KR101190461B1 - Production of High Purity and Ultra-High Purity Gas - Google Patents

Abstract

Passing the feed stream through the carbon monoxide adsorbent 33 before passing it through the supported metal catalyst 34 removes traces of carbon monoxide and optionally hydrogen from the gas phase feed stream. The present invention saves significant capital and operating costs over existing methods.

Adsorbent, Carbon Monoxide, High Purity, Feed Airflow, Metal Catalyst

Description

Production of High Purity and Ultra-High Purity Gases {Production of High Purity and Ultra-High Purity Gas}

preference

This application claims the priority of US Patent Application No. 60/384612, filed May 31, 2002 and Serial No. 60/385148, filed June 3, 2002. The contents of each application are incorporated herein by reference.

The present invention relates to the removal of impurities from the feed stream, and more particularly to the removal of hydrogen (H ₂ ) and carbon monoxide (CO) from the feed stream.

Typical concentrations of carbon monoxide (CO) and hydrogen (H ₂ ) in the air are on the order of 1 ppm CO and 1 ppm H ₂ , but may be present in the air at concentrations of about 50 ppm and 10 ppm or less, respectively. Conventional cryogenic distillation processes used to produce ultra-high purity (UHP) nitrogen (N ₂ ) do not remove hydrogen but only a small amount of CO. Unless removed by other optional means, these molecules will contaminate the product nitrogen at concentrations up to about 2.5 times its concentration in the feed air. The electronics industry requires very high purity nitrogen products (typically less than 5 ppb CO and less than 5 ppb H ₂ ), so CO and H ₂ must be removed from the feed air.

The air also contains other contaminants such as water (H ₂ O), carbon dioxide (CO ₂ ) and hydrocarbons. In cold parts of distillation separation processes (such as heat exchangers and separation columns), water and CO ₂ may solidify and clog the heat exchanger or other parts of the distillation column.

Acetylene and other hydrocarbons in the air also pose a potential problem because they can accumulate liquid oxygen (O ₂ ) and create an explosion hazard. It is therefore desirable to remove these impurities prior to cryogenic distillation of air.

The prepurification of air can be performed using pressure swing adsorption (PSA), temperature swing adsorption (TSA) or a combination thereof (TSA / PSA) involving one adsorbent or a plurality of adsorbents. If two or more adsorbents are used, the adsorbents may be arranged in discrete layers, as mixtures, composites or combinations thereof. Impurities such as H ₂ O and CO ₂ are typically removed from air using one or more adsorbent layers in a combined TSA / PSA process. The first layer of activated alumina or zeolite is commonly used for water removal and the second layer of zeolite, such as 13X molecular sieve, is used for CO ₂ removal. Prior art, such as US Pat. No. 4,711,645, teaches the use of various adsorbents and methods for removing CO ₂ and water vapor from air. And the adsorbent is allowed to pass through because of ineffective, the CO and H ₂ a distillation apparatus in order to remove the CO and H _2.

There are three main strategies in the prior art for the removal of CO and / or H ₂ from air to produce UHP nitrogen: removal upstream of the prepurifier adsorbent, removal in the prepurifier adsorbent using oxidation catalyst and cryogenic Removal from nitrogen product after air separation.

In a first approach, CO and H ₂ are typically removed by high temperature catalytic oxidation on a supported inert metal or hopcalite catalyst upstream of the prepurifier bed. Oxidation product of CO and H _2, i.e., CO ₂ and H ₂ O are removed with the periphery of the CO ₂ and H ₂ O in the prepurifier bed (FC Venet, etc., "Understand the Key Issues for High Purity Nitrogen Production," Chem. Eng. Prog., Pp 78-85, Jan. 1993). This approach requires significant power and additional capital that is substantially added to the cost of the process.

U.S. Patent No. 5,656,557 discloses compressed air before entering a catalyst tower containing a palladium (Pd) and / or platinum (Pt) supported catalyst for converting CO, H ₂ and hydrocarbons to H ₂ O and CO ₂ . A method of further heating to ° C is disclosed. The treated air is then cooled to 5 ° C. to 10 ° C. before entering the prepurifier where H ₂ O and CO ₂ are removed. Some of the emissions from the prepurifier may be used as air containing less than 1 ppm total impurities, while the remaining air is separated to cryogenic temperatures to produce N ₂ and O ₂ .

French patent FR 2 739 304 1) contacts hot steam compressed from a compressor with a bed of CO oxidation catalyst; 2) cool the intermediate air stream obtained to ambient temperature; 3) adsorbing CO ₂ and H ₂ O by contacting the CO-free air stream with an adsorbent; 4) Describes a method for removing CO and H ₂ from air, involving contacting the resulting air stream with a H ₂ trapping adsorbent. The CO catalyst may be a copper (Cu) or group Pt metal supported on alumina, silica or zeolite. The H ₂ trapping adsorbent may be osmium (Os), iridium (Ir), Pd, ruthenium (Ru), rhodium (Rh) or Pt supported on alumina, silica or zeolite.

US Pat. No. 6,074,621 describes a method similar to FR 2 739 304 except for the cooling step after the CO oxidation catalyst.

US Pat. No. 5,693,302 discloses a process for removing CO and H ₂ from composite gas by passing over particles containing gold and Pd supported by TiO ₂ .

US Pat. No. 5,662,873 describes a similar process using a catalyst consisting of at least one element and silver from the Pt group supported on alumina, silica or zeolites.

The second technique used in the prior art is an ambient temperature process for CO and H ₂ removal from air.

U.S. Patent 5,110,569 discloses CO and optionally hydrogen from air by 1) removing water, 2) catalytically oxidizing CO to CO ₂ and optionally H ₂ to H ₂ O, and 3) removing the oxidation product. Disclosed is a method for removal. The oxidation catalyst for CO can be a mixture of manganese and copper oxides such as hopcalite or carulite. Nickel oxide is also referred to as an effective CO catalyst. The oxidation catalyst for H ₂ is typically supported palladium.

U.S. Patent No. 5,238,670 discloses 1) water from air until its moisture content is less than 150 ppm, and 2) Cu, Ru, Rh, deposited by ion-exchange or impregnation on zeolite, alumina or silica; A method for removing CO and / or H ₂ from air at a temperature between 0 ° C. and 50 ° C. by removing CO and H ₂ on a bed of particles containing at least one metal element selected from Pd, Os, Ir and Pt. Initiate.

European patent application EP 0 454 531 describes a similar method which proposes to remove both H ₂ O and CO ₂ before the impregnated bed of particles. Trace amounts of H ₂ O and CO ₂ are removed downstream of the impregnated particle bed.

US Pat. No. 6,048,509 discloses 1) Pd or Pt and iron (Fe), cobalt (Co), nickel (Ni), manganese (Mn), Cu, chromium (Cr), tin (Sn) supported on large pore alumina. Contacting the gas with a CO catalyst composed of at least one selected from the group consisting of lead (Pb) and cerium (Ce); 2) contacting the CO-free gas with an adsorbent for water removal; 3) the gas that is obtained is brought into contact with the CO ₂ sorbent for the removal of CO _2, optionally; 4) surrounding the gas, passing H ₂ O, CO ₂ , CO and optionally H ₂ containing air, in contacting the gas with an H ₂ catalyst consisting of Pt or Pd supported on activated alumina or zeolite Disclosed is a method for removing CO and H ₂ from air at temperature. Water formed during the final stage of the hydrogen oxidation is either adsorbed on the catalyst support is H _2, H ₂ mixed with the catalyst and is removed by a physical or H ₂ O adsorbent located downstream thereof.

U. S. Patent No. 6,093, 379 discloses a process in which a prepurifier bed having a first layer of moisture adsorbent and a second layer of CO ₂ adsorbent acting at ambient temperature is enlarged by a third layer of catalyst / adsorbent to remove both CO and H ₂ . It is described. The third layer is exposed to substantially H ₂ O-free and CO ₂ -free air at ambient temperature. The dual catalyst / adsorbent layer is located in the third layer at the most downstream end of the prepurifier bed. The dual catalyst / adsorbent layer oxidizes CO, adsorbs the resulting CO ₂ , and chemically adsorbs H ₂ . The dual catalyst / adsorbent is a precious metal such as Pd on a support having a zero point charge (ZPC) of greater than 8.

U. S. Patent No. 6,511, 640 discloses a method in which the prepurifier is arranged to contain a series of layers that comprise a series of layers starting with the adsorbent at the feed inlet for H ₂ O removal. The second layer, followed by an oxidation catalyst for converting CO into CO ₂ , is followed by an adsorbent for CO ₂ removal. An oxidation catalyst is placed in the next layer to change H ₂ to H ₂ O, while the last layer is used to adsorb H ₂ O. The disclosed CO catalyst is hopcalite, while the H ₂ catalyst is Pd supported on activated alumina.

A third common method for producing UHP N ₂ in the prior art is to treat the cryogenically separated N ₂ product to remove H ₂ , CO, O ₂ and other contaminants that pass through the prepurifier and air separation unit.

US 4,579,723 discloses the use of Ni or Cu supported catalysts or getters to oxidize contaminants to CO ₂ and H ₂ O, which are then removed from the adsorbent.

European patent EP 0 835 687 teaches the regeneration of the catalyst bed with a hot N ₂ purge.

Adsorption of CO is known in the prior art mainly to recover CO in large quantities of separation, for example when the concentration (partial pressure) of CO is relatively high (typically = 1%) and CO is a more strongly adsorbed component in the gas mixture. Has been applied. Cuprous compounds in the cation form in zeolites or dispersed on a porous support have been widely applied for the recovery of CO from gas mixtures containing CO and N ₂ , methane (CH ₄ ), H ₂ and / or CO ₂ . Materials containing copper in one oxidation state (represented by Cu ⁺ , Cu (I) or cuprous) exhibit high CO adsorption performance, while adsorbents containing Cu (II) are not. Adsorbents are typically synthesized, treated or modified with Cu (II) compounds, and then exposed to reducing agents such as H ₂ at elevated temperatures to convert Cu (II) to Cu (I).

Xie et al. ("Highly Efficient Adsorbent for Separation of Carbon Monoxide," Fundamentals of Adsorption, Proc. IV ^th Int. Conf. On Fundamentals of Adsorption, Kyoto, May 17-22, 1992, pp. 737-741) An adsorbent formed by mixing CuCl dry powder at elevated temperature and dispersing on a zeolite support is described. High-purity CO separated with high recovery is shown for feed streams containing 9.0% CO / 91% N ₂ and 30.7% CO / 65.3% H ₂ /4% CH ₄ .

US Pat. No. 4,917,711 discloses adsorbents and methods using supported CuCl. U. S. Patent No. 5,531, 809 discloses a VSA process using CuCl dispersed on alumina to recover CO from the synthesis gas leaving the steam-methane reforming unit.

GK Pierce ("The Industrial Practice of Adsorption," in: Separation of Gases, 5 ^th BOC Priestley Conf., Birmingham, UK Sept. 19-21, 1989, Spec. Publ. No. 80, Royal Soc. Of Chemistry, Cambridge , 1990) describe the use of Cu (I) Y zeolites to recover CO from CO / N ₂ and CO / H ₂ feed streams containing percent levels of CO.

U.S. Patent 4,473,276 discloses Cu (I) Y and Cu-Mordenite with other exchanged zeolites with silica to alumina ratio (SiO ₂ / Al ₂ O ₃ ) = 10 for the recovery of CO. Doing.

U.S. Pat. No. 4,019,879 discloses H ₂ O and / or CO ₂ using Cu ⁺ containing zeolites having 20 = SiO ₂ / Al ₂ O ₃ = 200, for example ZSM-5, -8, -11, etc. Recovery of CO is started from the air stream it contains.

Another class of adsorbents with the potential for CO adsorption is materials containing silver (Ag ⁺ ). Y. Huang ("Adsorption in AgX and AgY Zeolites by Carbon Monoxide and Other Simple Molecules," J. Catal., 32, pp. 482-491, 1974) is the lowest temperature (0 ° C and 25 ° C) isotherms For CO and N ₂ isotherms for AgX and AgY zeolites at partial pressures as low as 0.1 to 1.0 Torr. The adsorption capacity for CO is significantly greater than that of N ₂ at 25 ° C. and 100 Torr.

US Pat. No. 4,743,276 discloses mordenite, A, Y and X type zeolites exchanged with varying amounts of Ag for mass separation (recovery) of CO from purification and petroleum-chemical exhaust-gases.                 

US Pat. No. 4,019,880 discloses H ₂ O and / or CO ₂ using Ag exchanged zeolites having 20 = SiO ₂ / Al ₂ O ₃ = 200, for example ZSM-5, -8, -11 and the like. It also relates to recovering CO from the air stream it contains. This invention applies to feed streams containing at least 10 ppm of CO at temperatures between 0 ° C and 300 ° C. The claimed process results in CO-depleted emissions, for example air.

U.S. Patent No. 4,944,273 discloses 1 = Si / Al = 100, especially for adsorption of nitrogen oxides (NO _x ) and CO as part of the NO _x and CO sensors in automotive exhaust gases. Zeolites doped with Ag, Pt or Ru are disclosed.

U.S. Patent No. 3,789,106 discloses that mordenite impregnated with copper is effective at removing CO from H ₂ at a CO partial pressure of less than 3 mmHg. This effect was measured by applying the adsorbent to a CO concentration of at least 100 ppm and measuring the capacity at saturation.

The above prior art regarding the adsorption of CO is rarely mentioned with respect to the purification of mixed air streams, especially those containing less than 10 ppm of CO in O ₂ and N ₂ .

Object of the Invention

It is therefore an object of the present invention to provide an improved process for removing trace amounts of CO and H ₂ , H ₂ O, CO ₂ , hydrocarbons and N ₂ O from a feed stream, preferably air.

Summary of the Invention

The present invention relates to the removal of CO and optionally H ₂ from air and / or other gases or gas mixtures using a combination of adsorptive separation and catalytic conversion, which is unexpected in capital and power costs compared to conventional techniques. Provide savings.

In one preferred embodiment, at least 90% of each of CO ₂ and H ₂ O is first removed from the feed gas (preferably air) to produce a CO ₂ and H ₂ O removed gas. In a particularly preferred embodiment, the CO ₂ and H ₂ O removed gas contains less than 1.0 ppm (more preferably less than 0.25 ppm) of CO ₂ and less than 1.0 ppm (more preferably less than 0.10 ppm) of H ₂ 0. It contains. Subsequently, a substantial amount of CO is removed from the CO ₂ and H ₂ O stripped gas through the use of a CO adsorbent to contain less than 100 ppb of CO in preferred embodiments and less than 5 ppb of CO in more preferred embodiments. Prepare a CO-degassed gas. Optionally, H ₂ and any remaining CO can then be removed using a catalyst. Because of this unique combination of catalyst and adsorption, the process of the present invention provides surprisingly superior CO and H ₂ removal efficiencies compared to the prior art processes.

Preferred apparatus for the practice of the present invention includes an adsorption apparatus for removing CO from a feed stream containing CO in an amount of less than 50 ppm. The apparatus comprises at least one adsorption vessel containing a CO adsorption layer, a CO adsorbent having a ΔCO working capacity of at least 0.01 mmol / g, wherein

a) when the feed stream further contains at least one gas selected from the group consisting of nitrogen, He, Ne, Ar, Xe, Kr, CH ₄ and mixtures thereof, the adsorbent is ion exchanged with a Group IB element And is preferably selected from the group consisting of AgX zeolite, Ag-mordenite, Cu-clinoptilolite, AgA zeolite and AgY zeolite;

b) if the feed stream further contains at least one gas selected from the group consisting of oxygen, air and mixtures thereof, the apparatus is preferably an air prepurifier and the adsorbent is less than 20 SiO ₂ / It is a zeolite having an Al ₂ O ₃ molar ratio and is ion-exchanged with Ag ⁺ or Au ⁺ and is preferably selected from the group consisting of AgX zeolite, Ag-mordenite, AgA zeolite and AgY zeolite.

In another embodiment, the apparatus comprises two or more adsorption vessels selected from the group consisting of vertical flow vessels, horizontal flow vessels, side flow vessels or radial flow vessels.

In another embodiment, the adsorption device further contains an adsorbent, preferably at least one alumina or NaX zeolite, which is selective for adsorption of water upstream of the CO adsorbent bed.

In another embodiment, the apparatus is a catalyst layer, preferably alumina, silica, natural or synthetic zeolite, titanium dioxide, magnesium, and calcium oxide for the catalytic oxidation to H ₂ O in the H ₂ at the downstream of the CO adsorbent layer At least one metal supported on a substrate selected from the group consisting of Os, Ir, Pt, Ru, Rh, Pd, Fe, Co, Ni, Cu, Ag, Au, Zn, Sn, Mn, Cr, Pb, Ce It further contains a metal supported catalyst comprising a.

In another embodiment, the apparatus further contains an auxiliary adsorbent, preferably at least one alumina or NaX for removal of water downstream of the catalyst bed.

In another embodiment, the ΔCO working capacity is at least 0.03 mmol / g.

In another embodiment, the device further contains at least one additional adsorbent for adsorption of at least one of H ₂ , CO ₂ , N ₂ O and hydrocarbons, wherein the additional adsorbent is alumina, silica gel, clinoptil It is preferably selected from the group consisting of olites, zeolites, composites thereof and mixtures thereof, and located downstream of the catalyst bed.

In another embodiment, the CO adsorbent has a ΔCO / ΔN ₂ separation titer of at least 1 × 10 ⁻³ .

In another embodiment, the CO adsorbent is AgX with at least 50%, preferably 100%, of its cations bound to Ag.

Preferred methods for the practice of the present invention include a removal method for removing CO from a feed air stream containing CO in an amount of less than 50 ppm, or even less than 1.0 ppm or 0.5 ppm. The process, which is preferably an air prepurification process, involves contacting the feed stream with a CO adsorbent having a ΔCO working capacity of at least 0.01 mmol / g, preferably at least 0.03 mmol / g, to produce a CO-removed air stream that can be recovered. Includes; From here                 

a) when the feed stream further contains at least one gas selected from the group consisting of nitrogen, He, Ne, Ar, Xe, Kr, H ₂ , CH ₄ and mixtures thereof, the adsorbent is group IB Zeolites ion-exchanged with an element and are preferably selected from the group consisting of AgX zeolite, Ag-mordenite, Cu-clinoptilolite, AgA zeolite and AgY zeolite;

b) when the feed stream further contains at least one gas selected from the group consisting of oxygen, air and mixtures thereof, the adsorbent is a zeolite having a SiO ₂ / Al ₂ O ₃ molar ratio of less than 20, Ion-exchanged with Ag ⁺ or Au ⁺ and is preferably selected from the group consisting of AgX zeolite, Ag-mordenite, AgA zeolite and AgY zeolite.

In another embodiment, the method includes recovering the CO-depleted air stream, wherein CO is CO-depleted air stream at a concentration of less than 100 ppb, preferably less than 5 ppb, most preferably less than 1 ppb. Exists in

In another embodiment, the feed gas further comprises water (H ₂ O) and the method comprises contacting the feed air stream with a moisture selective adsorbent, preferably alumina or NaX, located upstream of the CO adsorbent. It includes more.

In yet another embodiment, the feed gas further comprises hydrogen, the method comprising: removing a group CO to oxidation catalyst the H ₂ to H ₂ O H ₂ is removed, producing a the H ₂ O is rich gas The at least one metal Os, Ir, Pt, Ru, Rh, supported on a substrate selected from the group consisting of alumina, silica, natural or synthetic zeolites, titanium dioxide, magnesium oxide and calcium oxide, Contacting with a catalyst layer, which is a metal supported catalyst comprising Pd, Fe, Co, Ni, Cu, Ag, Au, Zn, Sn, Mn, Cr, Pb, Ce. The degassed gas preferably contains less than 100 ppb hydrogen, more preferably less than 5 ppb hydrogen.

In another embodiment, the process comprises contacting a catalyst prepared H ₂ O-enriched gas with a water removal adsorbent (preferably alumina or NaX) located downstream of the catalyst bed to produce CO, H ₂ and H _2. The method further includes preparing a gas from which O has been removed.

In another embodiment, the CO adsorbent has a ΔCO / ΔN ₂ separation titer of at least 1 × 10 ⁻³ , preferably at least 1 × 10 ⁻² .

In another embodiment, the method comprises

a) at least one adsorbent layer upstream of the CO adsorbent for at least one adsorption of H ₂ O and CO ₂ ;

b) the catalyst bed for catalytic conversion of H ₂ O of H ₂ in the downstream of the CO adsorbent layer; And

c) consisting of alumina, silica gel, clinoptilolite, zeolites, composites thereof and mixtures thereof for adsorption of at least one of H ₂ O, CO ₂ , N ₂ O and hydrocarbons downstream of the catalyst bed It further comprises passing through at least one further adsorbent preferably selected from the group.

In another alternative embodiment, the method is selected from the group consisting of pressure swing adsorption, temperature swing adsorption, or a combination thereof, wherein the method occurs in an adsorbent vessel selected from a vertical flow vessel, a horizontal flow vessel, or a radial flow vessel.

In a preferred embodiment, the adsorption step of the process is carried out at a temperature of 0 to 50 ° C.

In a preferred method, the CO adsorbent is AgX with at least 50%, more preferably 100% of its cations bound to Ag.

In a preferred embodiment, the feed stream contains air and the CO-removed stream is sent to the cryogenic distillation column.

In another alternative embodiment, the CO partial pressure of the feed stream is less than 0.1 mmHg, or even less than 0.005 mmHg.

Another alternative embodiment relates to a removal method for removing CO from a feed stream containing less than 50 ppm of CO and hydrogen, the method comprising a CO adsorbent having a ΔCO working capacity of at least 0.01 mmol / g. Contacting the feed air stream to produce a CO-depleted air stream; The adsorbent here is a zeolite exchanged with a Group IB element.

In a preferred embodiment, the present invention comprises an adsorption apparatus for removing CO from a feed stream containing less than 50 ppm of CO and hydrogen, the apparatus comprising a CO adsorbent having a ΔCO working capacity of at least 0.01 mmol / g. At least one adsorption vessel containing the bed; The adsorbent is a zeolite exchanged with an IB group element.

In a particularly preferred embodiment, the invention comprises an air prepurification adsorption apparatus for removing CO from an air supply air stream containing CO in an amount of less than 50 ppm, the apparatus having a ΔCO working capacity of at least 0.01 mmol / g. At least one adsorption vessel containing a CO adsorbent layer; The adsorbent is a zeolite (preferably AgX) which is ion-exchanged with Ag ⁺ or Au ⁺ with a SiO ₂ / Al ₂ O ₃ molar ratio of less than 20.

In a preferred embodiment, the air prepurification apparatus further comprises at least one adsorbent layer upstream of the CO adsorbent for adsorbing at least one of H ₂ O and CO ₂ .

In a preferred embodiment, pre-treating the catalyst layer is a catalyst for converting H ₂ in the downstream of the CO adsorbent layer in H ₂ O; And one or more additional adsorbents for adsorption of at least one of H ₂ O, CO ₂ , N ₂ O and hydrocarbons, wherein the additional adsorbent is downstream of the catalyst bed.

In another embodiment, the present invention includes a method for removing CO from a feed air stream containing less than 50 ppm of CO and air, wherein the process is performed at least 0.01 mmol / g in an adsorbent vessel. Contacting a CO adsorbent with a capacity to produce a CO-removed air stream, wherein the adsorbent has a SiO ₂ / Al ₂ O ₃ molar ratio of less than 20 and is ion-exchanged with Ag ⁺ or Au ⁺ , preferably It is AgX. The method comprising a) in order to adsorb at least one of H ₂ O and CO ₂ produced at least one of the supply air stream by passing the supply air stream H ₂ O and or CO ₂ removal on the adsorbent layer is placed upstream of the CO adsorbents and b) one or more additional adsorbents for the adsorption of one or more of H ₂ O, CO ₂ , N ₂ O and hydrocarbons and a catalyst layer downstream of the CO adsorbent layer for catalytic conversion of H ₂ to an H ₂ O layer And further passing the feed stream above, wherein the additional adsorbent comprises a feed stream from which one or more of CO, H ₂ , and H ₂ O, CO ₂ , N ₂ O, and hydrocarbons have been removed. Downstream of the catalyst bed for production, optionally the removed feed air stream is sent to a cryogenic distillation column for separation of air.

Combinations of any of the above embodiments are considered to be within the scope of the present invention.

Other objects, features and advantages will appear to those skilled in the art from the following description of the preferred embodiment (s) and from the accompanying drawings:

1 is a schematic diagram of a breakthrough test apparatus.

2 is a graph of the CO and H ₂ breakthrough curves for the AgX adsorbents in Table II.

3-6 are schematic diagrams of a preferred adsorbent arrangement in an adsorbent vessel / bed.

7 is a schematic diagram of a prepurification device useful for practicing the present invention.

The present invention relates to the removal of CO and optionally H ₂ from a feed gas containing ppm levels of CO and ppm levels of H ₂ . The present invention provides a feed stream having CO of 1 ppm or less; In particular, it is particularly applicable for removing CO in feed streams having a low CO partial pressure of less than 0.1 mmHg or even less than 0.005 mmHg.

The feed gas is typically air and may contain other contaminants including at least one of hydrocarbons such as CO ₂ , H ₂ O, N ₂ O and acetylene. Although the present invention may be applied to the removal of CO and optionally hydrogen impurities after cryogenic air distillation, the removal of these components typically takes place in a prepurification process prior to air separation by cryogenic means.

Feed streams containing less than 1 ppm or even less than 0.5 ppm of CO and less than 1 ppm or even less than 0.5 ppm of H ₂ can also be used, although the present invention is capable of approximately 50 ppm of CO and 10 ppm of H ₂ in the feed. It can be carried out with a feed air stream containing an amount of. The practice of the present invention using such a feed stream can produce product air streams containing less than 5 ppb of CO and less than 5 ppb of H ₂ , but an air stream containing 1 ppb of CO and 1 ppb of H ₂ will be obtained. It may be.

The use of H ₂ the catalyst after the use of a CO adsorbent in accordance with a preferred embodiment of the invention reduces the H ₂ catalyst requirements and eliminates the need of high temperature activation, the H ₂ catalyst by placing the H ₂ catalyst in the clean end of the adsorbent Maximum protection against loss of activity and lower capital and operating costs compared to conventional prepurification techniques. The result is that emissions from the adsorbent / adsorption process of the present invention (e.g., feed to an cryogenic distillation column (or "cold box") of an air separation unit (ASU)) are less than 5 ppb CO and 5 ppb. It contains less than H ₂ .

In one preferred embodiment of the present invention, the adsorbent bed from the feed stream containing CO, H ₂ , CO ₂ and H ₂ O impurities is first substantially all H ₂ O and optionally all prior to removing CO on the CO adsorbent. Arranged to remove CO ₂ . The CO ₂ , H ₂ O and CO degassed gas is then passed through a catalyst bed to remove any remaining CO and H ₂ . The purified gas is then sent to ASU to produce UHP nitrogen. For the purposes of the present disclosure, “high purity” gas means a gas having an individual contaminant level of less than about 100 ppb (parts in billions), depending on the intended application of the gas, and “ultra-high purity” (UHP) By gas is meant a gas having an individual contaminant level of less than about 10 ppb, preferably less than about 5 ppb, most preferably less than 1 ppb.

That is,

1) Emissions containing less than 5 ppb CO by removal of CO by adsorption from a feed stream containing trace amounts of CO on a suitable adsorbent in the presence of high N ₂ and / or O ₂ concentrations at temperatures ranging from 0 to 50 ° C. To prepare a gas;

2) optionally, two steps of removing trace amounts of H ₂ from the feed stream in the presence of high N ₂ and / or O ₂ concentrations by combination of adsorption and / or absorption and / or catalysts at temperatures ranging from 0 to 50 ° C. The process produces a gas from which H ₂ and CO are removed.

According to one embodiment of the invention, an appropriate CO adsorbent is selected that has a sufficient ΔCO working capacity to remove most or all CO from the feed stream. The ideal CO adsorbent will have a good CO load (ΔCO or working capacity) in the presence of high N ₂ and / or O ₂ concentrations. It is also desirable to have a low N ₂ load in order to maximize CO selectivity and minimize the cooling effect of N ₂ desorption during decompression. In a preferred embodiment, the adsorbent also has a high ΔCO / ΔN ₂ working separation titer at local steam conditions.

One method for evaluating adsorbent performance is by measuring the working capacity of each major adsorbate, ie N ₂ and CO. This is the methodology applied in the present invention. The separation titer α defined below is a preferred method for evaluating adsorbent efficiency. The methodology is discussed in detail in US Pat. No. 6,152,991, the contents of which are incorporated herein by reference.

The separation titer α is defined by the following equation:                 

The separation titer α is defined here as the ratio of the working capacity. The molecule of the equation is the working capacity of CO, which is equal to the difference in the load w between adsorption and desorption conditions. Adsorption and desorption conditions are characterized by composition y , pressure p and temperature T. In TSA air prepurification, the maximum regeneration temperature may vary from about 100 ° C to about 350 ° C. As a result, the adsorbate (particularly atmospheric gas) is expected to be completely thermally desorbed from the adsorbent. Under such conditions, equation (1) can be simplified as follows.

When the contaminants are removed in the shallow adsorbent bed of the TSA process, there is a significant resistance to mass transfer and the selectivity is redefined by Equation 3:

The molecule of equation (3) represents the working capacity of the adsorbent for contaminants. m _in represents the molar feed flowed into the bed, y _in and y _out are the feed and discharge mole fractions of the minor components, respectively. w _s is the mass of the adsorbent and t _b is the breakthrough time corresponding to the predetermined concentration. The denominator is the equilibrium capacity of the main components under conditions at the end of the adsorption step, ie assuming complete desorption of all components. The application of equation (3) is preferred when the length of the mass transfer region occupies a significant portion (eg, at least about 10%) of the total depth of the adsorbent bed.

The process differs because the working capacity is determined at the partial pressure of each individual component at the relevant process conditions, for example CO and N ₂ partial pressures in the feed air for prepurification, typically below 0.1 mmHg and above 1000 mmHg, respectively. It is preferred over the method. Coadsorption effects are also introduced into the determination of the load. Since the feed concentration of N ₂ in the air prepurification is overwhelming compared to CO, the cosorption effect of CO on N ₂ is negligible. Thus, the denominators of Equations 2 and 3 can be obtained directly from the measured pure-component N ₂ isotherms.

In contrast, the cosorption of N ₂ will have a very significant effect on the adsorption of CO. In the case of CO, where exactly low concentrations of pure-component isotherm data can or can be obtained, Equation 2 can be applied to assess the equilibrium working capacity and selectivity. However, it is desirable to determine the working capacity for CO directly under N ₂ coadsorption conditions and with the mass transfer effects involved. This can be done using a breakthrough test method known to those skilled in the art. The breakthrough test provides saturation and the equilibrium dose of the component at any defined breakthrough level, for example breakthrough capacity and time at 100 ppb. That is, the ΔCO working capacity and selectivity are preferably defined using equation (3). Since co-adsorption of nitrogen is introduced into the ΔCO working capacity, ΔCO (molecule of formula (3)) determined by the breakthrough method is a major factor in selecting an appropriate CO adsorbent for the purposes of the present invention.

Those skilled in the art will adsorbent which satisfies the criteria of acceptable operation remarkable feature of the CO is well recognized that act much more advantageous from the rest of the gas mixture to be less strongly adsorbed component than a large amount compared to the N ₂ N _2. Examples of such CO-containing gas mixtures include those having at least one of He, H ₂ , Ar, Ne, Xe, Kr, O ₂ and CH ₄ , and the present invention relates to CO from the gases or gas mixtures. It can also be applied to separation.

To evaluate the potential efficacy of the CO adsorbent at typical conditions of the air separation process, a CO and / or H ₂ breakthrough test was performed.

Breakthrough tests were performed at an inlet gas flow rate of 7.9 bara (114.7 psia), 10 ° C. or 27 ° C. and about 21 slpm (78.7 mol / m ² s) using an adsorption column of 5.9 cm or 22.9 cm in length. Changes in the conditions are shown in the examples. Feed conditions are representative of the conditions at the inlet of the air prepurifier for a typical cryogenic air separation plant. In some tests a complete breakthrough curve was generated for CO and / or H ₂ , while only a partial breakthrough was measured when using high volumes of material. Initial breakthrough was at 100 ppb CO and 20 ppb H ₂ . Next, the initial breakthrough and / or saturation load (capacity in mmol / g) was determined by mass balance as shown in the molecule of equation (3). The CO and H ₂ loads determined at the initial breakthrough represent a dynamic or working capacity, which introduces both cosorption and mass transfer effects.

The breakthrough test was performed in the following manner using the apparatus whose main elements are shown in the schematic diagram of FIG. 1. From the cylinder (3,4) containing the mixture of contaminant gas in N ₂ , the gas in question is metered in (flow rate of 0-5 standard liters per minute) through the flow regulator (7,8), there was allowed to mix with the high-purity diluent N _2, or helium (He) of the gas in the mixer (9). In some cases, the diluent was synthetic air, where O ₂ from source (2) was combined with N ₂ from source (1). The diluent was fed through flow regulators 5 and 6 at a defined flow rate (eg 0-30 slpm) to obtain a feed concentration of the desired contaminant (s). The mixed gas in question was then fed through a heat exchange loop to test bed 10 containing the adsorbent. The test bed 10 and the heat exchange loop were both located in a thermal bath (not shown), the temperature of which was controlled through a water cooler and thermocouple. Temperature control systems are known in the art and typically consist of a heat exchange loop, a thermal bath, a water cooler and a thermocouple. Any number of changes in the temperature control system can be effectively applied to maintain the test bed at a constant temperature as is readily familiar to those skilled in the art.

Gas pressure was controlled via a control valve / pressure regulator 11 during the test. A portion of the effluent was sent through valve 12 to one or more analyzers 13 to monitor the breakthrough concentration of CO and / or H ₂ as a function of time. A Servox 4100 gas purity analyzer equipped with both CO and CO ₂ sensors was used in Example 1. In Examples 2-4 a Trace Analysis (RGA5) analyzer was used for the trace CO and H ₂ measurements.

All zeolites were obtained from commercial sources (identified below), but some were further modified by ion exchange by the inventors. For the examples of the present invention the mordenite and Y zeolites were extrudate, the X and A zeolites were beads, and the clinoptilolite and chabazite were granular.

Highly exchanged zeolites were prepared using a laboratory ion-exchange column. In the column process, zeolite was charged inside the column and the ion exchange solution was pumped upward through the bed at a flow rate of 0.5 ml / min. The column was heated to between 70 and 90 ° C. to facilitate exchange. At least 10-fold excess solution was used to ensure high levels of ion exchange were obtained. The product was collected by filtration and washed thoroughly with deionized water.

Powder X-ray diffraction and inductively coupled plasma (ICP) chemistry analysis was used to verify the integrity of the sample and measure ion exchange levels. For Ag, Cu and zinc (Zn) samples, 0.1 molar (M) solution of each nitrate was used. For sodium (Na) exchange, 1.0 M sodium chloride solution was used. Zeolites were typically in their Na exchanged form before being exchanged with Ag, Cu or Zn. For clinoptilolite, a TSM-140 sample was selected and first exchanged with Na before Ag, Cu or Zn were introduced. Zeolite 13X HP, 13X APG and Na Mordenite were obtained from UOP and Zeolite 4A was purchased from Aldrich.

For Ag, which is a low exchange level on 13X HP (eg less than 85%), a batch process was used. In the batch process, the zeolite is immersed in a fixed volume of solution as opposed to regularly pumping fresh solution into the column. Silver nitrate solution with a concentration of 0.025 to 0.1 M was used.

The zeolite was stirred at 50-90 ° C. in silver nitrate solution. The concentration of the solution was chosen to reflect the level of exchange desired. After stirring for 6-8 hours at this temperature, the solution was freshly fed up to three times in total, or the sample was collected by filtration and washed sufficiently with deionized water. Fresh feed of solution increased the exchange level. Powder X-ray diffraction and ICP chemical analysis were performed to investigate the crystallinity and ion exchange level of the product sample.

Ion exchange methods of zeolites are known to those skilled in the art. The aforementioned columns and batch methods do not limit the invention in any way. Different methods can be used in accordance with the present invention to achieve the desired level of exchange for cations and zeolites.

The invention will now be described by way of the following comparative examples. These examples are not intended to limit the scope of the invention in any way.

Example 1 (outside the scope of the present invention)

Natural adsorbents (clinoptilolite and carbazite) were obtained from Steelhead Specialty Minerals, WA. Synthetic zeolites were obtained from various sources: Zeolist (HZSM5), Zeochem (CaX) and UOP (13X, LiX, 5AMG). All adsorbents were thermally activated for about 16 hours under 350 ° C., 1.0 bara pressure and N ₂ purge prior to each test. After regeneration, the adsorbent was allowed to cool to 27 ° C. Adsorbents of this group exhibit cross-sections of known adsorbent properties such as Si / Al, micropore channel opening size, zeolite structure type, number and type of cations, etc., which are known to perform adsorptive properties. Breakthrough tests were performed under the conditions described above to saturate with 1.0 ppm of CO in N ₂ or He. Breakthrough time was determined at 100 ppb CO. The results are summarized in Table I.

absorbent Flow
slpm T ℃ Carriers ⁽¹⁾ y _CO
ppm Bed
cm t _b ⁽²⁾
minute X _CO ⁽³⁾
mmol / g Clinoptilolite TSM-140 20.4 27 N ₂ 1.0 22.9 <2 1.4 x 10 ^-5 21 27 He 1.0 22.9 30 1.2 x 10 ^-3 Cabazite TSM-300 21 27 N ₂ 1.0 22.9 <2 2.2 x 10 ^-5 21 27 He 1.0 22.9 114 4.5 x 10 ^-3 LiX
(SiO ₂ / Al ₂ O ₃ = 2.0) 21 27 N ₂ 1.0 22.9 <2 2.5 x 10 ^-5 5A MG 5.5 10 N ₂ 1.0 22.9 <6 3.1 x 10 ^-5 CaX (Z100) 5.5 10 N ₂ 1.0 22.9 <4 3.9 x 10 ^-5 HZSM5 (CBV3024E) 21 27 air 2.0 5.9 <2 - (1) 0.25 ppm typical H ₂ concentration in carrier gas
(2) breakthrough measured at y _CO = 100 ppb
(3) CO equilibrium saturation capacity

The competitive effect of N ₂ on CO saturation capacity (X _CO ) is that the X _CO for the CO / N ₂ feed is reduced by a titer of at least 100 compared to the CO / He feed for small pore natural zeolites. It is obvious. The breakthrough time of CO in N ₂ was reduced by a titer of at least 15-50 compared to CO in He. The higher partial pressure of N ₂ relative to CO is a decisive factor in determining adsorbent efficacy in CO removal. All adsorbents in Table I only have a few minutes CO / N ₂ breakthrough time. Breakthrough times can be extended with longer adsorbent beds, but none of these adsorbents are likely to provide a practical solution for air prepurifiers with adsorption circulation times of more than one hour.

N ₂ isotherms were measured at 27 ° C. using known gravity balance methods for carbazite and clinoptilolite. The N ₂ capacity from the isotherm and the CO saturation capacity from Table I were combined in Equation 2 to obtain the equilibrium separation titer (ΔCO / ΔN ₂ ) <2 × 10 ⁻⁵ for the adsorbents of Table I.

Example 2

Several zeolites containing exchanged Ag, Cu and Zn cations were tested as described above to assess working CO capacity. AgX (P / N 38,228-0) was obtained from Aldrich, while clinoptilolite (TSM-140) and mordenite (large-pore from UOP) were exchanged in small-scale laboratory columns. The degree of exchange was measured by inductive pair plasma (ICP) analysis. The exchanged sample was air-dried and then activated overnight at 350 ° C. in a dry N ₂ purge. Approx. 21 slpm (78.7 mol / m ² s) using a 5.9 cm length column packed with adsorbent and applied at 2.0 ° C. and 7.9 bar in 2.0 ppm CO in synthetic air (79% N ₂ /21% O ₂ ) The breakthrough test was performed while flowing at a speed of).

For each test, the amount of H _{2 in} the feed is shown in Table II along with the results of the breakthrough test. Breakthrough time (t _b ) and working capacity (X _CO ) were measured at a CO breakthrough concentration of 100 ppb.

absorbent % exchange size
US mesh y _H2
ppm X _CO
@y _CO = 100 ppb
mmol / g t _b
hr % CO removed
@y _CO = 100 ppb AgX 100 10 x 14 3.0 0.052 5.4 98.8 Ag-Mor 89 1.8 mm ^* 3.0 0.034 3.2 98.3 Cu-
Clinoptilolite 100 8 x 14 0.5 0.037 3.0 99.3 Zn-
Clinoptilolite 86 8 x 14 0.6 - <0.2 - Ag-
Clinoptilolite 89 8 x 14 0.4 - <0.5 - * Extrudate diameter

AgX, Ag-mordenite and Cu-clinoptilolite all exhibit CO working capacities in excess of 0.03 mmol / g and consequently meet the criteria of the present invention. Zn-clinoptilolite and Ag-clinoptilolite had a working capacity of less than 0.01 mmol / g and consequently are outside the scope of the present invention.

Compared conservatively to the CO saturation capacity of Table 1, the CO working capacity (including dynamic effects) of Table II for these adsorbents is over 1000 times greater than that of Table 1. The breakthrough time is also several hours for a relatively short bed, ie the resulting CO working capacity allows the adsorbent to be easily integrated with the usual prepurifier circulation and with minimal additional adsorbent.

Cu-, Zn- and Ag-exchanged clinoptilolite appear to have little or no removal capacity for H ₂ . H ₂ breaks almost immediately in AgX and Ag-mordenite, but some degree of stagnation of H ₂ is apparent as shown in FIG. 2.

2 provides breakthrough records for both CO and H ₂ using AgX (Aldrich P / N 38,228-0). Although a CO breakthrough concentration of 100 ppb was chosen for the purpose of adsorbent evaluation, one skilled in the art will appreciate that the amount of CO impurities in the product can be adjusted by selecting a shorter adsorption step or increasing the amount of adsorbent.

After each breakthrough test, the beds are regenerated at 200 ° C. in air for 3.0 hours followed by 3.0 hours at ambient temperature purge in dry N ₂ , at a flow rate of 2.2 slpm and a pressure of about 1.7 bar (25 psia). The result of FIG. 2 is for the tenth breakthrough of the bed. Even in the presence of strong H ₂ reducing agents-there was no noticeable drop in the CO removal capacity due to circulation. Desorption was monitored by purging the bed with N ₂ at 27 ° C. and 50 ° C. prior to standard regeneration after the sixth breakthrough. Only CO was detected in the emissions, and more than 87% of the adsorbed CO was desorbed. The remaining CO was desorbed during normal regeneration.

N ₂ isotherms were measured for AgX (Aldrich P / N 38,228-0) at 0 ° C. and 27 ° C. as described above. The N ₂ dose at an N ₂ partial pressure of 6.25 bar at 27 ° C. was determined to be 0.81 mmol / g. The corresponding CO partial pressure for the test streams in Table II is 1.2 × 10 ⁻² mmHg. Using the results of Equation (3) and Table II, the work separation titer ΔCO / ΔN ₂ = 6.4 × 10 ⁻² .

For Cu-clinoptilolite, the results in Table II are for the first breakthrough after activation. After regeneration according to the conditions recited above, the performance dropped significantly at the second breakthrough. The performance of Table II recovered after re-activation at 350 ° C. in N ₂ . Regeneration at elevated temperatures in air is believed to result in oxidation of Cu ⁺ to Cu ⁺⁺ . That is, reactivation in N ₂ returned Cu to the lower oxidation state Cu ⁺ . Therefore, the adsorbent is a substantially O ₂ or after separation of the N ₂ to the extremely low temperature does not exist at all - can be more effectively applied to other alternative embodiments of the present invention to apply the concepts of the invention in the tablet.

Example 3

It was purchased from commercial AgX (Ag400B3), Ag- to know who a sample of nitro (Ag900E16) and AgY (Ag600E16) ^* Chem ^{(C * CHEM, A Division of} Molecular Products, Inc., Lafayette, CO) is. AgX was also prepared by exchanging 13 XHP zeolites with Ag at various Ag-exchange levels (indicated by AgX (HP)). AgA was prepared by exchanging 4A zeolite with Ag. The materials and corresponding Ag exchange levels are shown in Table III along with breakthrough time and CO working capacity.

absorbent Ag-exchange
% size
US mesh X _CO
@y _CO = 100 ppb
mmol / g t _b
hr % CO
remove AgX
(Ag400B3) 100 10 x 18 0.051 5.5 99.0 AgX (HP) 95 10 x 18 0.13 13.3 99.1 AgA 94 8 x 12 0.019 2.2 98.4 Ag-Mor 40 2.0 mm ^* 0.0096 0.85 97.9 AgY 61 1.7 mm ^* 0.018 1.4 98.8 * Extrudate diameter

For all tests the breakthrough test was performed under the same conditions as in Example 2 except that 3.0 ppm H ₂ was present in the feed. The results in Table III are for the first breakthrough after activation. AgX (Ag400B3) performance matches well with that of AgX in Table II. AgX (HP) appears to have significantly improved CO working capacity over commercially available AgX, but much of this advantage is lost after several breakthrough / regeneration cycles. After four cycles, the breakthrough time for AgX (HP) was reduced to 7.3 hours and the CO working capacity dropped to 0.073 mmol / g. However, the amount of decay was rapidly reduced and a larger final working capacity than that of AgX (Ag400B3) was presented. Such an improvement may be a result of the difference in the substrate zeolite starting material (13XHP), the macropore geometry and / or the final state of the exchanged zeolite and / or the position of the silver cation in the zeolite.

Different Ag-exchange levels varying from 10% to 100% were prepared using 13XHP base material. CO working capacity was found to be nearly directly proportional to Ag-exchange levels. That is, higher Ag-exchange levels are preferred for adsorption of CO, with 100% exchange being most preferred.                 

Commercial AgY and Ag-mordenite adsorbents performed much better than conventional zeolites for CO adsorption but were less effective than AgX. Significantly different CO working capacities for Ag-mordenite in Tables II and III resulted in higher Ag-exchange levels and the ΔCO working capacity of less than 0.01 mmol / g of the laboratory-produced Ag-mordenite in Table II. Have lower levels of exchanged material (in Table III) and are therefore outside the scope of the present invention.

Example 4

In this example, the bed consists of a 7.6 cm layer of AgX of Example 2 followed by a 15.2 cm layer of palladium (Pd / Al ₂ O ₃ ) oxidation catalyst supported on porous alumina, which is typical of the prior art. The second bed consists of the same overall length (22.9 cm) using only Pd / Al ₂ O ₃ oxidation catalyst. An oxidation catalyst (E221, 0.5 wt% Pd supported on the surface of activated alumina beads) was obtained from Degussa Corporation. Breakthrough tests were performed under the same conditions as in Example 3. The results are compared to Table IV. Breakthrough times of H ₂ were recorded at 1 ppb, 5 ppb and 20 ppb.

Bed t _b
(H ₂ )
hr t _b
(CO)
hr t _b
(CO ₂ )
hr X _H2
mmol / g
@ y _H2 = 20 ppb y _H2 = 1 ppb y _H2 = 5 ppb y _H2 = 20 ppb 7.6cm AgX + 15.2cm
0.5% PdAl ₂ 0 ₃ 21.3 23.2 26.5 > 30.7 17.3 0.218 22.9 cm
0.5% PdAl ₂ O ₃ 3.9 4.9 13.0 > 21.5 3.0 0.072

Surprisingly, replacing 1/3 of the catalyst bed with AgX for CO adsorption resulted in 2 to 5 times greater H ₂ breakthrough times than when the catalyst was used alone. At 20 ppb H ₂ breakthrough, the effective capacity of the catalyst was more than tripled when CO was removed prior to H ₂ oxidation. Over the test time, ie 30.7 hours (hr) for layered beds and 21.5 hr for catalyst only beds, there was essentially no CO breakthrough in the bed. Finally, the CO ₂ breakthrough time (y _CO2 = 100 ppb) was more than five times greater when CO was adsorbed on AgX. As is clear from the previous information, the use of AgX to replace some of the CO / H ₂ oxidation catalysts has unexpected advantages in purifying the airflow of CO and H ₂ while minimizing CO ₂ as oxidation byproducts. .

Based on the above examples, preferred CO adsorbents for the practice of the invention have a ΔCO working capacity of at least 0.01 mmol / g, preferably at least 0.03 mmol / g. In a more preferred embodiment, the CO adsorbent has a ΔCO / ΔN ₂ separation titer α as given in equation 3, which is at least 1 × 10 ⁻³ , preferably greater than 1 × 10 ⁻² . For small pore zeolites (eg, native zeolite clinoptilolite, carbazite, etc.), the pore diameter of the pore or zeolite “window” should be greater than the pore diameter of the CO molecule (0.376 nm). Zeolites exchanged with Group IB cations (Cu, Ag, Au (gold)) to obtain oxidation state +1 are preferred adsorbents for CO for the application of the present invention. Ag-exchanged zeolites (> 50% exchange) are preferred, and highly exchanged AgX (> 85% exchange) is most preferred. Although not disclosed in the above examples, gold is thought to be useful in the practice of the present invention because of its similar chemical structure to Cu and Ag. Group IB metals that act as charge balanced cations in the exchanged zeolites provide improved CO adsorption capacity, and zeolites requiring higher numbers of charge balanced cations are preferred in the practice of the present invention. Such zeolites are characterized by a molar ratio of SiO ₂ / Al ₂ O ₃ of less than 20.

Using the process of the present invention, the TSA adsorbents and systems provide a concentration of CO and, optionally, at least one of H ₂ O, CO ₂ and H ₂ from the feed stream introduced, to levels of up to 100 ppb in the exhaust or product gas, Or even reduced to less than 5 ppb. Preferably, the vessel is used in a TSA prepurifier for air feed airflow. One design of such a container is described below with reference to FIG. 3. Arrows (see also Figures 4-6) indicate the direction of gas flow through the adsorbent bed / vessel. A TSA prepurifier system introducing such a vessel is described below with reference to FIG. 7.

Returning to FIG. 3, the container 30 is visible. Vessel 30 includes a first layer of H ₂ O adsorbent 31, such as alumina, silica gel or molecular sieves or mixtures thereof to remove substantially all H ₂ O entering the vessel. A second layer 32 of CO ₂ adsorbent such as 13X (NaX) or 5A or mixtures thereof is used to remove substantially all CO ₂ . The CO ₂ adsorbent layer can also remove any residual moisture remaining from the H ₂ O adsorbent layer. The third layer 33 of CO adsorbent is located downstream of the CO ₂ adsorbent. (The term "downstream" means closer to the discharge or product end of the adsorbent vessel.)

Substantially H ₂ O-free and CO ₂ -free air flows enter the CO adsorbent bed. The CO adsorbent layer is adapted to produce product gas containing less than 100 ppb of CO by removing at least 50%, preferably at least 95%, of CO in the feed, and most preferably at least 99.8% of CO in the feed. Can be designed to produce product streams containing less than 5 ppb of CO. A feed air stream that is substantially free of H ₂ O, CO ₂ and CO enters the catalyst bed 34. The catalyst uses adsorption and / or a combination of adsorption and oxidation to remove H ₂ and any remaining small amounts of CO. One or more optional adsorbent layers 35 may be located downstream of the CO adsorbent and catalyst beds to remove any CO ₂ and H ₂ O oxidation products formed but not adsorbed to the catalyst. Optional layer (s) may also be selected to remove hydrocarbons, N ₂ O and / or other trace contaminants.

4-6 show another alternative embodiment for integrating a CO adsorbent bed into an adsorbent vessel for use in gas separation, such as prepurification. Of course, combinations of these embodiments are also within the scope of the present invention.

The thickness of the adsorbent and catalyst bed required varies depending on the process conditions experienced by the adsorbent during the separation process. These can vary widely from system to process. The depth of the bed shown in FIGS. 3-6 is not intended to mean or suggest any particular or relative amount of adsorbent and / or catalyst. The most important process conditions are the molar flow rate, circulation time, feed temperature and pressure and gas composition of the air. For example, one skilled in the art of designing a prepurifier may devise an adsorbent and catalyst bed for each prepurifier depending on the adsorbent / catalyst properties and prepurifier process conditions. These methods are known as Ruthven, Principles of Adsorption and Adsorption Processes , 1984).

3 has been described above, and FIGS. 4-6 will be described below.

Referring to FIG. 4, an embodiment is intended to use a vessel 40 for the removal of H ₂ O (in layer 41), CO ₂ (in layer 42) and CO (in layer 43). In this embodiment no catalyst layer is required for H ₂ removal.

Referring to FIG. 5, a vessel 50 is shown wherein the H ₂ O and CO ₂ adsorbents are combined into a single layer 51 as a single adsorbent, as a mixture of different adsorbents, or through the use of a complex adsorbent, and then This is followed by a CO adsorbent layer 52 and a catalyst layer 53 for H ₂ removal. Additional adsorbent layers (similar to the function of layer 35 shown in FIG. 3), such as layer 35 for cleaning and / or removal of other components (eg hydrocarbons), may also be used but are not shown.

Referring to FIG. 6, the vessel 6 contains a CO adsorbent layer 62 located prior to the CO ₂ adsorbent layer 63, both downstream of the H ₂ O adsorbent layer 61. If removal of H ₂ is required, an additional catalyst layer (not shown) may be used after the CO layer. Additional layers may be added for cleaning and removal of other components (eg hydrocarbons).

The catalyst used for the removal of H ₂ and remaining CO is a supported metal catalyst. Metal Os, Ir, Pt, Ru, Rh, Pd, Fe, Co, Ni, Cu, Ag, Au, Zn, Sn, Mn, Cr, Pb, Ce One kind of Ce Alumina, silica, natural or synthetic It may be deposited on a support selected from zeolite, titanium dioxide, magnesium oxide or calcium oxide. At least one of the metals is deposited on the support by techniques known in the art. Most preferred catalysts for the present invention are Pt and / or Pd supported on alumina.

The optional adsorbent layer described above may be at least one of alumina, silica gel or zeolite. Zeolite layers are preferred. Most preferably 13X (most preferably 13X APG available from UOP, Des Plaines, Illinois USA) is used in the layer where CO ₂ and H ₂ O oxidation products should be removed. The optional H ₂ O and CO ₂ removal layers may be located downstream of the catalyst bed or physically mixed with the catalyst bed. The adsorbent selected for the selective bed depends on the contaminants to be removed; For example, to remove trace amounts of N ₂ O and CO ₂ as described in PCT Application No. 02/40591, "Method of Removing N ₂ O from Gas Flow," which is pending together and assigned together, Low light; And / or NaY, alumina or selesov (complex of NaY / alumina available from SELEXORB, ALCOA, USA) for removal of hydrocarbons.

The present invention provides the advantage that existing prepurifiers can be easily retrofitted by placing the CO adsorbent and optional catalyst at the downstream end of the prepurifier. The invention also applies to the production of "high purity" products when ultra-high purity products are not required, i.e. when the standard purity of existing air separation facilities is not sufficient but electronic grade UHP is not required. It may be. In such applications, any of the embodiments, methods, and arrangements of the present invention can be used to purify feed mixtures and remove CO and / or H ₂ contaminants to less than 100 ppb each.

The method is preferably performed in a cyclic method such as pressure swing adsorption (PSA), temperature swing adsorption (TSA), vacuum swing adsorption (VSA), or a combination thereof. The process of the invention can be carried out in a single or a plurality of adsorption vessels operating in a cyclical process comprising at least the steps of adsorption and regeneration. The adsorption step is carried out at a pressure in the range from 1.0 to 25 bar, preferably from about 3 to 15 bar. The temperature range during the adsorption step is -70 ° C to 80 ° C. If a PSA process is used, the pressure during the regeneration step is in the range of about 0.20 to 5.0 bar, preferably 1.0 to 2.0 bar. In the case of the TSA process, regeneration is typically carried out at temperatures in the range of about 50 ° C to 400 ° C, preferably 100 ° C to 300 ° C.                 

One possible method is described herein with reference to FIG. 7. The supply air is compressed in the compressor 70 and cooled by the cooling means 71 before at least contaminants H ₂ O, CO ₂ and CO enter one of the two adsorbents 76 and 77 from which they are removed from the air. Each of the adsorbents 76 and 77 has the same adsorbent bed arrangement, which may be, for example, as described with reference to FIGS. 3-6 above. The purified air exits the adsorbent and then enters an air separation unit (ASU) where it is separated at cryogenic temperatures into its main components, N ₂ and O ₂ . In the special design of the ASU, Ar, Kr and Xe can also be separated and recovered from the air. One of the beds adsorbs contaminants from the air, while the other is regenerated using purge gas. A dry, contaminant-free purge gas may be supplied from the ASU or from a waste stream or from an independent source to desorb the adsorbed adsorbate to regenerate the adsorbent and prepare it for the next adsorption step of the circulation. The purge gas may be a mixture of N ₂ , O ₂ , N ₂ and O ₂ , air or any dry inert gas. In the case of heat swing adsorption (TSA), the purge gas is first heated in heater 82 before passing through the adsorbent in a direction opposite to the feed flow of the adsorption step. TSA circulation may include pressure fluctuations. If only pressure swing adsorption (PSA) is used, there is no heater.

The operation of a typical TSA cycle is now described with reference to FIG. 7 for one adsorbent 76. Those skilled in the art will appreciate that other adsorbent vessels 77 will operate in the same circulation only from the phase with the first adsorbent in such a way that the purified air is continuously available to the ASU. This work from phase cycling is indicated by reference to the numbers in parentheses.

Feed air is introduced into the compressor 70 to which it is pressurized. The heat of compression is removed in the cooling means 71, for example a mechanical cooler or a combination of direct contact post-cooler and evaporative cooler. The compressed, cold and H ₂ O-saturated feed air stream then enters adsorbent 76 (77). When adsorbent vessel 76 (77) is pressurized, valve 72 (73) is opened and valves 74 (75), 78 (79), and 80 (81) are closed. Once the adsorption pressure is reached, valve 78 (79) is opened and the purified product is directed to the ASU for cryogenic air separation. When the adsorbent 76 (77) has completed the adsorption step, close valves 78 (79) and 72 (73) and open valve 74 (75) to blow adsorbent 76 (77) to a lower pressure, typically near ambient pressure. send. Once depressurized, valve 80 (81) is opened and heated purge gas is introduced into the product end of adsorbent 76 (77). At some point during the purge circulation, the heater is turned off to allow the purge gas to cool the adsorbent near the feed temperature.

Those skilled in the art will also appreciate that there are many variations of such a typical circulation in which the above technique merely illustrates an example of a prepurifier cycle and which can be used with the present invention. For example, the PSA is used alone, where the heater 82 and cooling means 71 may be removed. Pressurization may be carried out by product gas, feed gas or a combination of both. Bed-to-bed equalization may be used, where a compounding step is introduced in which the freshly regenerated bed is sent on line in the adsorption step with another adsorbent close to the completion of the adsorption step. Such a compounding step eliminates pressure disturbances due to bed switching and also serves to minimize any thermal disturbances caused when the regenerated bed is not completely cooled to the feed temperature. In addition, the present invention may be practiced with a prepurifier cycle that is not limited to two adsorbent beds. The method of the present invention can be applied in horizontal, vertical or radial flow vessels.

The method of regeneration depends on the type of circulation process. In the case of the TSA process, regeneration of the adsorbent bed is carried out by passing heated gas through the bed in the opposite direction. Using a thermal pulse method, the cold purge step follows the hot purge step. The heated regeneration gas may also be provided at reduced pressure (relative to the feed) such that the combined TSA / PSA process is performed. The reduced pressure may be above or below ambient pressure. In cryogenic air separation processes, regeneration gas is typically taken from product or N ₂ or O ₂ waste streams.

In some cases, passing the inert or weakly adsorbed purge gas through the bed in the opposite direction may make the adsorbent bed cleaner. In the PSA process, the purge step generally follows the opposite pressure reduction step. In the case of a single vessel system, the purge gas may be introduced from the storage vessel, while in the case of a multi-bed system the purge gas may be obtained from another adsorbent in the adsorption bed.

The adsorption system can have more steps than the two basic steps of adsorption and desorption. For example, top to top equalization or bottom to bottom equalization can be used to conserve energy and increase recovery.

In a preferred embodiment of the present invention as applied to air prepurification, substantially all of the water vapor and substantially all of the CO ₂ pass over at least one layer of activated alumina or zeolite, before passing through the air stream through the CO adsorbent layer, Or from the air by a plurality of layers of activated alumina and zeolite. Optionally, the CO selective adsorbent layer can be extended and used to remove some or all of the CO ₂ from the air.

Otherwise, in the adsorption vessel, the first layer of adsorbent may be used to remove water vapor, and the next layer of a mixture of CO selective adsorbent and 13X (or other zeolites) may be used to remove both CO ₂ and CO from air. Can be used. Such adsorbent mixtures may consist of different adsorbents bound together in the form of physically separated adsorbents or composites.

In applications requiring only a small amount of CO adsorbent (eg low CO concentration in the feed), the CO adsorbent is mixed with another adsorbent (13X, etc.) to form a thicker layer of mixed adsorbent rather than a very thin layer of CO adsorbent alone. It may be advantageous to do so. Thicker mixed layers are advantageous in terms of ease of installation and less critical resistance to the overall layer thickness.

The process of the present invention can be used to remove one or both of CO and H ₂ , from air, N ₂ and other weakly adsorbable bulk gases. Weaker adsorbent gases for which CO is to be purified can be identified initially from the electronic properties of such gases. Using the criteria of CO / N ₂ selectivity and CO working capacity defined above to select the appropriate adsorbent, the system is characterized by N ₂ , helium (He), neon (Ne), H ₂ , xenon (Xe), krypton ( It is guaranteed to work as well or better in removing CO from weaker adsorbed gases such as Kr), argon (Ar), O ₂ and methane (CH ₄ ) and other gases having similar properties.

In addition, the preferred application of the present invention is the application in prepurification prior to cryogenic air separation, but the present invention as described herein may also be applied to remove CO and / or hydrogen present in nitrogen gas after cryogenic distillation.

The size of the H ₂ O and CO ₂ removal layers upstream of the CO adsorbent (the term “upstream” means closer to the feed end of the adsorbent vessel) means that the size of H ₂ O and CO ₂ in the gas to be purified It will depend on the concentration. An inert gas, such as N ₂ , may contain ppm levels of O ₂ if it contains oxygen. Thus, the mechanism for CO and H ₂ removal in CO adsorbents and catalysts will be a combination of adsorption / chemical adsorption and absorption. If CO is removed only by adsorption, the optional clean layer can be omitted, as shown in FIG. 4 above. In the absence of oxygen, the oxidation products CO ₂ and H ₂ O will not be formed, so the optional clean layer shown in FIG. 3 will not be necessary. The process of the invention may also be used for the subsequent purification of N ₂ in cryogenic facilities.

As pointed out above, the adsorbent bed / container used in the process of the present invention can have various arrangements, such as vertical flow bed, horizontal flow bed or radial flow bed, pressure swing adsorption, temperature swing adsorption, vacuum swing adsorption In a manner or a combination thereof.

In this method, the adsorbent can be shaped by a series of methods into various geometric shapes such as beads, granules and extrudates. This will entail adding the binder to the zeolite powder in a manner known to those skilled in the art. The binder may also be necessary to match the strength of the adsorbent. The type and morphology of the binder are known and the present invention does not place any limitation on the type and percentage of binder in the adsorbent.

CO adsorbents may also potentially adsorb some hydrocarbons and nitrogen oxides from air. To ensure complete removal of the hydrocarbon, the CO adsorbent may be physically mixed with the hydrocarbon selective adsorbent.

The process presented herein can be used for the cleaning of any gas containing pollutant levels of CO alone or in combination with H ₂ , H ₂ O, CO ₂ , hydrocarbons and N ₂ O.

The term "comprising" as used herein means "including but not limited to", ie one or more other features, integers, steps, when specifying the presence of a stated feature, integer, step or component, It does not exclude the presence or addition of ingredients or their groups.

Specific forms of the invention are shown in one or more figures for convenience only, each of which can be combined with other forms according to the invention. Other alternative embodiments will be appreciated by those skilled in the art, which are intended to be included within the scope of the claims.

Claims (58)

At least one adsorption vessel containing a CO adsorbent layer having a ΔCO working capacity of at least 0.01 mmol / g, wherein a) when the feed stream further contains at least one gas selected from the group consisting of nitrogen, He, Ne, Ar, Xe, Kr, CH ₄ and mixtures thereof, the adsorbent is a cation of a group IB element. Ion exchanged zeolites; b) when the feed stream further contains at least one gas selected from the group consisting of oxygen, air and mixtures thereof, the adsorbent is a zeolite having a SiO ₂ / Al ₂ O ₃ molar ratio of greater than 0 and less than 20 Ion-exchanged with, Ag ⁺ or Au ⁺ , Adsorption apparatus for removing CO from a feed stream containing CO in an amount of less than 50 ppm. The adsorption apparatus of claim 1, further comprising a moisture selective adsorbent layer, wherein the moisture selective adsorbent layer is upstream of the CO adsorbent layer. The adsorption apparatus of claim 1, further comprising a catalyst layer for catalytic oxidation of H ₂ to H ₂ O, wherein the catalyst layer is downstream of the CO adsorbent layer. 4. The adsorption apparatus of claim 3, further comprising an auxiliary adsorbent for removal of water, wherein said auxiliary adsorbent is downstream of said catalyst bed. The method of claim 1, a) at least one adsorbent layer upstream of said CO adsorbent for adsorption of at least one of H ₂ O and CO ₂ ; b) the catalyst bed for catalytic conversion to H ₂ O of H ₂ in the downstream of the CO adsorbent layer; And c) an apparatus further comprising at least one additional adsorbent for adsorbing at least one of H ₂ O, CO ₂ , N ₂ O and hydrocarbons downstream of the catalyst bed. Contacting the feed air stream with a CO adsorbent having a ΔCO working capacity of at least 0.01 mmol / g to produce a CO removed air stream; From here a) when the feed stream further contains at least one gas selected from the group consisting of nitrogen, He, Ne, Ar, Xe, Kr, H ₂ , CH ₄ and mixtures thereof, the adsorbent is group IB Zeolites ion exchanged with cations of elements; b) when the feed stream further contains at least one gas selected from the group consisting of oxygen, air and mixtures thereof, the adsorbent has a zeolite having a SiO ₂ / Al ₂ O ₃ molar ratio of greater than 0 and less than 20 Is ion-exchanged with Ag ⁺ or Au ⁺ , A method for removing CO from a feed stream containing CO in an amount less than 50 ppm. 7. The method of claim 6, further comprising recovering the CO stripped air stream wherein CO is present at a concentration of less than 100 ppb. 8. The method of claim 6 or 7, wherein the feed stream further comprises water (H ₂ O) and further comprising contacting the feed stream with a moisture selective adsorbent located upstream of the CO adsorbent. The method of claim 6 or 7, wherein the feed stream is further comprising: a hydrogen, the CO that was removed feed stream is brought into contact with the catalyst layer to an oxidation catalyst for H ₂ to H ₂ O H ₂ is removed and H ₂ O is Producing a enriched gas, wherein the catalyst bed is located downstream of the CO adsorbent bed. 10. The method of claim 9, further comprising contacting the H ₂ O-rich gas with a water removal adsorbent, wherein the water removal adsorbent layer is located downstream of the catalyst bed to form CO, H ₂ and H ₂ O. To produce a gas removed. 8. The method of claim 6 or 7, wherein said ΔCO working capacity is at least 0.03 mmol / g. 10. The method of claim 9, wherein said degassed gas contains less than 100 ppb hydrogen. 8. The method of claim 6 or 7, wherein the feed stream further contains at least one gas selected from the group consisting of nitrogen, He, Ne, Ar, Xe, Kr, H ₂ , CH ₄ and mixtures thereof. Wherein said CO adsorbent is selected from the group consisting of AgX, Ag-Mor, Cu-clinoptilolite, AgA zeolite and AgY zeolite. 8. The CO adsorbent according to claim 6 or 7, wherein the CO adsorbent further comprises at least one gas selected from the group consisting of air, oxygen and mixtures thereof, wherein the CO adsorbent is selected from AgX, Ag-Mor, AgA zeolite and AgY. Method selected from the group consisting of zeolites. 8. The process according to claim 6 or 7, wherein the feed air stream contains air and the CO-removed air stream is sent to the cryogenic distillation column. 8. The method of claim 6 or 7, further comprising recovering said CO removed air stream wherein CO is present at a concentration of less than 1 ppb. 8. The process according to claim 6 or 7, wherein the partial pressure of CO in said feed stream is less than 0.1 mm Hg. Contacting the feed air stream with a CO adsorbent having a ΔCO working capacity of at least 0.01 mmol / g to produce a CO depleted air stream, the adsorbent being in an amount less than 50 ppm, which is a zeolite ion exchanged with a cation of a Group IB element. A method for removing CO from a feed stream containing CO and hydrogen. At least one adsorption vessel containing a layer of CO adsorbent having a ΔCO working capacity of at least 0.01 mmol / g, wherein the adsorbent is a zeolite ion exchanged with a cation of a Group IB element, less than 50 ppm of CO, and hydrogen Adsorption apparatus for removing CO from the feed stream containing the. At least one adsorption vessel containing a CO adsorbent layer having a ΔCO working capacity of at least 0.01 mmol / g, wherein the adsorbent is a zeolite having a SiO ₂ / Al ₂ O ₃ molar ratio of greater than 0 and less than 20 and is an Ag ⁺ or Au ⁺ Air purifying adsorption apparatus for removing CO from an air feed air stream containing less than 50 ppm of CO, which is ion-exchanged. The apparatus of claim 20 further comprising at least one adsorbent layer upstream of the CO adsorbent for adsorption of at least one of H ₂ O and CO ₂ . Claim 20 or according to 21, wherein the CO for conversion catalyst for H ₂ in the downstream of the adsorbent layer with H ₂ O catalyst layer, and H ₂ O, CO _2, N ₂ O and a hydrocarbon of 1 for the absorption or more species Further comprising at least one additional adsorbent, said further adsorbent downstream of said catalyst bed. Contacting the feed air stream with a CO adsorbent having a ΔCO working capacity of at least 0.01 mmol / g in the adsorbent vessel to produce a CO-depleted air stream, the adsorbent having a SiO ₂ / Al ₂ O ₃ molar ratio of greater than 0 and less than 20 A method for removing CO in a feed stream containing less than 50 ppm of CO, and air, which is zeolite and ion-exchanged with Ag ⁺ or Au ⁺ . 24. The method of claim 23, wherein generating a feed gas stream at least one passing over the adsorbent layer, H ₂ O or CO ₂ removal in the feed gas stream to the CO adsorbent upstream for adsorbing at least one of H ₂ O and CO ₂ Further comprising. 25. The catalyst bed of claim 23 or 24, wherein the feed stream is downstream of the CO adsorbent bed for catalytic conversion of H ₂ to H ₂ O bed, and 1 of H ₂ O, CO ₂ , N ₂ O and hydrocarbons. Passing over at least one additional adsorbent for adsorption of at least one species, wherein the additional adsorbent is downstream of the catalyst bed to form CO, H ₂ , and H ₂ O, CO ₂ , N ₂ O and Wherein at least one of the hydrocarbons is removed to produce a feed stream. delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete

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