Membrane CO2 separation process

Abstract

Described herein are membrane processes for separating CO.sub.2 from flue gas. An exemplary process involves passing a fluid stream including the flue gas across a membrane permeable to CO.sub.2 and H.sub.2O, removing treated gas from a feed side of the membrane that has less CO.sub.2 than the flue gas, and removing permeate from a permeate side of the membrane comprising CO.sub.2 and H.sub.2O. Suitably, the permeate is removed at a sub-atmospheric vacuum pressure. The permeate is then cooled to remove at least some of the H.sub.2O from the permeate and form a smaller volume of H.sub.2O-depleted, CO.sub.2 enriched permeate.

Claims

1. A membrane process to separate CO.sub.2 from flue gas, the process comprising: passing a fluid stream including the flue gas across a membrane permeable to CO.sub.2 and H.sub.2O; removing treated gas from a feed side of the membrane, the treated gas having less CO.sub.2 than the flue gas; removing permeate from a permeate side of the membrane at a sub atmospheric pressure between 0.1 to 0.4 bar, the permeate comprising CO.sub.2 and H.sub.2O; cooling the permeate to remove at least some of the H.sub.2O from the permeate and form a smaller volume of H.sub.2O-depleted, CO.sub.2 enriched permeate; and using a vacuum pump to increase the gas pressure of the smaller volume of H.sub.2O-depleted, CO.sub.2 enriched permeate to at least about atmospheric pressure.

2. The membrane process as set forth in claim 1, wherein the fluid stream passed across the membrane contains at least 70% of its saturation concentration of water.

3. The membrane process as set forth in either of claim 1 or claim 2, wherein the temperature difference between the fluid stream passed across the membrane and the smaller volume of H.sub.2O-depleted CO.sub.2-enriched permeate is at least 40° C.

4. The membrane process as set forth in claim 1, further comprising adding H.sub.2O to the flue gas before passing the fluid stream across the membrane such that the fluid stream that is passed across the membrane includes the flue gas and the water.

5. The membrane process as set forth in claim 4, wherein said adding H.sub.2O to the flue gas comprises using a direct contact cooler to add H.sub.2O to the flue gas before passing the fluid stream including the flue gas through the membrane.

6. The membrane process as set forth in claim 5, wherein the direct contact cooler adjusts a temperature of the flue gas.

7. The membrane process as set forth in claim 1, further comprising bringing the fluid stream including the flue gas to a pressure of 0.8 to 1.5 bar before passing the fluid stream across the membrane.

8. The membrane process as set forth in claim 1, further comprising bringing the fluid stream including the flue gas to a temperature of greater than 50° C. before passing the fluid stream across the membrane.

9. The membrane process as set forth in claim 1, further comprising bringing the fluid stream including the flue gas to a temperature in range of from about 50° C. to about 80° C. before passing the fluid stream across the membrane.

10. The membrane process as set forth in claim 1, further comprising bringing the fluid stream including the flue gas to have greater than 10 mol % water vapor.

11. The membrane processes as set forth in claim 1, comprising bringing the fluid stream including the flue gas to have greater than 70% of its saturation water concentration.

12. The membrane process as set forth in claim 1, wherein the membrane has an H.sub.2O/CO.sub.2 selectivity of greater than 2, measured at the operating conditions of the membrane process.

13. The membrane process as set forth in claim 1, wherein the membrane has a CO.sub.2/N.sub.2 selectivity of greater than 10, measured at the operating conditions of the membrane process.

14. The membrane process as set forth in claim 1, wherein the membrane process removes at least 50% of the CO.sub.2 of the flue gas.

15. The membrane process as set forth in claim 1, wherein the membrane process removes about 50% to about 80% of the CO.sub.2 of the flue gas.

16. The membrane process as set forth in claim 1, wherein said cooling the permeate comprises cooling the permeate to a temperature in an inclusive range of from about 5° C. to about 30° C.

17. The membrane process as set forth in claim 1, wherein the smaller volume of H.sub.2O-depleted, CO.sub.2 enriched permeate has a CO.sub.2 concentration of greater than 35%.

18. The membrane process as set forth in claim 1, wherein the fluid stream passed across the membrane contains at least 70% of its saturation concentration of water and wherein the temperature difference between the fluid stream passed across the membrane and the smaller volume of H.sub.2O-depleted, CO.sub.2-enriched permeate is at least 30° C.

19. The membrane process as set forth in claim 18, wherein the fluid stream passed across the membrane contains at least 80% of its saturation concentration of water.

20. The membrane process as set forth in claim 19, wherein the temperature difference between the fluid stream passed across the membrane and the smaller volume of H.sub.2O-depleted CO.sub.2-enriched permeate is at least 40° C.

21. The membrane process as set forth in claim 18, further comprising adding H.sub.2O to the flue gas before passing the fluid stream across the membrane such that the fluid stream that is passed across the membrane includes the flue gas and the water.

22. The membrane process as set forth in claim 21, wherein said adding H.sub.2O to the flue gas comprises using a direct contact cooler to add H.sub.2O to the flue gas before passing the fluid stream including the flue gas across the membrane.

23. The membrane process as set forth in claim 22, wherein the direct contact cooler adjusts a temperature of the flue gas.

24. The membrane process as set forth in claim 18, further comprising bringing the fluid stream including the flue gas to a pressure of 0.8 to 1.5 bar before passing the fluid stream across the membrane.

25. The membrane process as set forth in claim 18, further comprising bringing the fluid stream including the flue gas to a temperature of greater than 50° C. before passing the fluid stream across the membrane.

26. The membrane process as set forth in claim 18, further comprising bringing the fluid stream including the flue gas to a temperature in a range of from about 50° C. to about 80° C. before passing the fluid stream across the membrane.

27. The membrane process as set forth in claim 18, further comprising bringing the fluid stream including the flue gas to have greater than 10 mol % water vapor.

28. The membrane process as set forth in claim 18, wherein the membrane has an H.sub.2O/CO.sub.2 selectivity of greater than 2, measured at the operating conditions of the membrane process.

29. The membrane process as set forth in claim 18, wherein the membrane has a CO.sub.2/N.sub.2 selectivity of greater than 10, measured at the operating conditions of the membrane process.

30. The membrane process as set forth in claim 18, wherein the membrane process removes at least 50% of the CO.sub.2 of the flue gas.

31. The membrane process as set forth in claim 18, wherein the membrane process removes about 50% to about 80% of the CO.sub.2 of the flue gas.

32. The membrane process as set forth in claim 18, wherein said cooling the permeate comprises cooling the permeate to a temperature in an inclusive range of from about 5° C. to about 30° C.

33. The membrane process as set forth in claim 18, wherein the smaller volume of H.sub.2O-depleted, CO.sub.2 enriched permeate has a CO.sub.2 concentration of greater than 35%.

34. The membrane process as set forth in claim 18, wherein the fluid stream passed across the membrane has a temperature of at least 50° C. and no more than 90° C., has greater than 10 mol % water vapor, and has a gas pressure less than or equal to 2 bar.

35. The membrane process as set forth in claim 34, wherein the fluid stream passed across the membrane has a temperature of no more than 80° C.

36. The membrane process as set forth in claim 35, wherein the fluid stream passed across the membrane has a temperature of no more than 70° C.

37. The membrane process as set forth in claim 34, wherein the fluid stream passed across the membrane has greater than 15 mol % water vapor.

38. The membrane process as set forth in claim 34, wherein the fluid stream passed across the membrane has greater than 25 mol % water vapor.

39. The membrane process as set forth in claim 35, wherein the gas pressure of the fluid stream passed across the membrane is less than 1.5 bar.

40. The membrane process as set forth in claim 34, wherein the temperature difference between the fluid stream passed across the membrane and the smaller volume of H.sub.2O-depleted CO.sub.2-enriched permeate is at least 40° C.

41. The membrane process as set forth in claim 40, wherein the fluid stream passed across the membrane has a temperature of no more than 80° C., wherein the fluid stream passed across the membrane has greater than 15 mol % water vapor, and wherein the gas pressure of the fluid stream passed across the membrane is less than 1.5 bar.

42. The membrane process of claim 41, further comprising pretreating the flue gas to obtain the fluid stream passed across the membrane.

43. The membrane process of claim 34, further comprising pretreating the flue gas to obtain the fluid stream passed across the membrane.

44. The membrane process of claim 18, further comprising pretreating the flue gas to obtain the fluid stream passed across the membrane.

45. A membrane process to separate CO.sub.2 from flue gas comprising: (i) bringing the flue gas to a pressure of 0.8 to 1.5 bar and a temperature of greater than 50° C., wherein the flue gas contains greater than 10 mol % water vapor; (ii) passing the flue gas from (i) across a membrane permeable to water and CO.sub.2, said membrane having an H.sub.2O/CO.sub.2 selectivity of greater than 2 and a CO.sub.2/N.sub.2 selectivity of greater than 10, measured at the operating conditions of the membrane process; (iii) removing from the feed side of the membrane of (ii) a treated flue gas stream from which at least 50% of the CO.sub.2 content of the flue gas has been removed; (iv) removing from the permeate side of the membrane in (ii) at a pressure of 0.1 to 0.4 bar, a permeate gas enriched in CO.sub.2 and H.sub.2O; (v) cooling the permeate gas from (iv) to a temperature of 5-30° C. to condense a portion of the H.sub.2O content of the permeate gas and so lowering the H.sub.2O concentration of the permeate gas to produce a water-depleted permeate gas; (vi) separating the condensed water from the water-depleted permeate gas; and (vii) using a vacuum pump to bring the water-depleted permeate gas from (v) to atmospheric pressure or above.

46. The process of claim 45, wherein the membrane process removes 50-80% of the CO.sub.2 content of the flue gas.

47. The process of claim 45, wherein the flue gas to the membrane has a temperature between 50-80° C.

48. The process of claim 45, wherein the water vapor content of the flue gas is 70-100% of the water saturation concentration of the flue gas.

49. The process of claim 45, wherein a direct contact cooler is used to adjust the temperature and water concentration of the flue gas in step (i).

50. The process of claim 45, wherein the flue gas is generated by a coal power plant, a natural gas power plant, a natural gas boiler, a cement plant, a steel plant or an oil refinery.

51. The process of claim 45, wherein the difference in the water concentration of the flue gas in (i) and the permeate gas enriched in CO.sub.2 and H.sub.2O in (iv) differ by at least a factor of 2.

52. The process of claim 45, wherein the difference in temperature between the flue gas passing across the membrane in (ii) and the cooled permeate gas in (v) is more than 30° C.

Description

BRIEF DESCRIPTION OF THE FIGURES OF THE INVENTION

(1) FIG. 1A is generic block diagram of a prior art membrane separation process.

(2) FIG. 1B is a block diagram similar to FIG. 1A, illustrating the prior art membrane separation process used for a feed gas at 1.0 bar and containing 10% CO.sub.2 and 90% N.sub.2.

(3) FIG. 2 is a simple block diagram of the unit operations utilized in the process of the invention.

(4) FIG. 3 is an illustration of a specific example of the process of our invention.

(5) FIG. 4 is a plot showing the benefit of raising the water concentration in the membrane feed gas prior to separation.

(6) FIG. 5 is a plot showing the effect of CO.sub.2 capture rate on system performance.

DETAILED DESCRIPTION OF THE INVENTION

(7) In the text that follows the concentration of the components in the gas are molar concentrations unless otherwise stated. Also, all process pressures are in bar absolute.

(8) The separation process of this invention as applied to the separation of CO.sub.2 from flue gas is shown in a simple form in the block diagram of FIG. 2. This diagram shows the process of the invention as four steps.

The Pretreatment Step

(9) Pretreatment: The incoming gas to the process (201) is a CO.sub.2-containing flue gas containing 4-25% CO.sub.2. The gas will normally be discharged to a chimney at close to atmospheric pressure however gas blowers may be used to raise the gas pressure to 1-1.1 bar. The gas will often already contain a relatively high concentration of water and it may be possible to treat the gas as is by the process of our invention. However, it may be required to bring the gas to a controlled temperature and water vapor content by sending the gas through a direct contact water spray tower in which water at an appropriate temperature is sprayed into the gas. Such a device has the additional benefit of removing particulates and other contaminants that may be present in the flue gas. Other devices, including heat exchangers, blowers, etc., can also be used to bring the gas to the required temperature, pressure and humidity content.

(10) In this step (203), the CO.sub.2, N.sub.2-containing flue gas is brought to a composition, temperature and pressure suitable for the separation steps that follow. The flue gas mixture may already contain some water, but additional water (202) may be added to increase water vapor, CO.sub.2 and N.sub.2 composition required. The gas leaving this operation will normally be between 70-100% of its water saturation content.

(11) In an exemplary embodiment, the treated flue gas (204) meets several requirements. First, the temperature of the gas is at least 30° C. and preferably 40° C. higher than the condensation step (208) that follows to allow a useful fraction of the water vapor in the membrane permeate (207) to be removed by cooling. In principle, the condensation step (208) can be carried out at any low temperature, but to be economically viable, separation of CO.sub.2 from flue gas has to be a low-cost process. The cooling available is usually provided by an evaporative cooling plant and the cooling water produced will normally not be below 15° C. This means that if the pretreated gas to the membrane unit (204) is 30-40° C. higher than the cooled permeate gas (210), the minimum temperature of the pretreated gas (˜20° C.) is about 50° C. to 60° C.

(12) The upper temperature bound for the pretreated gas is set by the cost of providing energy to heat and humidify the gas, and the stability of the available membranes (205) at high temperatures. In general, the maximum treated flue gas temperature is 90° C. and more preferably 70-80° C.

(13) The water content of the pretreated flue gas should be high to produce high concentrations in the water vapor enriched gas delivered to the permeate condensation step. The pretreated gas (204) should contain at least 10% water vapor, more preferably at least 15% water vapor, and most preferably at least 25 wt % water. Achieving these water vapor concentrations in gas streams between 50 and 90° C. means that the pressure of the pretreated flue gas cannot be more than 2 bar, and more normally will be below 1.5 bar. Also the gas should be at at least 70% of its saturation value and preferably close to 80 or 90% of its saturation value.

The Membrane Separation Step

(14) The pretreated flue gas (204) is passed through to the membrane separation step (205) fitted with membranes permeable to water vapor and CO.sub.2, and relatively impermeable to N.sub.2, O.sub.2, and Ar. Many polymeric membranes have these characteristics, but the most suitable membranes are made of polar rubbery materials such as the family of polyamide polyether block copolymers sold under the trade name Pebax®. The Polaris™ membrane made by Membrane Technology and Research, Inc. can also be used. Most of the membranes in current use for this type of application use these types of polymers fabricated into multilayer composite membranes. By making the selective layer of the membrane very thin, on the order of 0.1-0.5 μm, it is possible to produce membranes with CO.sub.2 permeances of 1000-2000 gpu at 30° C. (1 gpu=1×10.sup.−6 cm.sup.3 (STP)/cm.sup.2.Math.sec.Math.cmHg). The permeance of these membranes will increase 2-fold or more at temperatures of 50-80° C. At temperatures of 25-30° C., good quality CO.sub.2 separating membranes when operated with flue gas will have CO.sub.2/N.sub.2 selectivity in the range of 25-50. When operated at higher temperatures, permeance will increase but CO.sub.2/N.sub.2 selectivity may fall to the 20-30 range. Because water is a small and condensable molecule, water permeances through almost all membranes are high and significantly higher than CO.sub.2. Typical water/CO.sub.2 selectivities will be in the range of 2-10 under the high-water content, high-feed gas temperatures required for the process of this invention.

(15) Obtaining a useful separation requires a high-permeance, high-selectivity membrane, but it also requires that the pressure ratio across the membrane be at least above 5 for the reasons described earlier. Because the maximum pressure is 1.5 to 2.0 bar, it follows that generating pressure difference across the membrane of our process requires a low-pressure on the permeate side of the membrane below 0.3 to 0.4 bar. The lowest practical pressures for large industrial plant are in the range of 0.1 to 0.2 bar. So the preferred operating range on the permeate side of the membrane is 0.1 to 0.4 bar

The Condensation and Separation Step

(16) The easiest way to generate low-pressure on the permeate side of our process is to use a vacuum pump. However, such pumps are expensive and consume a large amount of energy. This problem is overcome in our invention by using a cooling and condensation step (208) before the gas is sent to the vacuum pump (211). By cooling and condensing much of the water content of the gas (207), the volume of gas (210) sent to the vacuum pump is significantly reduced.

(17) When the membrane feed gas meets the composition and temperature requirements described above, the membrane system will produce a permeate gas at a temperature of 50-90° C. containing 40-70% water. Cooling this gas to about 20° C. even at sub atmospheric pressures of 0.1 to 0.4 bar causes the bulk of the water vapor content in the gas to condense and be removed as liquid water. The volume of CO.sub.2 and N.sub.2 left in residual water sent to the vacuum pump is then much smaller, and so the size of the vacuum pump (211) needed is reduced. More importantly, condensing and removing the water vapor enriches the CO.sub.2 in the permeate gas. The CO.sub.2 content of the gas is twice concentrated, once in the membrane separation step (205) and again in the water vapor condensation step (208).

The Vacuum Step

(18) The final step in the FIG. 2 process is the vacuum step (212) where the CO.sub.2, N.sub.2 and residual water vapor are compressed to atmospheric pressure or above to be discharged or sent to some other processes. If required, a final cooling step can be used after the vacuum pump to remove the remaining water.

EXAMPLES

(19) In the example calculations used to illustrate our invention that follow, we will use the permeance properties shown in Table 3. However it is not implied that these permeance and selectivity values limit the scope of the invention. All the invention requires is a CO.sub.2/N.sub.2 selectivity of at least 10, recognizing that selectivities of up to 50 or more may be possible. It also requires an H.sub.2O/CO.sub.2 selectivity of greater than about 2, but recognizing that selectivities of up to 10 or more are also possible.

(20) TABLE-US-00003 TABLE 3 Gas CO.sub.2 H.sub.2O N.sub.2 O.sub.2 Membrane 2,000 5,000 80 160 permeance (gpu) CO.sub.2/- 1 0.4 25 12.5 selectivity

Example One. The Process of FIG. 3, Varying the Feed Gas Temperature

(21) This example illustrates the benefits of our invention by operating the membrane separation system shown in FIG. 3 at different temperatures. In this example, the treated flue gas feed (301) entering the membrane unit (309) will range from 50-80° C. In all calculations reported, the gas is saturated with water vapor at the feed temperature and pressure of 1.0 bar. However, in industrial plants, the gas may not be completely saturated and relative humidities of 70-100% are possible. To simplify the calculation, the effect of temperature on the permeance and selectivity will be ignored and all calculations will use the permeance numbers shown in Table 3.

(22) Table 4 shows the results at a membrane feed gas temperature of 80° C. and Table 5 shows the results at a feed temperature of 40° C. In both examples, the temperature of the condensation step is set at 20° C. Thus, the difference in temperature between the membrane feed (301) and condensation step (304) is 60° C. in Table 4 and 20° C. in Table 5. In both of the examples, the flue gas feed gas (312) on a dry basis has a composition of 10% CO.sub.2, 90% N.sub.2 and has a flowrate of 5,100 (std) m.sup.3/h. The gas contains 1 ton/h of CO.sub.2. Before delivery to the membrane module, the gas is brought to the required temperature and is saturated with water in the pretreatment unit (313). At 80° C., the gas (301) contains 47.4% water, at 40° C. the gas (301) contains 7.4% water. In both examples, the membrane unit (309) has the membrane area required to remove 80% of the CO.sub.2 content of this gas into permeate stream (303). The characteristics of the key streams of the process are given in Tables 4 and 5.

(23) TABLE-US-00004 TABLE 4 Membrane Process Treats 80° C. Water Saturated Feed Permeate Permeate Permeate Feed Residue before cond. before vac. after pump (301) (302) (303) pump (304) cond. (306) Component Conc. (mol %) CO.sub.2 5.3 2.3 7.9 35.6 39.1 N.sub.2 47.3 88.2 11.6 52.7 58.0 H.sub.2O 47.4 9.6 80.6 11.7 2.9 Pressure (bar) 1.0 1.0 0.2 0.2 1.0 Temp (° C.) 80 78 79 20 20 *Membrane Area = 3,925 m.sup.2, Vacuum pump power (@ 80% efficiency), 84.2 kWh

(24) TABLE-US-00005 TABLE 5 Membrane Process Treats 40° C. Water Saturated Feed Permeate Permeate Permeate Feed Residue before cond. before vac. after pump (301) (302) (303) pump (304) cond. (306) Component Conc. (mol %) CO.sub.2 9.3 2.8 22.1 24.0 26.3 N.sub.2 83.4 95.6 59.3 64.3 70.8 H.sub.2O 7.4 1.8 18.6 11.7 2.9 Pressure (bar) 1.0 1.0 0.2 0.2 1.0 Temp (° C.) 40 40 40 20 20 *Membrane Area = 6,450 m.sup.2, Vacuum pump power (@ 80% efficiency), 127.8 kWh

(25) Comparing these two tables, it is clear that even though the membrane properties and pressures across the membrane are the same, operating the process with water saturated gas at 80° C. produces a much better result than operating the process at 40° C. At 80° C., the membrane area required to remove 80% of the CO.sub.2 is 40% less than at 40° C. This benefit is a result of the permeate side dilution effect of the co-permeating water. The water that permeates the membrane dilutes the CO.sub.2 in the permeate gas. This dilution increases the CO.sub.2 partial pressure driving force across the membrane, increasing the CO.sub.2 flux. As a result, the membrane area required (310) to remove 80% of the CO.sub.2 from the membrane feed gas (301) is reduced. The process also uses 34% less power, and after cooling and condensation, that removes most of the co-permeated water. This benefit is caused by the increase in the CO.sub.2 concentration and the reduction in volume in the gas leaving the permeate condenser (304) going to vacuum pump (311). Finally, the gas leaving the process (306) has a significantly higher CO.sub.2 concentration at 39% versus 26.3%.

(26) FIG. 4 shows in graphical form results of additional calculations in which these key process benefits described above are shown over the membrane feed gas temperature range 30-80° C., all at 80% CO.sub.2 capture into stream (306). At operating temperatures below about 50° C., the temperature difference between the feed (301) and the condensed gas (304) is only 30° C., so the benefit of adding water to the flue gas stream is small. This is because the water content of the gas is less than 10%, so that the concentration of water in permeate (303) is not very high and condensation to 20° C. to remove the water is less effective. At higher temperatures, such as 80° C., the difference in the temperature is 60° C., so the water content of the feed gas is significantly higher, in the range of about 20-50% water and so the permeate side dilution effect of water is much more significant. The difference in temperature should normally be at least 40° C. In this process, the feed gas to the membrane should contain at least 15% water and most preferably at least 25% water to achieve a benefit large enough to make the process worthwhile. This is the most preferred operating range.

(27) The need to have a relatively high concentration of water on the feed side of the membrane to produce a useful improvement in process performance also explains why the process is limited to low pressures on the feed side and vacuum operation on the permeate side. At a feed pressure of 1 bar and 80° C., the maximum concentration of water in the feed gas is 47% water, at 2 bar and 80° C., the water content is only 24%, and at 3 bar only about 16%.

Example 2. The Process of FIG. 3 at 80° C. with Saturated Feed Gas and Varying the CO.SUB.2 .Capture Rate of the Process

(28) The capture rate in the process illustrated in FIG. 3 can be varied by changing the membrane area. Table 6 shows this effect for a flue gas feed at 80° C. in which the CO.sub.2 capture rate is varied from 30-90%. FIG. 5 shows these same results in graphical form. These results show that both the power consumption and membrane area used per ton of CO.sub.2 decrease as the capture rate is reduced. The CO.sub.2 concentration in the final permeate (306) on a dry basis is also increased as the capture rate is reduced. This makes downstream treatment of the permeate to bring the CO.sub.2 concentration to greater than 95% easier. However, users of the technology need to reduce their CO.sub.2 impact on the environment, and so plant builders are motivated to obtain high CO.sub.2 capture rates. Also, the cost of installing the process is about the same irrespective of the capture rate, so this also favors high capture rates. It follows that a tradeoff exists between capture rate and process cost. The best range for the process is in the range of 50-80% capture. Because of the rapid increase in area and power needed, capture rates above 80%, although possible by our process, are not preferred

(29) TABLE-US-00006 TABLE 6 The Process of Figure 3 with feed gas saturated with water vapor at 80° C. The membrane area is varied to change the CO.sub.2 capture rate from 30-90% CO.sub.2. The membrane permeances are shown in Table 3. Power Consumption Conc. of separated of Vacuum Pump Area of Membrane CO.sub.2 (kW) (m.sup.2) CO.sub.2 Capture Rate (%) (306) (311) (310) 90 29.8 126.7 6,490 80 39.4 84.5 3,920 70 46.9 61.6 2,640 60 52.6 46.9 1,890 50 56.6 36.0 1,400 40 59.6 27.2 1,030 30 61.8 19.6 730

Example 3 the Process of the Invention Applied to Coal Power Plant Flue Gas

(30) The flue gas from a modern coal power plant typically contains about 12% CO.sub.2, 18% H.sub.2O and 70% N.sub.2, O.sub.2, and Ar. The gas will usually have been treated by a flue gas desulfurization unit which it leaves at a few degrees above its dewpoint, generally at 56-58° C. Table 7 shows a trial calculation using the process of FIG. 3 to treat this gas. The gas (301) is at 58° C., so the operating temperature is at the lower end of what is desirable for our process. However the feed gas if used as is, without pretreatment, still contains 18% H.sub.2O, so the permeate contains more than 40% water. When the permeate (303) is cooled to 20° C., more than 80% of the water content is condensed, significantly reducing the size of the vacuum pump needed and enriching the CO.sub.2 content of the gas after cooling (304) to 42.1% CO.sub.2. The concentration of the CO.sub.2 rises to 46.6% if a final optional (20° C.) condenser (308) is used to produce gas (306).

(31) TABLE-US-00007 TABLE 7 Feed Gas (301) at 58° C. Permeate Permeate Permeate Feed Residue before cond. before vac. after pump (301) (302) (303) pump (304) cond. (306) Component Conc. (mol %) CO.sub.2 12.0 5.3 27.0 42.1 46.5 N.sub.2 67.0 85.8 27.5 42.9 47.4 H.sub.2O 18.0 5.5 43.3 11.7 2.3 O.sub.2 3.0 3.4 2.2 3.3 3.7 Pressure (bar) 1.0 1.0 0.2 0.2 1.0 Temp (° C.) 58 57 57 20 20 *Membrane area = 2,310 M.sup.2; Vacuum pump power (@ 80% effic.) = 62.8 kWh; 70% CO.sub.2 capture

(32) Depending on the power plant, low-grade heat in the range of 70-100° C. may be available. If this heat is available, it can be used to increase the temperature and water saturation concentration of the gas being treated by the membrane. An increase of even a few degrees from 58 to 65° C., for example by raising the water concentration in the feed gas to 25% H.sub.2O, produces useful improvements in the process as the data in Table 8 show.

(33) TABLE-US-00008 TABLE 8 Feed gas (301) at 65° C. Permeate Permeate Permeate Feed Residue before cond. before vac. after pump (301) (302) (303) pump (304) cond. (306) Component Conc. (mol %) CO.sub.2 11.0 5.1 21.9 42.3 50.0 N.sub.2 61.2 83.9 19.2 39.9 44.1 H.sub.2O 24.0 7.6 57.3 11.7 2.3 O.sub.2 27 3.4 1.5 3.2 3.5 Pressure (bar) 1.0 1.0 0.2 0.2 1.0 Temp (° C.) 65 64 65 20 20 *Membrane area = 2.050 m2; Vacuum pump (@ 80% effc.) = 58.3 kWh, 70% CO.sub.2 capture.

(34) The residue stream from the process (302) can be discharged to a chimney. The final permeate (306) contains 50% CO.sub.2 and may find use as in various algae or CO.sub.2 processing applications, in cement plants for example. More commonly, the gas can be sent for further concentration by absorption, membrane, or cryogenic processes to produce more than 98% CO.sub.2 for sequestration or use in enhanced oil recovery processes.

Example 4. The Process of the Invention Applied to a Natural Gas Boiler Exhaust or a Natural Gas Power Plant Fitted With Partial Exhaust Gas Recycle

(35) The exhaust from a natural gas plant boiler used to produce high temperature steam will typically be very hot, often about 150° C., and will have a typical composition of about 8% CO.sub.2, 16% H.sub.2O, 4% O.sub.2 and 72% N.sub.2. The dew point of the gas is about 56° C., but if the gas is cooled by contacting with H.sub.2O in a direct contact cooler, the saturation point of the gas when cooled from 150° C. is 62° C. At this temperature, the gas contains 20.9% H.sub.2O. The performance of our membrane process using the FIG. 3 design is given in Table 9.

(36) TABLE-US-00009 TABLE 9 Feed gas (301) at 62° C. Permeate Permeate Permeate Feed Residue before cond. before vac. after pump (301) (302) (303) pump (304) cond. (306) Component Conc. (mol %) CO.sub.2 7.5 3.3 16.9 31.2 34.5 N.sub.2 67.8 85.8 28.1 52.0 57.6 H.sub.2O 20.9 6.6 52.3 11.7 2.3 O.sub.2 3.8 4.2 2.7 5.0 5.6 Pressure (bar) 1.0 1.0 0.2 0.2 1.0 Temp (° C.) 62 61 61 20 20 *Membrane area = 3,770 m2; Vacuum pump (@ 80% effc.) = 85.0 kWh, 70% CO.sub.2 capture

Example 5. The Process of FIG. 3 Using Sequential Condensation Systems

(37) In the example calculations reported in Tables 7 to 9 to illustrate the process of the invention, a one-stage condenser represented as unit (314) is shown in FIG. 3. Such a simple system could be used, but in large systems, multiple condensers in series might be used to reduce the amount of refrigerated cooling water needed. For example, a first condenser using 25° C. cooling water could be used to bring the permeate gas to 30° C., a second condenser using 20° C. cooling water could be used to bring the gas to 25° C. while a final condenser using chilled water at 15° C. could then be used to bring the gas to 20° C., while refrigerated water could be used to bring the gas to even lower temperatures of 5-10° C. if desired. Using these systems reduces the volume of gas going through the vacuum pump, and so reduces the power consumption of unit (311). However, this benefit has to be offset against the cost of providing chilled water. By using sequential cooling steps, the cost of providing the chilled cooling water required is reduced.

(38) When introducing elements of the present disclosure or the preferred embodiment(s) thereof, the articles “a”, “an”, “the” and “said” are intended to mean that there are one or more of the elements. The terms “comprising”, “including” and “having” are intended to be inclusive and mean that there may be additional elements other than the listed elements.

(39) In view of the above, it will be seen that the several objects of the disclosure are achieved and other advantageous results attained.

(40) As various changes could be made in the above products and methods without departing from the scope of the disclosure, it is intended that all matter contained in the above description shall be interpreted as illustrative and not in a limiting sense.