SYNGAS MIXTURES

INCORPORATION BY REFERENCE

[0001] All publications, patents, and patent applications mentioned in this specification are herein incorporated by reference to the same extent as if each individual publication, patent, or patent application was specifically and individually indicated to be incorporated by reference.

BACKGROUND

[0002] CO is used as a feed stock for the production of a number of important chemicals such as methanol, acetic acid, phosgene and ethylene glycol. In the production of these chemicals, it is desirable that the CO feed contain at least 90% CO and preferably greater than 95% or 98% CO.

[0003] The lowest cost method of making CO is the reaction of natural gas, coal, coke or other carbonaceous fuels with steam and oxygen to produce a mixture of CO, H ₂ , C0 ₂ and small amounts of N ₂ , CH ₄ , and Ar. Depending on the initial fuel, the CO content of the gas is typically in the 25% to 50% range. The first step in producing CO from this mixture usually involves treating the mixture with an acid gas removal process. Absorption processes are most commonly used. These processes involved contacting the gas with a liquid solvent which selectively removes C0 ₂ , H ₂ S and COS. The absorption characteristics of the solvent used determine the efficiency of the process. Absorption can be simple physical sorption or involve a reversible chemical reaction between the acid gases and components of the solution.

[0004] Absorption processes generally reduce the C0 ₂ concentration in the syngas to less than 1% C0 ₂ and achieve almost complete removal of H ₂ S and COS. In some cases, this low level of C0 ₂ may be sufficient but if even lower levels are needed, then a final polishing step with a molecular sieve absorption unit can be used.

[0005] Syngas, after the absorption step, if made from natural gas, will generally contain 25% to 30% CO and 70% to 75% H ₂ . If made from coal, the gas will generally contain 30% to 50% CO and 50% to 70% H ₂ . These types of gas mixtures are used as the feed to the CO separation processes described herein.

[0006] In the past, absorption processes have been used to separate CO/H ₂ mixtures. Such processes are described in U.S Patent 2,519,284 by Ray and Johnson, and in U.S. Patent 4,950,462 by Nakao, et al. The processes described use cuprous salt solutions to selectively absorb CO. The CO reacts with the cuprous salt and is removed. This technology has been commercialized under the trade names COSORB and Copure®. However, these processes have not been widely used because of a number of reasons, for example, the poor stability of the cuprous salt absorption solutions.

[0007] Membrane separation has also been used to treat CO/H ₂ gas mixtures. Membranes are known that selectively permeate H ₂ and retain CO. Successful applications of this technology generally involve removing a portion of the H ₂ from a higher pressure H^CO mixture, to reduce the H ₂ content of the gas and so change the ratio of H ₂ to CO in the gas from 3/1 to 2/1 for example. This process is useful in producing a gas mixture having the ratio required for later chemical synthesis reactions. The process is called ratio adjustment. Changing the H ₂ /CO ratio from 75% H ₂ /25% CO (3/1) to 66% H ₂ /33% CO (2/1) means the feed gas entering the membrane unit contains 75% H ₂ and leaves containing 66% H ₂ . The average H ₂ concentration on the feed side of the membrane is therefore high and the CO concentration in the permeate gas is low, so relatively little CO is lost in the permeate gas.

[0008] However, when these same membranes are used to produce CO as the product gas containing greater than 90% CO, the concentration of CO in the feed is high and a significant amount of CO permeates with the separated H ₂ . At least for this reason, this membrane process has not been used for CO production from syngas.

[0009] Currently a widely used process for CO separation from CO/H ₂ mixtures is cryogenic distillation. This process involves pressurizing the feed gas to 50-60 bar and cooling to a very low temperature to liquefy the CO. The cost of providing the low temperature and high pressure needed make this process costly, and so better, lower cost processes are still sought. [0010] The membrane process described herein uses at least two membrane units to produce higher purified CO and to permeate streams of different compositions. These permeate gas mixtures are then treated separately to produce H ₂ and/or for H^CO mixtures of a desired composition. The process produces as its main product high-pressure CO. Depending on the process requirements, a H ₂ stream is made and sometimes a CO/H ₂ stream. The composition of the CO/H ₂ stream can be adjusted to make it easily usable in downstream chemical production processes.

SUMMARY

[0011] In one aspect, the disclosure relates to a process for separating a gas mixture to produce a high concentration CO stream, the process comprising the following steps: (a) passing a high- pressure feed gas mixture comprising of CO and H ₂ to a first membrane separation unit fitted with a membrane that selectively permeates H ₂ and retains CO; (b) producing a H ₂ -emiched gas as the first membrane unit permeate gas containing greater than 70% H ₂ , and producing a H ₂ - depleted residue gas from the first membrane unit; (c) passing a portion of the H ₂ -enriched permeate gas from (b) to a pressure swing absorption system to produce a twice enriched H ₂ product gas containing greater than 90%, preferably greater than 95%, H ₂ and low-pressure tail gas enriched in CO; (d) passing the first membrane unit H ₂ -depleted residue gas from (b) to a second membrane unit containing a membrane that selectively permeates H ₂ and retains CO; (e) removing a second permeate gas from the second membrane unit and producing a CO-enriched residue gas from the second membrane unit, wherein the CO-enriched product gas is the enriched residue gas containing greater than 90% CO; and/or (f) creating a gas mixture comprising H ₂ and CO with a preset H ₂ /CO ratio between 3.5 to 1.0 by mixing the second membrane unit permeate gas from (e) with a portion of the first membrane unit permeate from

(b) and the pressure swing adsorption tail gas from (c). In some embodiment, at least a portion of the low-pressure tail gas enriched in CO from (c) is recycled to the feed gas in (a). In some embodiments, at least a portion of the second permeate gas from (e) is recycled to the feed gas in (a). In another embodiment, at least a portion of the low-pressure tail gas enriched in CO from

(c) and at least a portion of the second permeate gas from (e) is recycled to the feed gas in (a). In some cases, at least one of the membrane units are fitted with a residue sweep gas. The first membrane unit can operate at a permeate pressure ratio of 5 to 10 and the second membrane unit operates at a pressure ratio 10 to 20. The second membrane unit can operate at a permeate pressure at least 50% lower than the first membrane unit operating pressure. The second membrane unit can operate at a permeate pressure no more than 50% lower than the first membrane unit operating pressure. In some cases, the first membrane separation unit is fitted with more than one membrane. The twice enriched H ₂ product gas in (c) may contain greater than 95% H ₂ . The low-pressure tail gas enriched in CO in (c) may contain C0 _2, H ₂ , N ₂ , CH ₄ , and Ar. In some embodiments, the feed gas mixture comprises primarily of CO and H _2.

[0012] In one embodiment, a portion of the second membrane unit permeate from (e) and a portion of the PSA tail gas from (c) are mixed with a portion of the first unit permeate gas from (b) to produce a mixed H ₂ /CO-containing syngas of an appropriate ratio determined by the amount of the first unit permeate added to syngas mixture. In a preferred embodiments, all of the second membrane unit permeate from (e) and all of the PSA tail gas from (c) are mixed with a portion of the first unit permeate gas from (b) to produce a mixed H^CO-containing syngas of an appropriate ratio determined by the amount of the first unit permeate added to syngas mixture.

[0013] In another embodiment, portions of the second membrane unit permeate from (e) and/or the PSA tail gas from (c) maybe be compressed and recycled to the incoming H ₂ /CO feed gas in (a).

BRIEF DESCRIPTION OF THE DRAWINGS

[0014] Figure la depicts two connected membrane units illustrating the employment of a small residue sweep gas recycle process on the permeate side of the membrane.

[0015] Figure lb depicts two connected membrane units illustrating a sweep recycle unit that recycles 5% of the residue gas leaving the unit as a low-pressure counter-flow sweep.

[0016] Figure 2 is an illustration of a two-step membrane separation process using adsorption built according to the invention. [0017] Figure 3 is a schematic illustrate a basic process in which the PSA tail gas and second membrane permeate gas are recycled to the incoming CO/H ₂ feed.

[0018] Figure 4 is a schematic drawing using a preferred process design in which the CO-rich tail gas from the PSA unit and the H ₂ -rich permeate from the second stage membrane unit are mixed with sufficient of the H ₂ -rich gas from the first stage membrane unit to produce a useful synthesis gas product.

[0019] Figure 5 is a schematic showing a preferred process design in which the permeate gas from the second membrane unit is mixed with sufficient of the H ₂ -rich first membrane unit permeate gas to produce a useful synthesis gas mixture and the tail gas from the PSA unit is recycled to the feed.

[0020] Figure 6 is a schematic of a preferred process design in which the CO-rich PSA tail gas is mixed with a portion of the H ₂ -rich first membrane unit permeate to produce a useful synthesis gas mixture while the permeate gas from the second membrane unit is recycled to the feed.

DETAILED DESCRIPTION

[0021] Gasification of coal to produce syngas for coal-to-chemicals or hydrogen for IGCC power production is an area of growing interest. These applications require large gas separation plants as part of the flow scheme. Despite advances in conventional separation technologies over the past several decades, costs of gas separation systems remain high. Membrane gas separation is an emerging new process that has a number of useful attributes, including but not limited to, low cost, straight forward flow sheet, small foot print, and ease of operation and no hazardous chemicals to worry about.

[0022] In general, membranes are best used to perform a low cost bulk separation. Sometimes this type of separation is all that is needed. In other cases, a second separation technology will be used in a subsequent polishing step. In some cases, membrane separation technology is combined with an additional process to form combination processes that are more efficient and cost effective. Such combination processes involves at least two processes, including but are not limited to, membrane separation plus cryogenic fractionation, membrane separation plus pressure swing adsorption, and membrane separation plus absorption processes.

[0023] In some cases, the combination process can be used to perform the separation required for gasification streams from a variety of coals and other feedstocks, such as biomass and municipal or industrial wastes.

[0024] The present disclosure relates to an improved process to produce CO from a H ₂ /CO gas mixture while simultaneously producing H ₂ and optionally a H ₂ /CO syngas mixture as coproducts. The process is a combination of membrane separation and pressure swing adsorption (“PSA”).

[0025] It will be apparent that the figures shown of the process schematic herein are very simple block diagrams intended to make clear the key operations of the embodiment process of the invention, and that actual process designs may include additional steps of standard types such as heating, chilling, compressing, condensing, pumping, various types of separation as well as monitoring the process, temperature, flows and the like. It will also be apparent to those of skill in the art that the details of the unit operations used may differ from process to process according to the needs of particular plant operations.

The Membranes Used

[0026] The membranes used in this process are anisotropic structures consisting of a thin dense polymer layer formed on a microporous support layer. The material used for the two layers may be the same or different. The key requirement is that the membranes are able to support a high- pressure difference across the membrane and that the membrane is permeable to H ₂ but relatively impermeable to CO, CH ₄ , N ₂ and Ar. Membranes of this type are available from a number of companies including Air Products (Permea), Air Liquide (Medal), UBE industries, and Membrane Technology and Research, Inc. [0027] Ceramic, palladium metal or microporous carbon membranes also have the permeability and selectivity requirements to perform this separation. However these membranes are much more expensive than simple anisotropic dense polymer membranes, and although occasions may occur where these membranes can be used, they are generally not preferred.

[0028] The membranes used may be manufactured as flat sheets or as hollow fibers and can be packaged in any convenient form including spiral-wound modules, plate and frame modules and potted hollow fiber modules. The making of all these types of membranes and modules can be any method used in the field. Hollow fiber modules are a preferred design for the counterflow with sweep operations used for the most part in this process.

[0029] Gas permeation in dense polymer membrane films can be rationalized using the basic solution diffusion equation:

where j _A is the molar flux (cm ³ (STP)/cm ² -s) of component A , £ is the film thickness, p ^eed and p ^Permeate _are the partial vapor pressures of component A on the feed side and permeate side of the membrane, and P _A is the permeability to component A of the membrane material, usually expressed in Barrer (where 1 Barrer = 1 x 10 ¹⁰ cm ³ (STP)-cm/cm ² -s-cmHg).

Rearranging equation 1 :

[0030] The expression on the left of equation 2 is the pressure-normalized flux, and is numerically equal to the thickness -normalized permeability, usually referred to as permeance, on the right. Pressure-normalized flux or permeance is usually expressed in gas permeation units or gpu (where 1 gpu = 1 x 10 ⁶ cm ³ (STP)/cm ² -s-cmHg).

[0031] The membrane selectivity, a, for one component over another is expressed as the ratio of the permeabilities (or permeances) of the components. Thus, for two components A and B:

[0032] Permeance and selectivity are properties that characterize a membrane, and the search for membrane materials with improved properties continues. In general, there is an inevitable tradeoff between permeability (or permeance) and selectivity. Materials that exhibit high permeability tend to exhibit low selectivity and vice versa.

[0033] The permeance values for the H ₂ selective membranes used in the examples of this invention are given in Table 1 below. These numbers are representative of what can be obtained with commercial H ₂ -selective membranes currently on the market.

Table 1. Membrane Permeances

• lgpu = lxlO^ cm ³ (STP)/cm ² sec-cmHg

[0034] Another factor that affects membrane performance is the pressure ratio which is the ratio of the total feed pressure divided by the total permeate pressure. A high pressure ratio can increase the overall separation performance and so a low permeate pressure can be desirable. However, a low permeate pressure means more compression will be required to bring the permeate gas to a suitable pressure for recycle or other use. For example, a pressure ratio of 50 would mean that the permeate stream would have to be recompressed 50-fold to be recycled within the process. Furthermore, to achieve a high pressure ratio will demand larger, more powerful pumps and compressors, and thus pressure ratio tends to be limited by cost considerations. For these reasons, in industrial gas separation processes the pressure ratio across the membrane is typically in the range of 5 to 20.

[0035] Pressure ratio is defined as qi , thus

where p[ ^esd is the total pressure on the feed side of the membrane and _lS the total pressure on the permeate side of the membrane. As mentioned above, a high pressure ratio may improve the separation performance of the process, but at the expense of greater energy to produce the high ratio.

[0036] The enrichment, E, of a component provided by a membrane separation operation is expressed as the ratio of the concentration C of that component on the permeate and feed sides. Thus, for component A , the enrichment E _g in the first membrane separation step is given by where C _A ^{p rmeate} is the concentration of component A on the permeate side and C _A ^ed is the concentration of component A on the feed side.

[0037] For component A to flow from the feed to permeate side, the partial pressure of A in the permeate must remain lower than the partial pressure of A in the feed, thus: permeate permeate ✓ feed f feed

Pi ^x - Pi ^x (6)

Rearranging

Or

[0038] Thus, the enrichment is always numerically less than the pressure ratio, and

[0039] In principle, with an infinitely selective membrane, the permeate concentration of component A could reach 100 %, but the permeate concentration in practice is limited by expression (9). With a pressure ratio of 5 and a feed concentration of 10 vol%, for example, the maximum permeate concentration, irrespective of membrane selectivity, is 50 vol%. Likewise, with a pressure ratio of 30 and a feed concentration of 3 vol%, the permeate concentration can not exceed 90 vol%.

[0040] In the process described herein, the feed gas is typically at a pressure of 40-50 bar and the permeate gas is at a pressure ratio 4 -5 bar, a pressure ratio of 10. The key separation is H ₂ and CO for which most membranes of the type represented in Table 1 have a selectivity of 40. It follows that, so long as the H ₂ concentration in the feed exceeds 10%, the permeate H ₂ concentration is predominantly determined by the membranes selectivity. The pressure ratio across the membrane has an effect but it is not limiting. However, at feed H ₂ concentrations of less than 10%, particularly at feed H ₂ concentrations of 1% to 3%, the permeate H ₂ concentration is increasingly limited by the low pressure ratio rather than the membrane selectivity.

[0041] One method of mitigating the impact of pressure ratio on membrane performance is to expand a gas leaving the membrane unit to the permeate side pressure and use it as a counterflow permeate side sweep. The process is illustrated in Figure 1. The separation depends on how much sweep gas is used; typically the optimum is about 5% as shown in the figure. The result is significant. The partial pressure of H ₂ on the permeate side of the module is reduced and as a consequence less membrane area is needed to perform the separation. Mixing separated feed gas with the permeate gas improves the separation. The cause of this paradoxical result, illustrated in Figure 1, has been discussed by Cussler and coworkers in J of Memb Sci., v72, p22l (1992) K.L. Wang, S.H. McCrey, S.H. Newbold and E.L. Cussler.

[0042] Figure 1 shows a two-step membrane process fitted with membranes having the permeation properties listed in Table 1. The feed gas consists of 1000 std m ³ per hour of 50% H ₂ , 50% CO gas at 40 bar. The permeate pressure in all cases is at 4 bar, so the pressure ratio is 10.

[0043] In the example shown in Figure la, both membrane units consist of simple counterflow membrane modules, 120 m ² of membrane is needed by the first membrane unit to bring the gas to 10% H _2, 90% CO and another 154 m ² of membrane is needed to bring this gas to 1% H ₂ , 99% CO. The first membrane permeate contains 91% H ₂ , the second membrane unit permeate contains 40% H ₂ and 60% CO.

[0044] In the example shown in Figure lb both of the counterflow membrane modules are fitted with a sweep recycle unit that recycles 5% of the residue gas leaving the unit as a low-pressure counterflow sweep. Using the sweep gas has a small effect on the separation achieved by the membrane; the permeate gas concentrations are almost unchanged, but the membrane area required to do the separation is significantly affected. The area needed by the first membrane unit drops by 16 m ² (13.3%), while the area needed by the second membrane unit drops by 42 m ² (27.3%). The effect is bigger for the second unit because this unit is in the pressure ratio limited region. The Pressure Swing Adsorption Unit

[0045] Pressure swing adsorption is a process often used to separate H ₂ gas mixtures. The process uses an adsorbent that selectively removes components such as CO, C0 ₂ , CH ₄ , N ₂ and H ₂ S but does not adsorb H ₂ . The feed gas consisting primarily of H ₂ is passed through the adsorbent bed at a pressure of 20-30 bar. H ₂ passes freely through the bed but the other impurities are absorbed and removed. Different adsorbent materials are used depending on the impurities in the gas. Oftentimes more than one adsorbent is used in the bed. For example, a layer of activated carbon at the feed end of the bed is often used to remove H ₂ 0 and C0 ₂ followed by a layer of zeolite to remove CO and CH ₄ . The amount of the different adsorbents can be set for a particular feed so both beds become saturated at the same time. In a well- designed PSA system, a H ₂ mixed gas containing 70% H ₂ and impurities such as CO, CH ₄ , C0 ₂ , H ₂ 0 and N ₂ can produce a treated H ₂ product stream containing greater than 99.9% pure H ₂ .

[0046] When the adsorbent beds become saturated with adsorbed impurities, some breakthrough of impurities into the H ₂ gas occurs. The bed is then taken off-line and a countercurrent flow of H ₂ at low pressure is used to strip the adsorbed impurities from the adsorbent bed. The low pressure gas removed from the bed is called tail gas and will contain the adsorbed impurities and a portion of the H ₂ in the feed gas. The amount of H ₂ lost with the tail gas is a limiting factor in the PSA process.

[0047] If the H ₂ concentration in the feed gas is high, greater than 70%, preferably greater than 80% H ₂ and even more preferably greater than 90% H ₂ for example, a large amount of H ₂ can be processed before the bed needs to be regenerated. H ₂ in the tail gas is also low, with a high H ₂ feed gas concentration, and may then only be 5% to 10% of the feed H ₂ content. At low H ₂ contents in the feed tail gas, losses increase and larger adsorbent beds are required to treat the same amount of feed gas. PSA processes are therefore generally not economical if the feed H ₂ concentrate is less than 70% H ₂ , and a PSA feed gas of greater than 85% H ₂ in the feed is very much preferred. [0048] A great deal of work has been done to minimize tail gas losses and to reduce the size of the adsorbent beds needed. At a minimum, a PSA plant will consist of two adsorbent beds; one bed is used to treat the gas, while a second bed is being regenerated. However, most industrial plants will consist of 4-8 interconnected beds. A complicated series of bed pressurization and depressurization cycles are used to minimize tail gas loss. In the figures and examples used to illustrate this invention, the PSA plant is shown as a simple box. Engineers will understand that the actual plant is a good deal more complicated.

[0049] A final important variable determining PSA system performance is the pressure ratio between the pressure of the H ₂ containing feed gas, typically at 20-30 bar, and the low pressure tail gas, typically 2-5 bar. In some cases, the pressure ratio between the pressure of the H ₂ containing feed gas is greater than 20, 21, 22, 23, 24, 25, 26, 27, 28, 29, or 30 bar, and the low pressure tail gas is greater than 2, 3, 4, or 5 bar. Thus, pressure ratio is typically between 4-10. Feed pressures above 30 bar are possible but the high pressures mean the adsorbent pressure vessels become increasingly costly. Lower tail gas pressures are also possible and reduce tail gas H ₂ loss but require larger and more expensive compressors to handle the low pressure gas.

[0050] Figure 2 illustrates a basic embodiment of the process. Feed gas (200) consisting of CO, H ₂ and small amounts of C0 ₂ , CH ₄ , N ₂ and other contaminants is delivered to the first membrane unit (201). The feed gas will normally be provided by contacting the carbonaceous fuel with H ₂ 0 and 0 ₂ to provide a raw syngas of C0 ₂ , H ₂ , H ₂ S, CO, CH ₄ , H ₂ 0 and smaller amounts of N _2, Ar and COS. Water and acid gas components will usually be removed by an Amine or physical absorption process followed by a dehydration step. It is this treated gas that is the feed (200). If the gas is made from a natural gas feed, it will typically contain about 30% to 35% CO and 60% to 65% H ₂ . If the gas is made by gasification of coal it will contain more CO, typically in the range 50% to 60% and 35% to 45% H ₂ . Both of these gas streams have too little H ₂ for treatment by PSA alone. The PSA beds would be too large and H ₂ loss in the tail gas too much for a useful process.

[0051] The objective of the first membrane unit (201) is to remove the bulk of the H ₂ in the feed gas and so produce a permeate stream (213) that can now be economically treated by the PSA system (204). With typical available membranes having the selectivities shown in Table 1, it is possible to reduce the H ₂ concentration in stream (206) to between 10% and 15%, or at least 10%, or at least 11%, or at least 12%, or at least 13%, or at least 14%, or at least 15% H ₂ while still producing a PSA feed gas containing at least 70% H ₂ , or at least 75%, or at least 80%, or preferably greater than 85% H ₂ . The H ₂ concentration in the PSA feed gas (213) is affected by the pressure ratio between stream (200) and (213). Increasing the pressure ratio by reducing the pressure (213) reduces the membrane area needed and increases the H ₂ concentration in stream (213) but at the expense of a larger PSA compressor (203). The optimal trade-off between these factors will change depending on the cost of the equipment involved and the selectivity of the membrane. However, as a rule of thumb, the pressure ratio will normally be between 5 to 10. If feed gas (200) has a high H ₂ concentration, for example feed gas from a natural gas reformer, a pressure ratio of 5 may be optimal. If the feed gas has a lower H ₂ concentration, for example gas made from a coal gasification plant, a pressure ratio of 10 may be required. The residue gas leaving the first stage membrane unit (214) still has a H ₂ concentration of between 10 % to 20%, or at least 10%, or at least 12%, or at least 14%, or at least 16%, or at least 18%, or or at least 20%. This H ₂ must be removed to bring the CO concentrate stream to the target concentration of greater than 90%, or greater than 91%, or greater than 92%, or greater than 93%, or greater than 94%, or greater than 95%, of CO. Preferably, the H ₂ is removed to bring the CO concentrate stream to the target concentration of greater than 95% CO.

[0052] The second stage membrane unit (202) treats the H ₂ -depleted gas (214) and removes the bulk of the remaining H ₂ to produce the CO product gas (209) containing greater than 90% CO and preferably greater than 95% CO, containing between 1% to 3%, or at least 1%, or at least 2%, or at least 3% H ₂ , and even more preferably greater than 98% CO. The remaining components will be most of the N ₂ , CH ₄ and Ar that enter the process with the feed gas (200).

[0053] The volume of the second membrane unit permeate gas (211) is generally between 10% to 25%, or at least 10%, or at least 12%, or at least 14%, or at least 16%, or at least 18%, or at least 20%, or at least 22%, or at least 24%, of the volume of the first membrane unit permeate gas (213), but this gas has a 5 to lO-fold higher CO concentration and so is still a major contributor to the CO recovery of the process. For successful operation of the process, it is important to minimize CO loss in this gas. CO loss can be reduced by increasing the pressure ratio across the membrane and so, while the first membrane unit (201) will normally operate at a pressure ratio of 5-10, the second membrane unit may operate at a pressure ratio of 10-20.

[0054] Even when the second membrane operating conditions are optimized, this permeate gas (211) has too much CO to be ignored. In the processes described here, over these values can be recovered in two ways. One way is to send the gas to another operation in the chemical plant where a high CO concentration syngas can be beneficial if used. If no such process is available, then the gas can be recycled to the incoming feed, where the CO content is then recovered in the CO product and the H ₂ content is recovered in the PSA H ₂ product gas.

[0055] The permeate from the first membrane unit (213) will typically be produced at a pressure of 4-8 bar and so will normally be at too low a pressure to be sent directly to the PSA unit. A permeate compressor (203) is then used to bring the pressure of the PSA feed gas (215) to a convenient value, typically about 20-30 bar. The PSA unit produces a product gas of a essentially pure H ₂ (216) and a low-pressure PSA tail gas (212) containing all of the CO that permeated the first membrane unit. This tail gas (212) and the second membrane permeate (211) are both at approximately the same pressure, 2-5 bar, and both contain a relatively high concentration of CO. Recycling all or a portion of these two gas streams to the incoming feed is very worthwhile.

[0056] The following examples are intended to be illustrative of the invention but are not intended to limit the scope or underlying principles in any way.

[0057] As described herein, concentrations in this patent are molar concentrations unless otherwise stated.

EXAMPLES

[0058] Example 1. CO recovery from a H ₂ -rich syngas produced by reforming natural gas. [0059] The process used to treat this gas has the form of the basic embodiment shown in Figure 2. The initial feed gas contains 63.5% H ₂ , 30% CO, 5% C0 ₂ and small amounts of N ₂ and CH ₄ . An acid gas removal unit (not shown) is used to remove 95% of the C0 ₂ , producing feed gas (200). The first membrane unit (201) fitted with 8000 m ² of membrane, reduces the H ₂ concentration in (214) to 14.3%, producing a permeate (213) containing 92.7% H ₂ . The permeate gas is at a pressure of 7 bar so the membrane pressure ratio is about 6. After compression to 20 bar by compressor (203), this gas is a perfect feed for the PSA unit (204). The PSA unit produces pure H ₂ (216) and a small tail gas (212).

[0060] The residue gas from the first membrane unit (214) is then treated with a second membrane unit (202) to produce a CO product gas containing 93.7 % CO, 0.95% H ₂ and 5.3% of the N ₂ , CH ₄ , and C0 ₂ contaminants that enter the process with the feed gas (200). The permeate from the second membrane unit contains approximately 50% CO and represents half of the CO loss of the process.

[0061] Overall the process as described above achieves a recovery of 70.5% of the CO in the original feed as 94% CO product (209) and 86.4% recovery of the H ₂ as essentially pure H ₂ in stream (216). The remaining CO and H ₂ are contained in syngas streams (211) and (212), which can be used elsewhere in the chemical complex or recycled to the incoming feed gas.

[0062] The compositions, feed flows and pressures of the key streams in the process are shown in Table 2 below.

Table 2

Membrane area (201) 11,000 m ² , Membrane area (202) 800 m ² ,

Compressor (203) 2.4MWe theoretical

[0063] Example 2. CO recovery from a H ₂ -enriched gas with syngas recycle.

[0064] This example uses the same feed gas and membrane permeances of Example 1 with the additional step of recycle of the syngas stream to the feed to maximize CO and H ₂ production and recovery. The process flow diagram is shown in Figure 3.

[0065] As in Example 1, two membrane units are used to produce a H ₂ product stream (316), a CO product stream (309) and two syngas streams (311) and (312). In this example however, the syngas stream is mixed and recirculated back to the initial membrane feed (300) with compressor (318). By recirculating these two streams, the recovery of H ₂ and CO is increased from 70 to 100% at the expense of a new 3.6 MWe compressor and an increase of 28% in membrane area. Table 3

Compressor (303) 3.7 MWe theoretical, Compressor (318) 3.6 MWe theoretical

[0066] Example 3. CO recovery from a high CO syngas feed produced from a coal gasification plant.

[0067] Figure 4 shows a block flow diagram of the process of this invention when used to produce CO from a syngas containing a relatively high content of CO. The process is also designed to adjust the composition of the syngas produced by the process to produce a single syngas product stream in which the molar ratio of H ₂ to CO is controlled at 2/1, making the gas suitable for chemical production such as methanol. [0068] The key stream compositions, flow rates and processes of the Figure 4 process design are listed in Table 4. The membrane has the permeation properties used in the earlier examples. The initial feed (400) is made from coal, and so after treatment to remove C0 ₂ , H ₂ S, and H ₂ 0, contains 44% H ₂ , 55% CO and about 1% of minor contaminants. The first membrane unit reduces the H ₂ concentration in stream (414) to about 12%. Because the average H ₂ concentration in the gas passing through the membrane unit is lower than in the previous examples, more CO permeates and the concentration of H ₂ in the permeate gas (413) is approximately 83%. However, this is still a good feed for the following PSA step.

Table 4.

[0069] The second membrane unit removes the remaining H ₂ from stream (414) and produces a CO product (409) containing approximately 1% H ₂ and 97.4% CO. To minimize CO loss in the permeate (406) from the second membrane unit, the pressure is set at 4.0 bar, a pressure ratio of 10. However, the permeate still contains 56% CO. If mixed with the tail gas from the PSA unit

(415), this gas could be recycled to the incoming feed gas as in Example 2. Fortunately, coal-to- chemicals plants are very large operations and often produce a range of products, some of which require pure CO and pure H _2, but also other products which use a H ₂ /CO mixed gas if it has the appropriate ratio of H ₂ to CO, commonly about 2: 1.

[0070] Thus the process shown in Figure 4 produces three product streams; A high concentration CO product (409). A pure H ₂ stream (404) from the PSA (416) and a H ₂ /CO syngas mixture (411) obtained by mixing the PSA tail gas (412) and the second membrane unit permeate (406) with a portion of the H ₂ -rich first membrane unit permeate to bring the final syngas product to the appropriate H ₂ /CO ratio. [0071] About 73% of the CO in the feed gas (400) is recovered as concentrated CO in stream (409) and about 32% of the H ₂ is recovered as pure H ₂ from the PSA in stream (416) while the remaining H ₂ and CO are produced as the 2/1 syngas mixture (411). [0072] Example 4. CO recovery with recycle of the PSA tail gas.

[0073] Figure 5 shows a block diagram of the process of this invention used with a feed gas containing about 30% CO. In this example, the membrane permeate streams are manipulated to produce a pure H ₂ stream and a syngas H^CO mixture having the molar ratio 2/1, so it can be easily used for methanol synthesis. The key components of this Figure 5 process are listed in Table 5. After an initial treatment of the gas to remove H ₂ S, C0 ₂ and H ₂ 0 to very low levels, the gas consists of 31.7% CO, 67.9% H ₂ , 0.3% N ₂ , 0.1% CH ₄ . The membranes used have the properties used in the earlier examples. Table 5

[0074] The process thus creates three product streams; a CO product (509) containing 97% CO, a H ₂ product (516) containing about 100% H ₂ , and a small syngas product gas (511) with a H ₂ /CO molar ratio of 2/1. About 83% of the feed H ₂ in the feed gas (500) is contained in the H ₂ product gas (516) and 83% of the CO in the feed gas is in the CO product gas (509). [0075] The first membrane unit (501) reduces the feed gas H ₂ concentration from 66.4 % H ₂ to 20% H ₂ producing a permeate (513) containing about 94% H ₂ . The second membrane unit (502) removes almost all of the remaining H ₂ , producing a concentrate (509) consisting of 97.2% CO, 0.94% H ₂ and about 1.9% of the minor impurities CH ₄ , C0 ₂ , Ar, and N ₂ . Argon and N ₂ are relatively inert and will not normally be a problem with most uses of the CO product. If removal of the C0 ₂ , H ₂ , and CH ₄ is needed, a small catalytic oxidation step followed by a molecular sieve adsorber to remove the C0 ₂ and H ₂ 0 is all that is required.

[0076] The permeate from the second membrane unit (512) contains about 40% CO, and 59.1% H ₂ . This gas could be compressed and recycled to the feed but by mixing with about 5% of the first membrane permeate (513), a syngas (511) with the fh/CO ratio of 2/1 can be produced. This gas is then easily used elsewhere in the plant.

[0077] The remaining 95% of the first membrane stage permeate (505) contains 94% H ₂ and so is an excellent feed for the PSA unit (504), which then produces a clean H ₂ product gas (516) and a PSA tail gas (517) which, after compression, can be recycled to the incoming feed gas (500).

[0078] Example 5. CO recovery with recycle of the second membrane unit permeate gas.

[0079] Figure 6 shows a block diagram of the invention used with a feed gas containing a high concentration of CO for example, a CO/H ₂ mixture produced by coal gasification containing about 62% CO, 37.5% H ₂ and 0.5% of CH ₄ and N ₂ . The membranes have the properties of those used in the earlier example calculations.

[0080] The first membrane unit (601) reduces the feed gas (600) concentration from 38% H ₂ to about 16% H ₂ . The second membrane unit (602) then reduces the H ₂ concentration to about 1%, producing a CO concentrate (608) containing more than 98% CO.

[0081] The first membrane unit permeate (609) contains about 83% H ₂ . Two-thirds of this gas is sent to the PSA unit (604) which produces a pure H ₂ product gas (614) and a low pressure tail gas (615). This tail gas is mixed with the second portion, one-third, of the first unit permeate to produce a syngas stream (612) with the H^CO ratio of 2 required for methanol production. The permeate from the second membrane unit (616) is compressed and recycled to the feed gas (600). [0082] The CO product gas (608) contains about 89% of the CO content of the original feed. The hydrogen product gas (614) contains about 61% of the H ₂ content of the original feed.

[0083] The flows and compositions of the key process streams are shown in Table 6.

Table 6