PROCESSES TO SELECTIVELY REMOVE AND RECOVER A CONTAMINATING GAS FROM A GAS OF ORIGIN CONTAINING CONT

专利摘要:
processes for selectively removing and recovering a contaminating gas from a source gas containing contaminants, and process for removing a contaminating gas from a source gas containing contaminants. this invention relates to processes for the selective removal of contaminants from effluent gases. more particularly, various embodiments of the present invention relate to the selective removal and recovery of sulfur dioxide from effluent gases in a regenerative sulfur dioxide absorption / desorption process that achieves favorable energy efficiency. energy is recovered from a wet stripper overflow gas flow produced in the desorption cycle by indirect heat transfer from the stripper gas to a cooling medium and used to generate steam for use in stripping contaminants from the absorption liquor. the absorption zone, optionally, can be cooled to increase the capacity of the absorption medium for the absorption of a polluting gas, thus decreasing the volume of the absorption medium and the absorption liquor enriched with contaminants that must be pumped, handled, heated and cooled in the absorption / desorption cycle.
公开号:BR112014027446B1
申请号:R112014027446-0
申请日:2013-05-02
公开日:2021-04-06
发明作者:Ernesto Vera- Castaneda
申请人:Mecs, Inc.；
IPC主号:

专利说明:

[0001] [001] This invention relates to processes for the selective removal of contaminants from effluent gases. More specifically, several embodiments of the present invention concern the selective removal and recovery of sulfur dioxide from effluent gases in a regenerative absorption / desorption of sulfur dioxide that achieves favorable energy efficiency. The recovery schemes of the invention are applicable for the removal and recovery of other acidic gases such as hydrogen sulfide, carbon dioxide and hydrogen chloride, as well as other contaminating gases such as ammonia. BACKGROUND OF THE INVENTION.
[0002] [002] Gaseous effluents containing contaminating gases are produced by a variety of operations. For example, sulfur dioxide is generated in a number of chemical and metallurgical operations, including sulfur-burning sulfuric acid processes, spent sulfuric acid plants, passed sulfuric acid plants, roasting or smelting metal sulphide ores and concentrates and combustion of carbon fuels containing sulfuric (eg flue gases from coal-fired power plants). Carbon fuels play a significant role in generating electricity, providing heating energy and transport fuels. Most carbon fuels contain sulfur which, when burned, turns into sulfur dioxide. The sulfur dioxide emitted contributes to a wide range of environmental and health problems. As emerging economies expand, their energy demands are growing rapidly, and as carbon fuels with lower sulfur levels are depleted, more and more oil and coal reserves with increasing sulfur levels are used, causing the emission of sulfur dioxide.
[0003] [003] There are also increasing regulatory pressures to reduce sulfur dioxide emissions around the world. The most commonly used method to remove sulfur dioxide is through absorption or adsorption techniques. A common approach is to contact sulfur dioxide with an aqueous stream that contains a low-cost base. The sulfur dioxide dissolves in water, forming sulfurous acid (H2SO3), which in turn reacts with the base to form a salt. Common bases are sodium hydroxide, sodium carbonate and lime (calcium hydroxide, Ca (OH) 2). The pH starts at about 9 and is reduced to about 6 after the reaction with sulfur dioxide. A wet one-stage scrubber usually removes 95% of the sulfur dioxide. Wet scrubbers and, similarly, dry scrubbers require capital investment, variable costs due to the consumption of lime and solids disposal, and consume energy and utilities to operate such sulfur dioxide removal systems.
[0004] [004] Instead of reacting with base-type lime, sulfur dioxide and effluent gases can be recovered and sold as a refined sulfur dioxide product, used as part of the feed gas for a contact sulfuric acid plant , and recovered as sulfuric acid, and / or oil to meet the growing global demand from the fertilizer industry, or to feed a Claus plant for the preparation of elemental sulfur. In addition to addressing the health and environmental problems associated with sulfur dioxide emissions, this approach recovers the sulfur values of coal and other sulfur-containing carbon fuels. However, these gas flows often have a relatively low concentration of sulfur dioxide and a high concentration of water vapor. Where the concentration of sulfur dioxide in the gas fed to a sulfuric acid plant is less than about 4 to 5 percent by volume, problems may arise with respect to both the water balance and the energy balance at the acid plant . More specifically, the material balance of a conventional sulfuric acid plant requires that the molar ratio of H2O / SO2 in the gas stream containing sulfur dioxide fed to the plant be greater than the molar ratio of H2O / SO3 in the acid produced. If the desired acid concentration of the product is 98.5 percent or more, this ratio cannot be greater than about 1.08 in the gas flow containing sulfur dioxide being fed to the plant. As generated, effluent gases from metallurgical processes and combustion gases from sulfurous carbon fuels often have a water vapor content well above the 1.08 ratio, which cannot be sufficiently reduced by cooling the gas without significant energy expenditures. and capital. In addition, if the sulfur dioxide gas strength of the source gases is less than about 4 to 5 percent by volume, it may not be sufficient for the autothermal operation of the catalytic converter. That is, the heat from the conversion of sulfur dioxide to sulfur trioxide may not be large enough to heat the gases entering the catalyst operating at temperature and, consequently, heat from an external source must be supplied. This, in turn, also increases both operating costs and the capital requirement for the installation of sulfuric acid.
[0005] [005] The strength of sulfur dioxide in gaseous effluents can be improved by selective absorption of sulfur dioxide in a suitable solvent and, subsequently, removing the absorbed sulfur dioxide to produce a regenerated solvent and a gas enriched in dioxide content. sulfur. A variety of aqueous solutions and organic solvents and solutions have been used in regenerative sulfur dioxide absorption / desorption processes. For example, aqueous solutions of alkali metals (eg sodium sulphite / bisulphite solution), amines (eg alkanolamines, tetrahydroxyethylalkylenediamines, etc.), amine salts and various salts of organic acids have been used as dioxide absorbers regenerable sulfur.
[0006] [006] Inorganic aqueous buffer solutions are also effective in absorbing sulfur dioxide. Fung et al. (2000) provides data on the solubility of sulfur dioxide for a 1 Molar solution of phosphoric acid and sodium carbonate in the proportion of about 1.57 Na / PO4 as a function of temperature. The data are for the virgin mixture and the mixture where 1,000 ppm of adipic acid is added to increase the solubility of sulfur dioxide. Fung et al. they also indicate that, when brought to the boiling temperature, 95% and 65% of the sulfur dioxide is removed, respectively, for the virgin mixture and the mixture containing adipic acid. Calculations about the pH of the solution show that the pH changes from 6 to about 3 once the sulfur dioxide is absorbed. As with organic solvents, there is a slight reaction of sulfur dioxide with oxygen, forming sulfur trioxide. Although this reaction is very limited and, when Na2CO3 is used, it is further inhibited by its reaction with free radicals formed during oxidation, the sulfur trioxide formed leads to the formation of sodium sulfate, which if sodium sulfate is removed by crystallization, it is removed as sodium sulfate decahydrate (Na2SO4 10H2O), also known as Glauber salt. This salt can be removed with a suction cone being cooled to force precipitation of the Glauber salt, which is easily crystallized and removed by a screen, filtration, centrifugation or other solid / liquid separation technique.
[0007] [007] US Patent No. 4,133,650 (Gamerdonk et al). discloses a regenerative process for recovering sulfur dioxide from exhaust gases using a regenerable aqueous dicarboxylic acid scrubbing solution (eg, phthalic acid, maleic acid, malonic acid and glutaric acid and mixtures thereof) buffered to a pH of about 2 , 8 to 9. The recovered sulfur dioxide can be used in the production of sulfuric acid.
[0008] [008] Similarly, US Patent No. 2,031,802 (Tyrer) suggests the use of salts of substantially non-volatile acids, having a dissociation constant between 1 x 10-2 and 1 x 10-5, measured at a dilution of 40 liters per gram molecule and a temperature 25 ° C (for example, lactic acid, glycolic acid, citric acid and orthophosphoric acid), in a regenerative process for the recovery of sulfur dioxide from effluent gases.
[0009] [009] US Patent No. 4,366,134 (Korosy) discloses a regenerative flue gas desulfurization process that uses an aqueous buffered potassium citrate solution to a pH of about 3 to about 9.
[0010] [0010] Organic solvents used in sulfur dioxide adsorption / desorption processes include dimethyl aniline, tetraethylene glycol, dimethyl ether and dibutyl butyl phosphonate. Like most solvents, the capacity of organic solvents is enhanced by high pressures and lower temperatures. The sulfur dioxide gas is then recovered (and the solvent regenerated) by decreasing the pressure and / or increasing the temperature. These organic solvents require the use of a metallic construction and often require solvent regeneration due to the formation of sulfuric acid and, in some cases, due to the reaction of the solvent with sulfur trioxide, formed by the side reaction of sulfur dioxide with oxygen during the absorption / desorption process. Organic solvents are generally more expensive than aqueous absorption solutions.
[0011] [0011] The significantly large flue gas flow rates emitted from a coal-fired power generation plant make the size of the equipment to recover sulfur dioxide very large. Organic solvents that require metal construction generally do not compete well economically with wet scrubbers, which commonly use fiber-reinforced plastic (FRP) in their construction, coated containers and low-cost alloys.
[0012] [0012] Conventional organic solvents are also hampered by one or more deficiencies with regard to desirable characteristics in an absorbent used in a sulfur dioxide absorption / desorption cycle. Many of these solvents have a relatively low sulfur dioxide absorption capacity, especially for the partial pressures of sulfur dioxide normally found in weak sulfur dioxide containing effluents (for example, from about 0.1 to about 5 kPa). These solvents often absorb significant amounts of water vapor from the sulfur dioxide-containing effluent, resulting in a significant reduction in the sulfur dioxide absorption capacity of the solvent. As a result, the molar flow rates of these solvents needed to satisfy the desired sulfur dioxide absorption efficiency are higher. In addition, the absorption of large amounts of water vapor in the solvent can lead to excessive corrosion of the process equipment used in the sulfur dioxide absorption / desorption process. In addition, some of these organic solvents are susceptible to excessive degradation, such as hydrolysis, or other side reaction or decomposition when the solvent is exposed to high temperatures in acidic environments and / or suffers from high volatility, leading to great solvent losses.
[0013] [0013] Co-opened and co-signed US Serial No. 13 / 283,671, filed October 28, 2011 and published as 2012/0107209 A1, describes a sulfur dioxide recovery process that uses a buffered aqueous absorption solution comprising weak organic or inorganic acids or salts thereof, preferably certain polyprotic carboxylic acids or salts thereof, to selectively absorb sulfur sulfate from the effluent gas. Subsequently, the absorbed sulfur dioxide is removed (stripped) to regenerate the absorption solution and produce a gas enriched in sulfur dioxide content. Sulfur dioxide enriched gas can be used as part of the feed gas for a contact sulfuric acid plant or a Claus plant for the preparation of elemental sulfur, or it can be used for the production of refined sulfur dioxide . The process described in US 2012/0107209 A1 is particularly useful in the production of a sulfur dioxide enriched gas from effluent gases relatively low in sulfur dioxide content. The application also describes processes for the simultaneous removal of sulfur dioxide and nitrogen oxides (NOx) from effluent gases and recovery of sulfur dioxide. The process uses a buffered aqueous absorption solution, including a metal chelate to absorb sulfur dioxide and NOx from the gas and, subsequently, reducing NOx to form nitrogen.
[0014] [0014] Although the US 2012/0107209 A1 process operates at a high energy efficiency, there remains a need for further savings in the use of energy in regenerative sulfur dioxide processes. SUMMARY OF THE INVENTION
[0015] [0015] The present invention is directed to new processes that comprise features that improve energy efficiency in regenerative adsorption / desorption cycles for the recovery of sulfur dioxide and other contaminants from gaseous effluents. In certain process modalities, energy is recovered from a gas stream of liquid contaminants in a desorption cycle. In this and other embodiments, the absorption zone can optionally and advantageously be cooled to increase the capacity of the absorption medium for the absorption of a contaminating gas, thus decreasing the volume of the aqueous absorption medium and the enriched absorption liquor. with contaminants that must be pumped, handled, heated and cooled in the absorption / desorption cycle.
[0016] [0016] A prominent application of the processes of the invention is in the recovery of sulfur dioxide from various chemical and metallurgical effluent gases, as mentioned above. However, the improvements described here are also applicable to the recovery of other acid gases, such as H2S, CO2, NOx, or HCl, and also to the recovery of other contaminating gases, such as ammonia.
[0017] [0017] In short, therefore, the present invention is directed to a process of selectively removing and recovering a contaminating gas from a source gas containing a contaminant, in which the feed gas stream comprises a source gas coming into contact with a contaminant absorber, with an aqueous absorption medium comprising an adsorbent for contaminant gas, thereby absorbing the contaminant gas from the feed gas stream in the absorption medium, and producing an exhaust gas from which a contaminant gas has been removed and a absorbent liquor enriched with contaminant. The contaminant-enriched absorption liquor comes in contact with stripping vapor to desorb the contaminant from the contaminant-enriched absorption liquor and thus produce a regenerated contaminant absorption medium and a primary stripper gas effluent comprising water vapor and contaminating gas. The regenerated absorption medium is removed from a liquid outlet of the absorption liquor stripper and effluent from the primary stripper gas is removed from a vapor outlet of the absorption liquor stripper; The water is condensed from the primary stripper gas effluent by indirect heat transfer from the primary stripper gas effluent to a cooling medium in a primary stripper gas cooler / condenser, thereby producing a condensate containing contaminant. Contaminant-containing condensate exiting the primary stripper gas cooler / condenser comes in contact with the steam in a condensed stripper to produce a stripped condensate and a condensed stripper gas effluent containing water vapor and contaminant gas. The cooling medium to which heat is transferred from the primary stripper gas effluent in the primary stripper gas cooler / condenser comprises at least part of the stripped condensate, thereby generating steam from the stripped condensate. The steam generated from the stripped condensate in the primary stripper gas cooler / condenser is introduced into the absorption liquor stripper as stripping vapor for contact with absorption liquor enriched with contaminants to desorb contaminant from it.
[0018] [0018] In one embodiment of the present invention, the primary stripper gas effluent removed from the absorption liquor stripper is compressed, and water is condensed from the compressed primary stripper gas effluent by indirect heat transfer from the stripper gas effluent. primary for the cooling medium, comprising at least a portion of the condensate stripped in the condenser / refrigerant of the primary stripper gas, thus generating vapor from the stripped condensate at a pressure in excess of the pressure within the absorption liquor stripper at the liquid outlet thereof. . The vapor generated from the stripped condensate in the primary stripper gas cooler / condenser is then introduced into the absorption liquor stripper as stripping vapor for contact with absorption liquor enriched with contaminants to desorb contaminant from it.
[0019] [0019] According to another embodiment of the present invention, the vapor generated from the stripped condensate in the primary stripper gas cooler / condenser is compressed at a pressure higher than the pressure inside the absorption liquor stripper at the outlet of the liquid the same. The compressed vapor is then introduced into the absorption liquor stripper as stripping vapor for contact with the absorption liquor enriched with contaminants to desorb the contaminant from it.
[0020] [0020] In this and these and other embodiments, the absorption zone can be cooled to improve the ability of an aqueous absorption medium to absorb a contaminating gas. In such embodiments, at least part of the contaminant-enriched gas-absorbing liquor is circulated between the absorber and a heat exchanger, where the absorption heat is removed by transferring it to a refrigerant.
[0021] [0021] Other objects and characteristics will be partly apparent and partly indicated hereinafter. BRIEF DESCRIPTION OF THE FIGURES
[0022] [0022] Figures 1 and 2 are alternative schematic flowcharts of the absorption / desorption processes to selectively remove and recover sulfur dioxide from a source gas containing sulfur dioxide, in which the absorption of sulfur dioxide from the absorption liquor is achieved by contact with live steam in a stripping column, and live steam is generated by the indirect transfer of heat from an upper gas stripper to the cooling medium, comprising a stream of boiling water in a gas stripper cooler / condenser .
[0023] [0023] Figures 3 and 4 are curves tracing the solubility of sulfur dioxide in certain absorption solvents as a function of temperature;
[0024] [0024] Figure 5 is a flow chart of an absorption / desorption process to selectively remove and recover dioxide and a source gas containing sulfur dioxide, in which an absorption liquor is circulated between the absorber and one or more heat exchangers external to cool the absorption liquor and improve the ability of the absorption medium to transfer sulfur dioxide from the gas phase;
[0025] [0025] Figure 6 represents the sulfur dioxide content in the gas phase and the percentage recovery of sulfur dioxide from a gas phase as a function of distance from the bottom of a countercurrent absorber of various combinations of gas, absorbent medium compositions , and liquid flow rate; and
[0026] [0026] Figure 7 illustrates absorption liquor temperature profiles and mol% of sulfur dioxide in the vapor phase of an absorption / desorption process for the recovery of sulfur dioxide in which a series of cooling loops are provided for the absorbent.
[0027] [0027] Corresponding reference numerals indicate corresponding component parts during drawings. DETAILED DESCRIPTION OF THE PREFERENTIAL MODALITIES
[0028] [0028] In accordance with this invention, several new process schemes have been developed for the recovery of a contaminating gas from a source gas at relatively high energy efficiency. The processes of the invention are particularly applicable for the recovery of acidic gases such as sulfur dioxide, nitrogen oxides, hydrogen sulfide, carbon dioxide and the like, but they are also useful and valuable in the recovery of other contaminating gases, such as, for example, ammonia. The generic term "contaminant" is used in this document because, typically, the processes of the invention are used to clean effluent gas streams from chemical or metallurgical power generation facilities in order to minimize emissions of acid gases or other gaseous components that otherwise would be contaminants in the atmosphere. However, as recognized by those skilled in the art, contaminant gases removed from effluent gas streams often have economic value and are recovered by processes of the invention, and then applied in commercial uses such as, for example, the conversion of carbon dioxide. sulfur to sulfur trioxide and sulfuric acid, recovery of elemental sulfur to hydrogen sulfide dioxide, recovery of hydrochloric acid or aqueous ammonia for use in chemical processing, recovery and conversion of hydrogen chloride to elemental chloride and hydrogen, etc.
[0029] [0029] The processes of the invention can be illustrated by the particular case of the recovery of sulfur dioxide. In the practice of the present invention, a variety of aqueous and organic solvents can be used as a means of absorbing sulfur dioxide. For example, the absorption medium can comprise aqueous solutions of alkali metals (for example, solution of sulfite / sodium bisulfite), amines (for example, alkanolamines, tetrahydroxyethylalkylenediamines, etc.), salts of amines or salts of various organic acids Alternatively , the sulfur dioxide absorption medium may comprise an organic solvent, including, for example, dimethyl aniline, tetraethylene glycol dimethyl ether or dibutyl butyl phosphonate. Some organic solvents require the use of a metallic construction and often require solvent regeneration due to the formation of sulfuric acid and, in some cases, due to the reaction of the solvent with sulfur trioxide, formed by the side reaction of sulfur dioxide with oxygen during the absorption / desorption process, and are generally more expensive than inorganic absorption media. The significantly large flue gas flow rates emitted from a coal-fired power generation plant make the size of the equipment to recover sulfur dioxide very large. Conventional organic solvents can also be hampered by one or more deficiencies with regard to desirable characteristics in an absorbent used in a sulfur dioxide absorption medium, as noted above.
[0030] [0030] Taking these and other considerations into account, and in accordance with the preferred embodiment of the present invention, the sulfur dioxide absorption medium comprises an aqueous solution buffered with a salt of a relatively weak polyprotic carboxylic acid (e.g., malate sodium), as described in the aforementioned Ser. US No. 13 / 283,671, filed on October 28, 2011 and published as on 2012/0107209 A1, all content of which is expressly incorporated into this document by reference. The following description refers to the preferred absorption medium comprising a polyprotic carboxylic acid salt, as well as to an absorption medium comprising tetraethylene glycol dimethyl ether (tetraglyme). However, it should be understood that the various characteristics of the processes described in this document are easily adaptable to systems in which other absorption media are employed. As noted above, it should also be understood that the improvements described in this document are equally applicable to systems for removing and recovering other acidic gases and contaminants, using suitable conventional contaminant absorption means known in the art. For example, the processes described in this document can be used for the regenerative absorption and desorption of various contaminants from effluent gas streams, including hydrogen sulfide, carbon dioxide and hydrogen chloride, nitrogen oxide, as well as other contaminating gases such as ammonia and mixtures thereof.
[0031] [0031] As shown in Figure 1, the process optionally conditioned process feed gas flow 10, which comprises a source gas containing sulfur dioxide being introduced into the sulfur dioxide absorber 11, having one or more theoretical stages where it comes into contact with an aqueous absorption medium comprising a sorbent for sulfur dioxide to absorb sulfur dioxide. The sulfur dioxide absorber 11 comprises a vertical column or tower 12 containing a gas / liquid contact zone 13 which comprises a means for promoting mass transfer between the gas and liquid phases which may comprise a bed of random packages, such as such as saddles or rings, structured packaging, or other contact devices. Preferably, in order to maximize the transfer of sulfur dioxide, the gas flow of the process feed process comes in contact with the current with the aqueous absorption solution. As shown in Figure 1, the process feed gas flow 10 is introduced through an inlet 14 near the bottom of the tower 12 and enters the bottom of the gas / liquid contact zone 13, while flow 15 comprises regenerating the medium aqueous absorption recirculated by the sulfur dioxide stripper 30 (described later in this document) be introduced through a liquid inlet 16 near the top of the tower, and it is distributed and enters the top of the gas / liquid contact zone. A stream of sulfur dioxide-enriched absorbent liquor 17 exiting the bottom of the gas / liquid contact zone 13 is drawn from a liquid outlet 18 near the bottom of the tower 12, and an exhaust gas stream 19 substantially free of carbon dioxide. sulfur leaving the top of zone 13 is taken from gas outlet 20 near the top of the tower. Although a conventional tower, randomly packaged, can be used as an absorbent 11, those skilled in the art will appreciate that other configurations can be employed appropriately. For example, the absorbent tower 12 may contain a structured package or comprise a tray tower, in any of which process flows preferably flow against the current. Although the countercurrent flow between the process feed gas flow 10 and the aqueous absorption medium in the absorbent is preferred, the absorbent can be operated simultaneously. However, this organization tends to negatively affect efficiency and absorption capacity and is generally less preferred.
[0032] [0032] Where an acid salt absorber or other species that combine chemically with sulfur dioxide are present as the main sorbent in an aqueous absorption medium, the concentration of the sorbent in the absorption medium and the absorption medium flow rate it must be such that, at the temperature prevailing at the liquid outlet of the absorber, excessive absorption capacity remains in the absorbent liquor. Preferably, the remaining capacity is at least 10%, preferably at least 20% of the total absorption capacity entering the absorbent. For this purpose, the adsorbent concentration and the flow rate of the absorption medium entering the absorbent must be sufficient to provide a stoichiometric excess in the flow rate of the sorbent by the absorbent relative to the rate at which the sulfur dioxide must be recovered by the flow process feed gas, preferably in excess relative to the total sulfur dioxide content of the feed stream, thus compensating for several factors such as: the sulfur dioxide content remaining in the absorption medium after its regeneration; the concentration of sulfur dioxide in the stripper gas enriched with sulfur dioxide; the possible presence of slightly acidic components such as carbon dioxide; but mainly to compensate for the desirable relatively low absorption affinity or preferred sorbents, such as an aqueous carboxylic acid / polyprotic salt absorption system. A relatively weak absorption affinity is preferred to facilitate the subsequent desorption of sulfur dioxide through a slight increase in temperature and / or pressure reduction. Accordingly, the concentration of sorbent in the aqueous absorption medium required to achieve a desirable removal efficiency varies with the acid employed, the amount of sulfur dioxide in the gas to be treated, as well as the mass transfer characteristics of the absorbent, and can be easily determined by someone skilled in the art. Typically, the ratio of stoichiometric equivalents of sulfur dioxide absorbed in mole of polyprotic carboxylic acid salt in the absorption solution ranges from about 0.1 to about 1. In the case of an aqueous absorption solution comprising the sodium acid salt malic to treat a gas comprising about 2600 ppmv (parts per million by volume) of sulfur dioxide, the concentration of malate in the absorption solution can suitably vary from about 1 mol% to about 7 mol%.
[0033] [0033] The mass flow rate (L / G) ratio of the absorbent aqueous solution flow 15 and process feed gas flow 10 introduced into the sulfur dioxide absorber 11 necessary to achieve a substantial transfer of sulfur dioxide from a source gas for an absorbent solution can be determined by conventional design practice. More particularly, the L / G can be selected based on the gas flow content entering the absorber, the concentration of the sorbent in the aqueous absorption medium, and the absorption capacity of the sorbent unit at the prevailing temperature of the liquid / gas in the absorber. . Normally, L / G is selected, such that the flow of sorbent into the absorbent is at least 10 to 20% excess under the flow of the contaminant gas in the absorbent. The ideal extent of the excess depends on the mass transfer rate and heat transfer in the gas / liquid contact zone.
[0034] [0034] Preferably, the sulfur dioxide absorber is designed and operated such that the sulfur dioxide content of the exhaust gas stream 19 exiting the absorber is less than 500 ppmv, and preferably less than 200 ppmv ( for example, as low as 10-20 ppmv). This small amount of sulfur dioxide, together with the carbon dioxide, oxygen, nitrogen and other inert substances contained in the feed gas flow process are eliminated from the system as part of the exhaust gas stream 19 exhaled from the top of the absorber. The exhaust gas is in substantial equilibrium with the absorption solution, and depending on the water vapor content of the gas flow process fed to the absorber, and the absorption conditions, there may be a net gain or loss of water in the absorber. If necessary, a fan 21 is used to conduct the gases to the stack. To achieve satisfactory emission standards, the exhaust gas flow 19 can be passed through a mist eliminator or similar device to recover the entrained liquid before being discharged through the stack. In addition or alternatively, in some cases the exhaust gas flow 19 can be heated by indirect heat exchange in a heat exchanger 22 with the inlet flow of the gas supply process or using other heating means or in the exchanger heat 64, as described below, so that any plume will not have the tendency to descend after being emitted through the stack.
[0035] [0035] As shown in Figure 1, where the sorbent comprises a polyprotic carboxylic acid, a 23 base metal replacement source, such as sodium hydroxide, potassium hydroxide, sodium carbonate, etc., it is combined with flow 15 comprising an aqueous absorption medium regenerated in a solvent tank 24, before being introduced near the top of an absorbent tower 12. The metal base reacts with the polyprotic carboxylic acid to form a metal salt absorber. In accordance with US 2012/0107209 A1, a sufficient amount of metal base is introduced to neutralize at least some of the acid groups, such that the acid is neutralized within 20%, more preferably within about 10% , from the acid dissociation equivalence point, having a pKa value of about 3 to about 10 to 25 ° C, preferably about 4 to about 7 to 25 ° C. Someone skilled in the art can use known pH control techniques and instruments to add base to the regenerated absorption solution, contacted with gas containing sulfur dioxide in the absorbent, to maintain a desired level of neutralization, with respect to the equivalence point of the pKa value. In addition, sufficient base must be added to maintain the metal ion concentration. For example, as described below, some of the metal ions are lost with the sulfate salt removed in a crystallization operation. Two moles of base (for example, sodium hydroxide) are added per mole of sodium sulfate removed. The concentration of the metal ion can be properly monitored and controlled by taking samples and performing metal analysis in the plant laboratory.
[0036] [0036] The absorption liquor enriched with sulfur dioxide 17 exiting the absorber 11 is heated to an intermediate temperature (as described below) and the preheated absorption liquor is introduced to the sulfur dioxide stripper 30 where the sulfate dioxide is disassociated of a sorbent and desorbed of an absorption liquor. Stripper 30 comprises a vertical column or tower 31 containing a liquid / vapor counting zone 32, comprising a means for promoting mass transfer between the gas phase and the liquid phase. Like the absorbent 11, the stripper 30 can be configured in the form of a packaged tower containing a conventional random packing bed, structured packing, trays or any other gas / liquid containment device. The lower section (stripping) of the liquid / vapor contact zone 32 within the tower 31 can be fed with live steam generated in accordance with the present invention (as described below) and used to remove sulfur dioxide from the absorption liquor. The upper (refining) section of the liquid / vapor contact zone 32 is used to reduce the amount of water in the desorbed sulfur dioxide. A stripper gas effluent enriched with sulfur dioxide 33, comprising sulfur dioxide substantially saturated with water vapor, is produced in the upper stripper 30 above the liquid / vapor contact zone 32, and removed from the steam outlet 34 at the top tower 31; and the regenerated absorption solution 15 leaving the liquid / vapor counting zone is withdrawn from a liquid outlet 35 at the bottom of the tower and recirculated back to the absorber 11, completing the cycle. Although countercurrent flow between the dioxide-absorbing liquor enriched with sulfur dioxide and the stripping vapor in the stripper, as shown in Figure 1, is preferred, the stripper can be operated simultaneously. However, this organization tends to negatively affect stripping efficiency and is generally less preferred.
[0037] [0037] The average temperature of the sulfur dioxide absorber in absorbent 11 is generally maintained in the range of about 10≡C to about 70≡C. In accordance with the present invention, the average temperature of the sulfur dioxide-absorbing liquor in the absorbent is preferably maintained at about 20 ° C to about 60 ° C. Although, in general, the absorption of sulfur dioxide is enhanced at lower temperatures of absorption medium, the absorption liquor needs to be heated from the absorption temperature to a temperature high enough and / or under reduced pressure to release the sulfur dioxide and providing this sensitive heat leads to higher energy demands. During regeneration, it is also desirable to reduce the amount of water vaporized to reduce energy consumption and to avoid low concentrations of water in the absorption medium that can cause precipitation of the sulfur dioxide sorbent (for example, weak polycarboxylic acid or salts). The overall efficiency of the sulfur dioxide absorption / desorption process is improved when absorption is relatively strongly dependent on temperature and within a narrower range of temperatures between the absorption and desorption steps of the cycle.
[0038] [0038] The average temperature of the sulfur dioxide-absorbing liquor in stripper 30 is generally maintained in the range of about 60 ° C, up to the boiling point of the absorption solution at stripper operating pressure.
[0039] [0039] The absorption and desorption of sulfur dioxide can be improved by increasing or decreasing the operational pressures of absorber 11 and stripper 30, respectively. The appropriate operating pressure in the absorber 11 ranges from about 70 to about 200 kPa absolute. Increasing the pressure in the absorbent increases the fraction of sulfur dioxide that the absorption medium can absorb, but the absorption is preferably carried out at relatively low pressures, thereby reducing equipment costs. Likewise, the proper operating pressures on stripper 30 are from about 40 to about 200 kPa absolute, but higher or lower operating pressures can be employed.
[0040] [0040] Temperature control inside absorber 11 and stripper 30 can be achieved by controlling the temperature of various process flows fed to these operations. Preferably, the temperature in stripper 30 is kept within the desired range by controlling the temperature of the absorption liquor enriched with sulfur dioxide 17 and steam introduced near the bottom of the stripper in a stripping section of the vapor / liquid counting zone. 32. Again referring to Figure 1, the absorption liquor enriched in sulfur dioxide 17 leaving the absorber 11 at a temperature of about 10 ° C to about 70 ° C, preferably from about 20 ° C to about 60 ° C is passed through a heat exchanger 40, where it is preheated to an intermediate temperature by indirect heat transfer from a regenerated absorption medium 15 being recirculated from the stripper 30 to the sulfur dioxide absorber. The transfer of heat from the regenerated absorption medium to the absorption liquor inside the exchanger increases the absorption capacity of the regenerated absorption medium and heats the absorption liquor, to help promote the resulting sulfur dioxide stripping. If additional heating is required to reach the desired temperature in the stripper, a liquor enriched in sulfur dioxide 17 can be passed through a solvent heater 41, where it is preheated (for example, by indirect heat transfer from a flow of sulfur dioxide product leaving the process), and / or further heated by indirect heat exchange with steam or with the hot condensed flow 70. In certain advantageous modalities, the absorption liquor enriched with sulfur dioxide is heated by heat transferred from a feed process gas flow and / or a regenerated sulfur dioxide absorption medium, without the addition of external. In such an embodiment, the temperature of the feed process gas stream is preferably not reduced below 50 ° C, and the temperature difference between the sulfur dioxide-enriched absorption liquor introduced into the stripper and the regenerated absorption medium is less about 40 ° C
[0041] [0041] The regenerated aqueous absorption medium 15 from leaving the bottom of stripper 30, at a temperature of about 60 ° C to about 140 ° C is cooled in exchanger 40 by transferring heat to the absorption liquor enriched with dioxide sulfur dioxide 17, leaving the sulfur dioxide absorber 11. Similarly, if an additional refrigerator is needed to maintain the desired temperature in the absorber, the regenerated absorption medium exiting the changer 40 can be passed through the solvent refrigerator 42 and cool further through indirect heat exchange with cooler tower water. Use of the heat exchanger 40 reduces the energy demand of the system so that the use of the solvent heater and / or solvent cooler may not be necessary.
[0042] [0042] In the preferred embodiments of the present invention, the levels of sulfate salt contaminants in the aqueous absorption solution comprising at least one salt of a carboxylic acid are maintained at an acceptable level by diverging at least a purge fraction 90 from a regenerated absorption 15 leaving stripper 30 for sulphate removal treatment. The relative volume of the purge fraction varies with the concentration of the sorbent in the regenerated absorption medium and the susceptibility of sulfur dioxide to oxidation during absorption and stripping. Typically, in an operation using malate as an absorbent, the purge fraction can represent less than 5% of the flow of regenerated absorption medium.
[0043] [0043] The purge fraction treatment comprises evaporating the water from the purge fraction 90 in an evaporator crystallizer 92 to produce a supersaturated concentrated solution in a sulfate salt. Crystals of sulfate salt are then precipitated from the concentrated aqueous absorption solution in the crystallizer to form a crystallization slurry 94, comprising crystals of sulfate salt and mother liquor. Sodium sulfate crystals are separated from the slurry in a conventional solid / liquid separation device 96, such as a vacuum or centrifugal filter and the mother liquor fraction 98 is recirculated to the solvent tank 24, where it is mixed with the flow main of the regenerated absorption medium to return to the absorber 11. The concentration of the aqueous absorption solution can be adequately achieved by heating and / or decreasing the pressure or increasing the flow of steam to the cooler, to evaporate the water by flash. Typically, the aqueous absorption solution is heated to a temperature of at least about 40 ° C, and preferably at least about 60 ° C, and preferably to the boiling point of the absorption solution at the operating pressure of the stripper, during concentration to inhibit the formation and precipitation of sodium sulfate decahydrate or Glauber salt (Na2SO4-10H2O). Glauber salt tends to form a gelatinous or sticky precipitate that is not easily separated from the mother liquor by centrifugation or filtration.
[0044] [0044] The crystallizer can be operated at atmospheric pressure or under vacuum. As an alternative to separating crystals of sodium sulfate salt by centrifugation or filtration, the crystallizer can be designed to continuously decant the mother liquor from the crystallization slurry. In addition, the sulfate salt crystals can be washed with water, and the resulting wash water comprising the polyprotic carboxylic acid salt absorber directed, likewise, to the solvent tank, to return to the absorber. The upper vapor flow from the crystallizer can be condensed and returned to the absorber. Alternatively, the airflow from the crystallizer can be routed to the stripper as a source of stripping steam.
[0045] [0045] Although the treatment described above is effective in maintaining acceptable sulfate salt levels in the circulating absorption solution, in accordance with some embodiments of the present invention, an oxidation inhibitor can be included in the absorption solution to reduce oxidation bisulfite and sulfite and bisulfate and sulfate contaminants, respectively. There are several types of oxidation inhibitors that can be useful in the practice of the present invention, including: oxygen absorbers and free radical scavengers, such as p-phenylenediamine and hydroquinone; NOx-catalyzed oxidation inhibitors, such as ascorbic acid; and chelate agents, such as ethylene diaminetetraacetic acid (EDTA), which sequester and inhibit metal-catalyzed oxidation. Such oxidation inhibitors can be used individually or in various combinations, and can be added as needed for the regenerated aqueous absorption solution introduced into the absorbent. Depending on the type of inhibitor (s) employed, the concentration in the absorption solution typically ranges from a few ppm to about 1 to about 10 percent by weight. An excess is usually added (for example, at least about 1000 ppm), since the inhibitors will gradually be consumed by oxidation. Ascorbic acid and hydroquinone are particularly effective in inhibiting oxidation in a sodium malate absorption solution. EDTA is expected to be an effective oxidation inhibitor when metals are present in the absorption solution.
[0046] [0046] The increase in acidity of the absorption solution has the effect of increasing the stripping efficiency of sulfur dioxide. Thus, leaving a small concentration of dissolved sulfur dioxide or keeping some sulfate in the absorption solution leads to greater efficiency in the stripper. For example, a small concentration of sodium sulfate and / or sulfurous acid in the stripper makes regenerating the absorbent solution less energy consuming. In accordance with an embodiment of the invention, the sulfate salt concentration is maintained at about 0.5 to about 11 weight percent, preferably about 3 to about 11 weight percent in the absorption solution and a small fraction of sulfur dioxide is left in the regenerated aqueous absorption solution, thus making the solution slightly more acidic and, consequently, making the sulfur dioxide desorption consume less energy. Generation of Stripping Vapor and Condensing Stripping
[0047] [0047] To provide a source of energy to generate stripping vapor, effluent from stripper main gas 33 from absorption liquor stripper 30 is compressed in a suitable device to increase the temperature of the primary stripper gas effluent. Suitable devices include mechanical compressors and thermal compressors (for example, steam jet ejectors). As shown in Figure 1, the main stripper gas effluent is preferably compressed by passing it through a steam jet ejector 36. Where sulfur dioxide is recovered from the residual gas of a contact sulfuric acid plant, a flow generated in the Heat recovery from sulfur trioxide absorption can provide motivating steam for the ejector.
[0048] [0048] Even though absorption / desorption systems for sulfur dioxide recovery in which the wet sulfur dioxide stripper gas is compressed and the latent heat of water vapor condensation is transferred from the compressed gas to the liquor of absorption enriched with sulfur sulfate, in such systems the condensate leaves the sulfur sulfate saturated system. Unless the sulfur dioxide emanating from the condensate is captured in a separate system, this scheme creates unacceptable emissions that also equate to the loss of sulfur dioxide values.
[0049] [0049] The process described in said US 2012/0107209 A1, sulfur dioxide is recovered from the condensate in a condensate stripping column, but this implies additional energy consumption.
[0050] [0050] According to the process of the present invention, the energy required for condensate stripping is substantially recovered by using the stripped condensate as a source of stripping vapor for the absorption liquor stripper. Additional energy input is required to vaporize the condensate at a pressure sufficient to flow to the stripper base. In the process of the invention, the latent heat in the water vapor component of the stripper gas provides this energy source. The modest compression of the stripper gas exiting the absorption liquor stripper creates the modest differential temperature sufficient to transfer heat from the stripped compressed gas to the stripped condensate, thereby vaporizing the stripped condensate at a pressure sufficient to conduct the resulting vapor in the stripper .
[0051] [0051] The compression of the gas effluent containing wet sulfur dioxide from the stripper preferably increases the flow pressure by an increase of about 30 kPa to about 65 kPa. Separation of sulfur dioxide is enhanced if stripper 30 is operated at lower pressures (for example, under vacuum) to increase the relative volatility of sulfur dioxide in relation to water and increase desorption and decrease the number of theoretical stages required for a reflux. In addition, lower pressures lead to lower system temperatures, allowing the use of low pressure steam to heat the absorption liquor enriched with sulfur dioxide. However, energy recovery is optimized at moderately higher operating pressures, and this also reduces the required tower 31 diameter and associated capital cost. As an example, the operation of the stripper under a slight vacuum (for example, -35 kPa gauge) and the modest increase in the pressure of the sulfur dioxide-enriched stripper gas exiting the stripper (for example, gauge of about 20 kPa) represents an economical approach. However, operating the stripper at or above atmospheric pressure can also be an attractive approach. Economic optimization can determine specific operating conditions. Balancing these considerations, the pressure of the primary stripper gas effluent exiting the absorption liquor stripper is preferably maintained from about 40 to about 170 kPa absolute).
[0052] [0052] The pressurized flow of stripper gas containing sulfur dioxide is directed to a primary stripper gas cooler / condenser 50. A substantial part of the water vapor is condensed from the primary gas stripper effluent into the refrigerator / condenser 50 by indirect heat transfer to a cooling medium. In accordance with the present invention, condensate stripped in flow 51 flowing to the cooler / condenser 50 of a condensate stripper or water column 60 (whose operation is described below) serves as a cooling medium, and the latent heat of condensation is transferred for the stripped condensate, generating steam which is used as a stripping medium in an absorption liquor stripper 30. As shown in Figure 1, the stripped condensed flow 51 from column 60 is directed to a liquid-vapor separator 52 (for example, example, steam drum) and circulates via line 54 between the separator and the cooler / condenser 50 where the heat transfer from the primary stripper gas generates steam 53 for the stripper. Steam and stripped condensate are separated in separator 52, steam is directed to stripper 30 via line 57, at least part of the condensate circulates to the primary stripper gas cooler / condenser 50 via line 54 and another part, optionally, can be recirculated and combined with the regenerated sulfur dioxide absorption solution 15 via line 55 and returned to absorber 11 and / or a part 56 can be purged from the system. Alternatively, the condensate side of the primary stripper gas cooler / condenser 50 can be designed to allow the removal of water vapor within the heat exchanger itself, allowing a free flow of entrained water to flow directly from the refrigerator / condenser. for the absorbent, without the need for a separate liquid / vapor separator.
[0053] [0053] Steam generated in the primary stripper gas cooler / condenser 50 is introduced into stripper 30 via line 57, where it comes into contact with the absorption liquor in the liquid / vapor contact zone 32, both supplying heat to the absorption liquor while functioning as a stripping gas for the removal of sulfur dioxide from the liquid phase. Heating of the liquid phase in the absorption liquid stripper reduces the concentration in the sulfur dioxide equilibrium contained therein and increases the driving force for the transfer of sulfur dioxide to the vapor phase. In the heat transfer to the liquid phase, steam generated from stripped condensate in the refrigerator / with condenser 50 partially condenses inside the stripper, thus functioning essentially as a condensable stripping gas. Optionally, the stripping heat provided by the steam generated from the primary stripper gas cooler / condenser can be supplemented by the heat supplied from a foreign source in a refector 37, through which the liquid phase of the absorption liquor stripper is circulated. . The auxiliary cooler provides total flexibility in controlling the water balance of the process. Normally, absorption liquor to be passed through the cooler is removed from a stripper reservoir and returned to the bottom of the liquid / vapor contact zone 32 above the reservoir.
[0054] [0054] In the primary stripper gas cooler / condenser 50, most of the water vapor content of the primary stripper gas effluent 33 is condensed and, therefore, most of the latent heat removed by transfer to stripped condensate returning from the condensed stripper 60. The aqueous condensate obtained by condensing water vapor from the primary stripper gas effluent comprises dissolved sulfur dioxide. This condensate is removed from the cooler / condenser 50 and fed through line 58 to the condensate stripper or water column 60 and heated (for example, with steam or a cooler) to desorb sulfur dioxide and produce a condensed stripper gas comprising water vapor and sulfur dioxide desorbed from the aqueous condensate. As shown in Figure 1, condensate stripper gas is combined with vent gas containing wet sulfur dioxide 59 from the primary stripper gas cooler / condenser 50. The combined final condenser stripper gas 61 exiting the top of the condensed stripper column 60 is cooled to a temperature normally below about 70 ° C in a low temperature condenser 62 (for example, with cooling water at 50 ° C) to condense the water vapor and produce a product stream 63 comprising sulfur dioxide recovered. As shown in Figure 1, additional marginal condensate can be squeezed out of the condensed stripper gas, or the combined final condensed stripper gas 61 leaving the top of the condensed stripper column 60, passing the gas first through a heat exchanger 64 where the condensed stripper gas is cooled by transferring heat to a portion of the vent gas 19 exiting the absorber 11. After cooling, the stream of recovered sulfur dioxide product 63 is removed from the sulfur dioxide recovery process and directed to a destination where it can be used, for example, for the drying tower or a catalytic phase of a sulfuric acid contact plant for conversion to sulfur trioxide, for a Claus process operation for the generation of elemental sulfur, the an alkali metal sulphite or bisulfite fabrication process, for a papermaking operation, or for a compression and refrigeration unit for liquefaction of liquid sulfur dioxide.
[0055] [0055] The stripped condensate stream 51 depleted in sulfur dioxide exits the bottom of the condensed stripper column 60 and is directed to the primary stripper gas cooler / condenser 50 in which the condensation of water vapor from the gas effluent of compressed primary stripper 33 transfers heat to the stripper condensate, thereby generating steam for use as a combined heating medium and stripping gas (for example, as a condensing stripping medium) in an absorption liquor stripper 30. Optionally , a part 56 can be purged from the system.
[0056] [0056] The extent of compression of the primary gas effluent stripper 33 of the absorption liquor stripper 30 is necessarily sufficient to bring the compressed steam to a temperature high enough to have a higher pressure than the pressure at the bottom (stripping) ) of the liquid / steam contact zone section 32 within tower 31 can be generated by the stripped condensed heating of stripping in the main cooler / condenser stripper gas 50. But the extent of compression is preferably controlled to a minimum necessary for the steam generated at from the stripping condensate flow into the stripper. More particularly, it is preferable that the steam is generated from the stripping condensate at a temperature no more than 30 ° C higher than the temperature of the liquid phase within the absorption liquor stripper at the liquid outlet 35 of these, or more particularly, not more than about 20 ° C or no more than about 5 to about 10 ° C higher than the temperature of the liquid phase exiting the bottom of the vapor / liquid contact zone 32 within the stripper. In certain particularly preferred embodiments, the temperature of the steam produced by heating the stripping condensate in the primary cooler / condenser gas stripper 50 is no more than, or may be even less than, the temperature of the liquid phase inside the stripper of liquor absorption at the liquid outlet, or at the bottom of the vapor / liquid contact zone. In general, it is preferable that the temperature of the steam generated in the primary cooler / condenser gas stripper 50 varies from the temperature of the regenerated absorption medium inside the stripper at the outlet of the respective liquid or the temperature of the liquid phase, leaving the bottom ( stripping) of the liquid / vapor contact zone within the absorption liquor stripper, by not more than about ± 10 ° C. In order for the vapor to flow through the absorption liquor stripper, the pressure of the vapor generated in the evaporator / condenser 50 is necessarily greater than the total pressure in the stripper and, therefore, greater than the equilibrium vapor pressure of the liquid phase within the liquid / vapor stripping section of the contact zone, even with the output of the liquid phase of the stripping section, where the partial pressure of sulfur dioxide approaches zero as a limit.
[0057] [0057] The consequent driving force in the water pressure in the vapor phase, therefore, causes condensation of water vapor to occur in the stripper regardless of temperature differences between the vapor phase and the liquid phase, resulting in condensation and heating of the liquid phase within the vapor / liquid stripping section of the contact zone, even if the vapor is introduced into the zone is a temperature not exceeding, or even slightly below, the temperature of the liquid phase. Because of the depressive effect of the solute, that is, an adsorbent such as a salt of polyprotic carboxylic acid, in the liquid phase, the vapor pressure of the liquid phase may be slightly lower than the vapor pressure at the same temperature, or even where, the The temperature of the liquid phase is slightly higher than the temperature of the vapor.
[0058] [0058] To meet these preferred conditions, the average log temperature differential (Δt) in the main cooler / condenser stripper gas is no less than about 1.5 ° C, about 2 ° C, about 3 ° C , about 4 ° C, or about 5 ° C and not more than 10 ° C, about 8 ° C, about 6 ° C or about 5 ° C. For example, the average log temperature differential (Δt) in the main cooler / condenser gas stripper is about 1.5 ° C to about 10 ° C, or about 2 ° to about 9 ° C, or from about 2.5 ° C to about 8 ° C.
[0059] [0059] Depending on the overall energy balance of the process and the water, the volume of the condensate from the stripping of the stripper condensate 60 may exceed the demand for steam in the absorption liquor stripper 30. Thus, the stripping condensate can be usefully divided between flow of (i) a condensate directed to the primary cooler / cooler gas stripper 50 as a cooling fluid for water condensation of the stripper gas, thereby converting the stripping condensate at least in part to the steam for introduction to the absorption liquor stripper; and (ii) a flow of discharge water to remove water from the process.
[0060] [0060] A portion of the stripped condensate from the condensed stripper 60 as discharge water can also be used, optionally, to condition the source gas containing sulfur dioxide or process feed gas 10. As shown in Figure 1, the stripped condensate the liquid-vapor separator 52 is passed through line 70 and introduced into a saturator 71 upstream of the sulfur dioxide absorber 11 in relation to the feed gas flow. Saturation may comprise a phase contact (for example, usually consisting of a packed column or tower containing random or structured packaging or a spray column), in which the stripped condensate comes into contact with the gas flow, thereby increasing humidity of the feed gas entering the sulfur dioxide absorber. The flow of water out of the saturator can be removed from the process. The saturator also cools the gas containing sulfur dioxide by evaporative cooling and removes acid gases (for example, sulfuric acid, hydrochloric acid, sulfur trioxide) before entering the absorber. The saturator advantageously allows the humidification of the feed gas flow using water of inferior quality, which provides an incremental cost saving compared to the humidification of the absorbent gas, where the water used must be ionized or distilled to avoid the accumulation of impurities. Although the flow of water from the saturator is saturated with sulfur dioxide, the volume of the flow is small. In addition, where, for example, sulfur dioxide is recovered from the residual gas of a sulfuric acid plant, the flow of sulfur dioxide-laden water out of the saturator can be used as dilution water in an SO3 absorber. Advantageously, in an interpass plant, water is used for dilution in the interpass absorber, the minimum liquid flow of sulfur dioxide involved back through the sulfur dioxide recovery unit is not lost from the process.
[0061] [0061] The process of Figure 1 compresses the primary stripper gas effluent in order to provide the temperature differential, whereby the latent heat recovered by the condensation of water vapor from the primary stripper gas is transferred to the stripped condensate for steam generation. is introduced for absorption liquor stripping in the absorption liquor stripper. In accordance with the invention, other alternatives are provided to generate this temperature differential and direct the stripping operation.
[0062] [0062] Figure 2 illustrates an alternative to the process of Figure 1 in which the steam generated from the stripped condensate is compressed by a compressor 39 during the flow between the steam outlet from the cooler / condenser 50 and the absorption liquor stripper 30 The drawing shows the compression of steam by a mechanical compressor, but steam could also be introduced into the neck of a steam jet ejector to achieve the required compaction. The diameter of stripper 30 is dimensioned, and the packaging or other mass transfer promotion structure within the liquid / vapor contact zone 32 of stripper 30 is designed to prevent excessive pressure drop during the passage of the gas phase / steam upwards through the zone. The primary stripper gas outlet 34 and the line used to transfer the primary stripper gas effluent 33 to the chiller / condenser 50 are also sized to prevent excessive pressure drop. Preserving a pressure on the primary gas stripper side of the cooler / condenser 50 greater than the pressure on the stripped condensate side of that changer, a temperature differential is established by which heat is transferred to the stripped condensate as condensed water vapor from the effluent from primary stripper gas and steam is generated on the condensate side for use in stripper 30. The steam generated in the cooler / condenser 50 is introduced into the suction side of the compressor 39 which compresses the steam for introduction into the line via stripper 57.
[0063] [0063] To recover the latent heat of condensation from the water vapor from the stripping gas, compressor 39 increases the vapor pressure to a level such that, when the primary stripper gas reaches cooler / condenser 50, the pressure on the side of The refrigerator / condenser stripper gas is greater than the pressure of the vapor generated from the stripped condensate on the side of the refrigerator / condenser stripped condensate. More particularly, the extent of compression is sufficient so that the water saturation pressure at which the water vapor condenses on the primary stripper gas side of the cooler / condenser is greater than the pressure at which the steam is generated on the side stripped condenser from the refrigerator / condenser.
[0064] [0064] The temperature and differential pressure achieved in the process of Figure 2 is, preferably, essentially the same as that which prevails in the refrigerator / condenser 50 in the embodiment of Figure 1 in which the effluent of the primary stripper gas is compressed during the flow from from the stripper gas outlet to the cooler / condenser gas inlet. The predominant absolute pressure in the liquid / vapor contact zone is, preferably, also in the same range for each of the modalities respectively shown in Figures 1 and 2. In both cases, it is desirable to maintain a pressure slightly above atmospheric, for example. example, from about 15 to about 18 psia (from about 100 to about 125 kPa absolute), in the stripper. However, because only the vapor is compressed in the process of Figure 2, the ideal pressure within the stripping zone of the absorption liquor in the process of Figure 2 may be slightly less than the ideal pressure in the process of Figure 1 in which the dioxide component sulfur content of the primary stripper gases must also be compressed by bringing the partial pressure of the water vapor to a level at which the water vapor will condense at a temperature above the boiling temperature of the water on the stripped condensate side of the refrigerator / condenser 50 .
[0065] [0065] The remainder of the process in Figure 2 is operated in a substantially identical manner to that described above with respect to Figure 1.
[0066] [0066] Although the processes in Figures 1 and 2 provide comparable energy efficiency, an advantage of the process in Figure 2 is the substantial absence of sulfur dioxide from the flow subject to compression. This means that the fluid to be compressed is generally less corrosive than the compressed fluid in the process of Figure 1 and therefore provides savings in maintenance and selection of construction materials for the compressor or ejector.
[0067] [0067] Dependence on saturated steam, generated from stripped condensate in the primary stripper gas cooler / condenser as the only energy source for the sulfur dioxide stripping from the absorbing liquor can result in a liquid accretion of water in the medium of regenerated absorption circulated back to the absorbent, ultimately in the medium sorbent circuit between the absorbent and the stripper. In fact, any stripper operation that relies solely on live steam necessarily has this effect due to the increase in steam that needs to be added to provide the heat of vaporization of sulfur dioxide and the resulting increase in heat loss from the environment. Thus, the control of the water balance in this circuit requires some measure for the removal of the water fraction that can be otherwise acquired in this operation scheme. Several options are available for this purpose. For example, the energy supplied from an external source in refueler 37 can marginally increase the temperature of the primary stripper gas so that it carries a slightly larger load of steam, and the primary stripper gas cooler / condenser can be operated with an Δt and a slightly higher vent gas temperature to remove sufficient water vapor build-up, thus maintaining water balance. This may require a slightly higher compression of the primary stripper gas in the embodiment of Figure 1, or a slightly greater compression of a stripping vapor in the embodiment of Figure 2. Alternatively, some or all of the absorption liquors may regenerated absorption liquor divert the need for exchanger 40 and / or cutting cooler 42, thus allowing the absorber to operate at a slightly higher temperature which incrementally increases the water vapor content of the exhaust gas to maintain equilibrium.
[0068] [0068] In a typical operation of the process of Figure 1, about 2% of the water volume occurs during each rotation of the absorber / stripper circuit. In a modality in which the flue gas containing sulfur dioxide, at levels that reflect the sulfur content of coal or other sulfur-containing carbon fuel, is delivered to the absorber at 27 ° C, a balance can be achieved by avoiding the use of the regenerated absorption medium around the interchange 40 and cutting cooler 42 and feeding the absorption medium in the absorber at 40 ° C. The exhaust gas exiting the absorber at 35 ° C carries enough water vapor to balance the gain from the increase in steam needed to vaporize sulfur dioxide from the absorption liquor in the absorption liquor stripper. Sulfur Dioxide Recovery from Rich Gas Flows
[0069] [0069] The process of the invention is suitable for the recovery of sulfur dioxide from the residual gas of a contact sulfuric acid plant and other operations that generate relatively weak sulfur dioxide effluents. However, it is applicable to other process operations that require sulfur dioxide recovery, including operations that generate relatively rich sulfur dioxide gas streams. Since the sulfur dioxide absorption reactions of a feed gas are typically exothermic, significant heat reactions are generated in the absorber where the process is used to recover sulfur dioxide from rich gases containing, for example, from about from 2 to about 4 vol%. sulfur dioxide or more, including gas streams where the sulfur dioxide content can be as high as 10% by vol., 15% by vol., 20% by vol., 25% and, vol., 30 % by vol., 40% by vol., or more. For example, the concentration of sulfur dioxide can be at least about 4% by vol., Or at least about 5% by vol., Or at least about 10% by vol., Or at least about 15%. in vol., or at least about 20% by vol. or at least about 30% by vol.
[0070] [0070] The process of the invention is very easily adaptable to recover sulfur dioxide from such streams of rich gases containing sulfur dioxide. However, where the sulfur dioxide content of the gas flow is high, the sensitive heat generated in the reaction of the exothermic absorber can sharply increase the temperature of the absorption liquor, in some instances to levels that can seriously compromise the absorption efficiency and / or the absorbent capacity of the circulating absorption medium. For example, in an absorption system using tetraglyme as the sorbent, where the sulfur dioxide concentration of the incoming feed gas reaches 2.9% by vol., The temperature of the absorption liquor can rise from a temperature of 17 ° C, typically preferred, for a temperature of 30 ° C in otherwise acceptable L / G ratios. Where the sulfur dioxide content of the inlet gas is 43 mol%, the temperature can typically rise from 17 ° to 49 ° C. For a tetraglyme absorption system, this increase in temperature can seriously compromise the capacity of the sulfur dioxide absorption medium.
[0071] [0071] Figures 3 and 4 illustrate the adverse effect of temperature on the equilibrium absorption capacity of two known sulfur dioxide absorption solvents. As illustrated in Figure 3, using 100% by weight of tetraglyme (100S) as a sorbent in 4% mol SO2 in the gas, the sorption capacity of the absorbent medium decreases significantly as the temperature rises, even in a narrow range of 20 ° to 30 ° C. Absorption capacity continues to decrease even at higher temperatures, although this decline is less drastic. As illustrated in Figure 4, where the feed gas contains 30% mol SO2, the absorption capacity of pure tetraglyme (100S) decreases more evenly as the temperature increases. As also shown in Figures 3 and 4, comparable declines in absorption capacity are incurred using another tetraglyme sorbent, for example, 955_5 W (95% by weight of tetraglyme and 5% by weight of water) Thus, for rich gases containing more than 2% by vol. sulfur dioxide, an increase in absorption medium flows are generally needed to reduce the amount of temperature rise in the liquid phase passing through the absorbent, which results in a lower sulfur dioxide concentration in the absorption liquor enriched with sulfur dioxide sulfur.
[0072] [0072] The increased flow of the absorption medium and the absorption liquor rates the absorption liquor stripper in two important ways. It increases the energy demand for heating the absorption liquor to the appropriate temperature for the stripping of the sulfur dioxide from it, thus reducing the energy efficiency of the process. But it also imposes a greater mass flow across the stripping column, thus increasing the column diameter needed to accommodate the liquid flow without flooding the vapor / liquid contact zone. Higher rates of liquid phase flow also dictate a larger diameter of the absorption column.
[0073] [0073] In accordance with an even more preferred feature of the sulfur dioxide absorption process, refrigeration is provided at the base of the absorbent in order to reduce the increase in temperature in the absorption medium as it passes through the absorption zone (ie , gas / liquid contact) and thus allow the absorber and stripper to be operated at relatively low L / G ratios. Controlling the rise in temperature in the absorption medium, especially at the bottom of the absorption zone, preserves the equilibrium capacity of the absorption medium and thus preserves the driving force for the mass transfer of sulfur dioxide from the gas phase to the liquid phase within the absorption zone, as well as the driving force for the reaction of sulfur dioxide with the sorbent in the liquid phase. The relatively low liquid phase temperatures also favor the degree of conversion of the sulfur dioxide adduct within the liquid phase where the reaction between sulfur dioxide and the sorbent is an exothermic equilibrium reaction. Preferably, the absorption liquor is removed from the gas / liquid contact zone inside the absorber, circulated through an external heat exchanger and returned to the absorption zone. More particularly, the circulating absorption liquor is removed from the gas / liquid contact zone in a region spaced below the region to which the refrigerated circulating absorption liquor is returned to the zone, thereby defining a section within the absorption zone below the region. to which the refrigerated absorption liquor is returned, within which most of the absorption of sulfur dioxide occurs preferentially and most of the absorption heat is generated.
[0074] [0074] For example, as shown in Figure 5, a portion of the absorption liquor enriched with hot sulfur dioxide 17 is removed from the liquid outlet 18 or removed from a region 13.1 near the bottom of the vertical gas / liquid contact zone 13 in the absorber 11 and circulated through an external heat exchanger 80, where the absorption heat is removed by transferring it to a refrigerant. The refrigerated absorption liquor is returned to the absorber in a region 13.2 of the gas / liquid contact zone that is spaced above the region from which the hot absorption liquor is withdrawn, but spaced below the upper part of the gas / liquid contact zone. More preferably, the region 13.2 to which the refrigerated circulating absorption liquor is returned is at the bottom of the gas / liquid contact zone.
[0075] [0075] The circulation of the absorption liquor between the sulfur dioxide absorber 11 and the external heat exchanger 80 causes an increase in mass flow and the inevitable return of the absorption liquor mixture in the circulation section of the absorption zone that is between regions 13.1 and 13.2, and this may slightly offset the gain in mass transfer for the removal of sulfur dioxide in this section of the zone. Preferably, therefore, the return region 13.2 is spaced by the height of at least one transfer unit below the top of the gas / liquid contact zone, thus defining a rectification section of the absorption zone, comprising at least one unit of transfer below the top of the zone. Preferably, the grinding section comprises at least two transfer units. It is also preferred that the return region 13.2 is spaced by the height of at least one transfer unit, more preferably at least two transfer units above the withdrawal region 13.1. To accommodate the proper mass transfer capacity in the circulation section of the absorption zone between the return region 13.2 and the withdrawal region 13.1 and the rectification section between the return region 13.2 and the upper part of the absorption zone, the The absorption as a whole preferably comprises at least three, more preferably at least four transfer units. Because both gas and liquid flows are in substantial plug flow within the grinding section, maximum driving force for mass transfer is provided in that section, allowing the sulfur dioxide concentration in the exhaust gas to be reduced to a level that meets the emission standards. The appropriate site selection for the return region of the circulating liquid 13.2 is based on the selection of a region, in which the level of sulfur dioxide in the gas flowing upwards from there is not high enough to generate absorption / reaction heat. in the grinding section that would have a significant adverse effect on the absorptive capacity of the aqueous absorption medium, or on the driving force of mass transfer in the grinding section.
[0076] [0076] Preferably, where the sorbent is tetraglyme, region 13.2 to which the refrigerated circulating absorption liquor is returned to the gas / liquid contact zone is maintained at a temperature no higher than about 40 ° C, more preferably not higher than about 30 ° C, more usually from about 15 ° to about 25 ° C. In a tetraglyme system, the temperature of the 13.1 region from which the hot circulating absorption liquor is removed from the gas / liquid contact zone is preferably maintained at a temperature not greater than about 45 ° C, more preferably not greater than 35 ° C, more usually from about 15 ° to about 30 ° C. Those skilled in the art will recognize that different temperature ranges, in some substantially different cases, are optimal for other sorbents. For example, where the adsorbent sodium malate, region 13.2 to which cooling liquor absorption circulation is returned to the gas / liquid contact zone is maintained at a temperature not exceeding about 45 ° C, more preferably not exceeding about 45 ° C, more usually from about 20 ° to about 40 ° C. In this case, the temperature of the 13.1 region of the circulating hot absorption liquor is removed from the gas / liquid contact zone is preferably maintained at a temperature not exceeding approximately 50 ° C, more preferably not exceeding 40 ° C, more usually from about 25 ° to about 35 ° C. In each case, the rate of circulation between regions 13.1 and 13.2 is dictated by these temperature and power generation restrictions the unit of the absorption process.
[0077] [0077] Conveniently, a fraction of the front flow of absorption liquor enriched with hot sulfur dioxide 17 is removed from the flow of absorption liquor circulating upstream of the external heat exchanger 80 and directed to the absorption liquor stripper 30.
[0078] [0078] The location of the return region of circulating absorption liquor 13.2 can be selected based on the absorption profile for the sulfur dioxide absorption zone. Typical profiles using different absorption media are shown in Figure 6.
[0079] [0079] Where absorption is immediate and substantially quantitative until the contact of the feed gas with the absorption medium in the gas / liquid contact zone, a single cooling circuit of the absorption liquor is normally sufficient to preserve the absorption efficiency and control the volumetric flow of the absorption liquor at a level consistent with energy efficient use in the absorption liquor stripper. However, where the affinity of the sorbent for sulfur dioxide is more limited, as it is also desirable for the purpose of efficient functioning of the absorption liquor stripper, the sulfur dioxide concentration gradient across the absorption zone, that is, the The rate at which the concentration of sulfur dioxide in the gas stream (and the liquid stream) decreases with the distance above the gas inlet to the absorption zone, can only be modest. In such circumstances, greater efficiency in the operation of the absorber and stripper can be achieved by means of two or more vertically spaced cooling loops along the gas flow path within the absorption zone (ie gas / liquid contact) . For example, as shown in Figure 5, two of these cooling loops are shown. In the second refrigeration loop, a second part of absorption liquor enriched with hot sulfur dioxide descending to the gas / liquid contact zone 13 of the absorber 11 is removed from a region 13.3 above the region 13.2 to which the refrigerated circulating absorption liquor is returned to the gas / liquid contact zone in the first cooling loop and circulated through an external heat exchanger 81, where the absorption heat is removed by transferring it to a cooling fluid. The refrigerated absorption liquor is returned to the absorber in a region 13.4 of the gas / liquid contact zone that is spaced above the region 13.3 from which the hot absorption liquor is removed, but spaced below the upper part of the gas / liquid contact zone .
[0080] [0080] Figure 7 illustrates the operation of an absorbent / stripper system in which sulfur dioxide has only a modest affinity for the adsorbent, so that the sulfur dioxide gradient is relatively shallow. Figure 7 graphically represents the temperature of the absorption liquor and the concentration of sulfur dioxide in the gas flow within the absorption zone, in each case depending on the location in the absorption zone expressed as the distance in transfer units from the top , that is, the gas outlet in the zone, with different curves for the systems containing no cooling loop, a cooling loop, two cooling loops and three cooling loops, respectively. Data on the effect of one, two or three cooling loops are also set out below in Table 1. Table 1: Impact of Cooling Loops on Flow Requirements
[0081] [0081] The data plotted in Figure 7 and tabulated in Table 1 are from a sulfur dioxide absorption system in which the absorbent comprises 15 stages (essentially corresponding to the transfer units). In each case where the circulating absorption liquor is cooled, there is at least one loop in which the withdrawal region is stage 15 and the return region is stage 13, that is, the return region is spaced by the height of essentially two transfer units from the bottom of the absorption zone and spaced by the height of 12 units from the top of the zone. Where a second loop is added, the withdrawal region is stage 10 and the return region is stage 8, and where a third loop is used, the withdrawal region is stage 5 and the return region is stage 3.
[0082] [0082] These charts and graphs graphically illustrate the value of one or more cooling loops in contributing to the overall energy efficiency of the process. As shown in Table 1, a refrigeration loop decreases the flow usage in the absorption liquor stripper by about 15% compared to operation without refrigeration. Operation with two refrigeration loops reduces the flow consumption by 24% compared to operation without refrigeration; and operation with three loops reduces flow consumption by 25% compared to operation without refrigeration. Without refrigeration, the temperature reaches a maximum of 31 ° C. The maximum temperature drops from 27 ° C, 22.5 ° C and 19 ° C, respectively, with the introduction of one, two or three cooling loops.
[0083] [0083] Compared to the system whose operation is reflected in Figure 7 and Table 1, only a single cooling loop would normally be justified in a sulfur dioxide absorption process that uses a polyprotic acid, such as sodium malate as the sorbent.
[0084] [0084] The remainder of the process, as illustrated in Figure 5 is operated substantially in the manner described above with reference to Figure 1 or Figure 2. However, it should be understood that controlling the increase in temperature in the absorption medium within the absorbent 11 in compliance with the present invention can be practiced independently of providing a source of energy to generate a stripping flow, compressing the gas effluent from the primary stripper or flow generated from the stripped condensate (ie, the process may depend entirely on the refector 37 as an energy source for the absorption liquor removal column 30).
[0085] [0085] When introducing the elements of the present invention or the preferred modality (s) of the same, the articles "one", "one", "the (a)" are intended to mean that there is a or more of the elements. The terms "comprising", "including" and "having" are intended to be inclusive and mean that there may be additional elements other than those listed.
[0086] [0086] In view of the above, it will be seen that the various objects of the invention are obtained and other advantageous results are achieved.
[0087] [0087] As several changes can be made to the previous compositions and processes without deviating from the scope of the invention, it is intended that all the matter contained in the description above should be interpreted as illustrative and not in a limiting sense.

权利要求:
Claims (25)
[0001]
Process for selectively removing and recovering a contaminating gas from a source gas containing contaminant, characterized by the fact that it comprises the following steps: contacting a feed gas stream (10) comprising the source gas in a contaminant gas absorber (11) with an aqueous absorption medium comprising a contaminant gas sorbent, thereby absorbing contaminant gas from the gas stream. feeding into the absorption medium and producing an exhaust gas (19) from which contaminating gas has been removed and an absorption liquor enriched with contaminants (17); contact the contaminant-enriched absorption liquor (17) with stripping steam in an absorption liquor stripper (30) to desorb the contaminant from the contaminated enriched absorption liquor and thereby produce a regenerated contaminant absorption medium ( 15) and a primary stripper gas effluent (33) comprising water vapor and contaminating gas; removing the regenerated contaminant absorption medium (15) from a liquid outlet (35) from the absorption liquor stripper (30) and primary stripper gas effluent (33) from a vapor outlet (34) from the liquor stripper absorption (30); compressing the primary stripper gas effluent (33); condensing water from the compressed primary stripper gas effluent by indirect heat transfer from the compressed primary stripper gas effluent to a cooling medium in a primary stripper gas cooler / condenser (50), thereby producing a condensate containing contaminant ( 59); contacting the condensate containing contaminant (59) exiting the cooler / condenser (50) of primary stripper gas with steam in a condensed stripper (60) to produce a stripped condensate (51) and a condensed stripper gas effluent (61 ) containing water vapor and contaminating gas; wherein the cooling medium to which the heat is transferred from the compressed primary stripper gas effluent in the primary stripper gas cooler / condenser (50) comprises at least part of the stripped condensate (51), thereby generating steam at from the stripped condensate (51) at a pressure higher than the pressure inside the absorption liquor stripper (30) at the liquid outlet (35), and introduce steam generated from the stripped condensate (51) into the cooler / condenser (50) of primary stripper gas in the absorption liquor stripper (30) as stripping vapor for contact with contaminant-enriched absorption liquor (17) to desorb contaminant from it.
[0002]
Process according to claim 1, characterized in that the absorption liquor stripper (30) comprises a column comprising a zone of vertical liquid / vapor contact and steam generated in the primary stripper gas cooler / condenser (50) is introduced in the lower part of the liquid / vapor zone, and absorption liquor enriched with contaminants (17) is introduced in the upper part of the liquid / vapor zone.
[0003]
Process according to claim 2, characterized in that the primary stripper gas effluent (33) is removed from the vapor outlet (34) of the absorption liquor stripper (30) in the upper part of the liquid / contact zone steam, and the absorbent medium of the regenerated contaminant (15) is removed from the liquid outlet (35) of the absorption liquor stripper (30) at the bottom of the liquid / vapor contact zone.
[0004]
Process according to claim 2 or 3, characterized in that the steam generated from the stripped condensate in the primary stripper gas cooler / condenser (50) is introduced at the bottom of the liquid / vapor contact zone of the stripper. absorption liquor (30), and at least part of the vapor condenses within the liquid / vapor contact zone to heat the liquid phase, thereby reducing the equilibrium contaminant concentration in the liquid phase and increasing the driving force for the transfer of contaminants for the vapor phase.
[0005]
Process according to any one of claims 1 to 4, characterized by the fact that it also comprises the circulation of a part of the absorbed medium of the regenerated contaminant (15) removed from the absorption liquor stripper (30) by means of a refiller where it is heated with steam from an external source.
[0006]
Process according to any one of claims 2 to 5, characterized in that the temperature of the steam introduced from the primary stripper gas cooler / condenser into the absorption liquor stripper (30) is not more than 5 ° C at 10 ° C higher than the temperature of the liquid phase inside the absorption liquor stripper (30) at the liquid outlet (35), or at the bottom of the vapor / liquid contact zone.
[0007]
Process according to claim 6, characterized in that the temperature of the vapor introduced from the primary stripper gas cooler / condenser (50) into the absorption liquor stripper (30) is equal to or less than the temperature of the liquid phase within the absorption liquor stripper (30) at the liquid outlet (35), or at the bottom of the vapor / liquid contact zone.
[0008]
Process according to claim 6 or 7, characterized in that the temperature of the vapor introduced from the primary stripper gas cooler / condenser (50) into the absorption liquor stripper (30) varies from the temperature of the liquid phase within the stripper absorption liquor (30) at the liquid outlet (35), or the temperature of the liquid phase at the bottom of the vapor / liquid contact zone, by no more than ± 10 ° C.
[0009]
Process according to any one of claims 1 to 8, characterized by the fact that the difference in mean logarithmic temperature (Δt) in the primary stripper gas cooler / condenser (50) is 1.5 ° C to 10 ° C .
[0010]
Process according to any one of claims 1 to 9, characterized in that the pressure of the primary stripper gas effluent (33) exiting the absorption liquor stripper (30) is 40 and 170 kPa absolute.
[0011]
Process according to any one of claims 1 to 10, characterized in that the compression of the primary stripper gas effluent (33) from the absorption liquor stripper (30) increases its pressure from 30 to 65 kPa.
[0012]
Process according to any one of claims 1 to 11, characterized in that the primary stripper gas effluent (33) from the absorption liquor stripper (30) is compressed by passing through a steam jet ejector and the compressed primary stripper gas is introduced into the primary stripper gas cooler / condenser (50).
[0013]
Process according to any one of claims 1 to 12, characterized in that a final stripper gas is passed through a cutting condenser for the condensation of the water vapor contained therein, the final stripper gas comprising a flow combining condensed stripper gas effluent (61) and a vent gas from the primary stripper gas cooler / condenser (50).
[0014]
Process according to claim 13, characterized in that condensate from the cutting condenser is returned to the condensed stripper (60).
[0015]
Process according to any one of claims 1 to 14, characterized in that the stripped condensate of the condensed stripper (60) is divided to provide: (i) a flow of condensate directed to the gas cooler / condenser (50) primary stripper as a cooling medium for condensing water from the primary stripper gas effluent (33) and generating steam for introduction to the absorption liquor stripper (30); and (ii) a flow of discharge water to remove water from the process.
[0016]
Process according to claim 15, characterized in that at least a part of the discharge water flow is brought into contact with the source gas containing contaminants or the feed gas flow (10) in a saturator upstream of the contaminant gas absorber (11) in relation to the feed gas flow, thereby increasing the moisture in the feed gas flow (10) entering the contaminant gas absorber (11).
[0017]
Process according to claim 16, characterized in that the flow of water leaving the saturator is removed from the process.
[0018]
Process according to any one of claims 1 to 17, characterized in that the regenerated aqueous absorption medium is recirculated to the contaminant gas absorber (11) for further absorption of the additional flow contaminant from the feed gas stream ( 10).
[0019]
Process according to claim 18, characterized in that the recirculated regenerated aqueous absorption medium is passed through an absorption liquor exchanger in the course of the recirculation of the absorption liquor stripper (30) to the contaminant gas absorber ( 11), the contaminant-enriched absorption liquor (17) is passed through the exchanger in the course of transferring the contaminant gas absorber to the absorption liquor stripper (30), and heat is transferred from the regenerated absorption medium to absorption enriched with contaminants (17) inside the exchanger.
[0020]
Process according to any one of claims 1 to 19, characterized in that the exhaust gas (19) from which the contaminant gas has been removed leaves the contaminant gas absorber (11) and is passed through a cut-off cooler. contaminating gas in which heat is transferred from a final stripper gas effluent to the exhaust gas (19), the final stripper gas effluent comprising a flow combining condensed stripper gas effluent (61) and a vent gas of the primary stripper gas cooler / condenser (50).
[0021]
Process according to claim 20 when dependent on claim 14, characterized by the fact that the contaminant gas cutting cooler is upstream of the cutting condenser in relation to the flow of the final stripper gas effluent.
[0022]
Process according to any one of claims 1 to 21, characterized by the fact that the contaminant-enriched gas absorption liquor is circulated between the contaminant gas absorber (11) and an external heat exchanger where the absorption heat is removed by transfer to a coolant.
[0023]
Process according to any one of claims 1 to 22, characterized by the fact that the contaminating gas is selected from the group consisting of SO2, CO2, NOx, H2S, HCl and ammonia.
[0024]
Process for selectively removing and recovering a contaminating gas from a source gas containing contaminants, characterized by the fact that it comprises the following steps: contacting a feed gas stream (10) comprising the source gas in a contaminant absorber (11) with an aqueous absorption medium comprising a contaminant gas sorbent, thereby absorbing the contaminating gas from the feed gas stream in the absorption medium and producing an exhaust gas (19) from which the contaminating gas has been removed and an absorption liquor enriched with contaminants (17); contact the contaminant-enriched absorption liquor (17) with stripping steam in an absorption liquor stripper (30) to desorb the contaminant from the contaminant-enriched absorption liquor and thereby produce a contaminant absorption medium regenerated (15) and a primary stripper gas effluent (33) comprising water vapor and contaminating gas; removing from the regenerated contaminant absorption medium (15) from a liquid outlet (35) from the absorption liquor stripper (30) and primary stripper gas effluent (33) from a vapor outlet (34) from the liquor stripper absorption (30); condensing water from the primary stripper gas effluent by indirect heat transfer from the primary stripper gas effluent (33) to a cooling medium in a primary stripper gas cooler / condenser (50), thereby producing a condensate containing contaminant (59); contacting the condensate containing contaminant (59) exiting the cooler / condenser (50) of primary stripper gas with steam in a condensed stripper (60) to produce a stripped condensate (51) and a condensed stripper gas effluent (61 ) containing water vapor and contaminating gas; wherein the cooling medium to which the heat is transferred from the primary stripper gas effluent in the primary stripper gas cooler / condenser (50) comprises at least part of the stripped condensate (51), thereby generating steam from the stripped condensate (51); compressing the steam generated from the stripped condensate in the cooler / condenser (50) of primary stripper gas at a pressure higher than the pressure inside the absorption liquor stripper (30) at the outlet of the liquid (35) thereof; and introduce the compressed steam into the absorption liquor stripper (30) as stripping vapor for contact with absorption liquor enriched with contaminants (17) to desorb contaminant from it.
[0025]
Process for selectively removing and recovering a contaminating gas from a source gas containing contaminants, characterized by the fact that it comprises the following steps: contacting a feed gas stream (10) comprising the source gas in a contaminant absorber (11) with an aqueous absorption medium comprising a contaminant gas sorbent, thereby absorbing contaminant from the feed gas stream in the absorption medium and producing an exhaust gas (19) from which the contaminant has been removed and an absorption liquor enriched with contaminants (17); contact the contaminant-enriched absorption liquor (17) with stripping steam in an absorption liquor stripper (30) to desorb the contaminant from the contaminant-enriched absorption liquor and thereby produce a contaminant absorption medium regenerated (15) and a primary stripper gas effluent (33) comprising water vapor and contaminating gas; remove regenerated contaminant absorption medium (15) from a liquid outlet (35) from the absorption liquor stripper (33) and primary stripper gas effluent (33) from a vapor outlet (34) from the liquor stripper from absorption (30); condensing water from the primary stripper gas effluent by indirect heat transfer from the primary stripper gas effluent (33) to a cooling medium in a primary stripper gas cooler / condenser (50), thereby producing a condensate contaminant (59); contacting the condensate containing contaminant (59) exiting the cooler / condenser (50) of primary stripper gas with steam in a condensed stripper (60) to produce a stripped condensate (51) and a condensed stripper gas effluent (61 ) containing water vapor and contaminating gas; wherein the cooling medium to which heat is transferred from the primary stripper gas effluent in the primary stripper gas cooler / condenser (50) comprises at least part of the stripped condensate (51), thereby generating steam from the stripped condensate (51); and introduce steam generated from the stripped condensate into the cooler / condenser (50) of primary stripper gas in the absorption liquor stripper (30) as stripping vapor for contact with contaminant-enriched absorption liquor (17) to desorb contaminant from the same.

类似技术:

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BR112014027446B1|2021-04-06|PROCESSES TO SELECTIVELY REMOVE AND RECOVER A CONTAMINATING GAS FROM A GAS OF ORIGIN CONTAINING CONTAMINANTS, AND PROCESS TO REMOVE A CONTAMINATING GAS FROM A GAS OF ORIGIN CONTAINING CONTAMINANTS

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同族专利:

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公开号 | 公开日

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WO2013166301A1|2013-11-07|

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PH12014502451B1|2015-01-12|

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CN104519979A|2015-04-15|

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US20130315807A1|2013-11-28|

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CN104519979B|2017-06-09|

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EA201492008A1|2015-02-27|

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JP6310907B2|2018-04-11|

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PH12014502451A1|2015-01-12|

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EP2844369B1|2020-09-02|

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MX2014013257A|2015-05-15|

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MA20150295A1|2015-08-31|

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AU2013256233B2|2018-02-01|

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TN2014000456A1|2016-03-30|

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MX355791B|2018-04-27|

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CL2014002926A1|2015-07-10|

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MA37598B1|2016-05-31|

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ES2821501T3|2021-04-26|

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EA029381B1|2018-03-30|

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KR20150005675A|2015-01-14|

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MY186373A|2021-07-19|

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US8940258B2|2015-01-27|

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IN2014DN09595A|2015-07-31|

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ZA201408437B|2016-08-31|

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EP2844369A1|2015-03-11|

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KR102096680B1|2020-04-02|

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AU2013256233A1|2014-12-04|

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JP2015519194A|2015-07-09|

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CA2871987C|2018-04-03|

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CA2871987A1|2013-11-07|

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BR112014027446A2|2017-06-27|

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法律状态:
2018-03-06| B06F| Objections, documents and/or translations needed after an examination request according [chapter 6.6 patent gazette]|

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2018-03-13| B06F| Objections, documents and/or translations needed after an examination request according [chapter 6.6 patent gazette]|

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2018-03-20| B06I| Publication of requirement cancelled [chapter 6.9 patent gazette]|Free format text: ANULADA A PUBLICACAO CODIGO 6.6.1 NA RPI NO 2462 DE 13/03/2018 POR TER SIDO INDEVIDA. |

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2019-09-17| B06U| Preliminary requirement: requests with searches performed by other patent offices: procedure suspended [chapter 6.21 patent gazette]|

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2020-09-24| B06A| Patent application procedure suspended [chapter 6.1 patent gazette]|

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2021-02-23| B09A| Decision: intention to grant [chapter 9.1 patent gazette]|

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2021-04-06| B16A| Patent or certificate of addition of invention granted [chapter 16.1 patent gazette]|Free format text: PRAZO DE VALIDADE: 20 (VINTE) ANOS CONTADOS A PARTIR DE 02/05/2013, OBSERVADAS AS CONDICOES LEGAIS. |

优先权:

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申请号 | 申请日 | 专利标题

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US201261641833P| true| 2012-05-02|2012-05-02|

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US61/641,833|2012-05-02|

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PCT/US2013/039293|WO2013166301A1|2012-05-02|2013-05-02|Regenerative recovery of contaminants from effluent gases|

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