5” relies on the better quality of the natural gas grid due to the lower hydrogen concentration going back to the natural gas grid (0.90%).

[0050] The aim of configuration “A6” is to reduce the hydrogen flow rate separated by the EHP, at the expense of the membrane surface and final separation cost, to reduce the energy consumption. The membrane area to keep the same HRF is increased to 2.41 m.sup.2, with a reduction of the energy consumption from 5.19 of configuration “A”, to 4.72 kWh/kgH.sub.2.

[0051] Configuration “B1” differs from configuration “B” because of the type of membrane employed. In configuration B, a carbon molecular sieve membrane with a perm-selectivity of 550 was considered, while in case “B1”, an ultra-thin Pd based ceramic supported membrane with a selectivity of 5000 is adopted in the first membrane module. The energy consumption raises due to the higher operating temperature of Pd-based membrane (400° C.) in contrast to carbon molecular sieve membrane (CMSM). The energy consumption is 6.03 instead of 5.62 kWh/kgH.sub.2. Even if in case “B1” a lower surface area is required, due to the larger membrane costs and higher energy consumption, it results less economic convenient to adopt Pd-based membrane in the first membrane module. On the other hand, carbon molecular sieve membrane seems to be promising especially for separating hydrogen at high pressure grid.

[0052] Configuration “B2” is based on configuration “B” with the main difference of lower permeate pressure of the first membrane module (2 bar instead of 3). The surface area of the first module is reduced thanks to the more relevant driving force, on the other hand, the membrane surface of the second module raises because the retentate inlet pressure decreased.

[0053] Configuration “C” can guarantee relatively high purity because of the two membrane modules in series which assure a substantial purification of the stream. Configuration “C1” is based on a lower retentate pressure of the second membrane module, which is 4 bara instead of 8. In this configuration, the membrane area is reduced from 2.42 to 1.87 m.sup.2 with a decrease in energy consumption from 7.95 to 6.38 kWh/kgH.sub.2.