WO2001036324A1 - Procede de fabrication du trioxyde de soufre, de l'acide sulfurique, et de l'oleum a partir du dioxyde de soufre - Google Patents

Procede de fabrication du trioxyde de soufre, de l'acide sulfurique, et de l'oleum a partir du dioxyde de soufre Download PDF

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Publication number
WO2001036324A1
WO2001036324A1 PCT/US2000/030095 US0030095W WO0136324A1 WO 2001036324 A1 WO2001036324 A1 WO 2001036324A1 US 0030095 W US0030095 W US 0030095W WO 0136324 A1 WO0136324 A1 WO 0136324A1
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Prior art keywords
gas
absoφtion
sulfuric acid
set forth
zone
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PCT/US2000/030095
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English (en)
Inventor
Adam V. Menon
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Monsanto Company
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Application filed by Monsanto Company filed Critical Monsanto Company
Priority to EA200200516A priority Critical patent/EA200200516A1/ru
Priority to BR0015265-0A priority patent/BR0015265A/pt
Priority to EP00975531A priority patent/EP1230150A1/fr
Priority to JP2001538282A priority patent/JP2003517419A/ja
Priority to CA002387988A priority patent/CA2387988A1/fr
Priority to AU13573/01A priority patent/AU1357301A/en
Priority to MXPA02004408A priority patent/MXPA02004408A/es
Priority to KR1020027005602A priority patent/KR20020049001A/ko
Publication of WO2001036324A1 publication Critical patent/WO2001036324A1/fr

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    • C—CHEMISTRY; METALLURGY
    • C01—INORGANIC CHEMISTRY
    • C01B—NON-METALLIC ELEMENTS; COMPOUNDS THEREOF; METALLOIDS OR COMPOUNDS THEREOF NOT COVERED BY SUBCLASS C01C
    • C01B17/00—Sulfur; Compounds thereof
    • C01B17/48—Sulfur dioxide; Sulfurous acid
    • C01B17/50—Preparation of sulfur dioxide
    • C01B17/60—Isolation of sulfur dioxide from gases
    • C—CHEMISTRY; METALLURGY
    • C01—INORGANIC CHEMISTRY
    • C01B—NON-METALLIC ELEMENTS; COMPOUNDS THEREOF; METALLOIDS OR COMPOUNDS THEREOF NOT COVERED BY SUBCLASS C01C
    • C01B17/00—Sulfur; Compounds thereof
    • C01B17/69—Sulfur trioxide; Sulfuric acid
    • C01B17/74—Preparation
    • C01B17/76—Preparation by contact processes
    • C01B17/765—Multi-stage SO3-conversion

Definitions

  • This invention relates to a novel process for preparing sulfur trioxide (SO 3 ) by oxidizing sulfur dioxide (SO 2 ).
  • This invention also relates to a process for preparing liquid sulfuric acid (H 2 SO 4 ) and/or oleum from SO 3 by the contact process, wherein SO 2 is oxidized to form SO 3 , which, in turn, is contacted with water or a solution of sulfuric acid to produce additional sulfuric acid and/or oleum.
  • This invention further relates to recovering high grade energy from the heat produced during such a contact process.
  • Sulfuric acid is the highest volume chemical manufactured in the world. Much of the sulfuric acid is used to produce phosphoric acid in integrated fertilizer complexes. Sulfuric acid is also used, for example, in dyes and pigments, industrial explosives, etching applications, alkylation catalysis, electroplating baths, and nonferrous metallurgy. Current worldwide production is reported to be about 570,000 tons per day, with about 30% being produced in the United States.
  • the contact process has been one of the most popular methods for making sulfuric acid and oleum ("oleum” is a solution of SO 3 in sulfuric acid, and also is known as “fuming sulfuric acid” or “H 2 S 2 O 7 ").
  • This process generally comprises 3 steps: (1) forming SO 2 from a sulfur-containing raw material, (2) catalytically oxidizing the SO 2 to form SO3 and (3) contacting the SO 3 with water or concentrated sulfuric acid to hydrate the SO 3 and form sulfuric acid and/or ileum.
  • a wide variety of sulfur-containing raw materials have been used in the contact process to form SO 2 .
  • Most sulfuric acid plants for example, form SO 2 by oxidizing an oxidizable sulfur-containing material (e.g., elemental sulfur or metal ores containing sulfides) in a thermal combustion zone.
  • an oxidizable sulfur-containing material e.g., elemental sulfur or metal ores containing sulfides
  • a significant number of other plants e.g., sulfuric acid regeneration plants
  • burn a carbonaceous material i.e., a fuel
  • a decomposable sulfate to provide the heat necessary to decompose the sulfate into SO 2 and various byproducts.
  • the SO 2 is normally oxidized to SO 3 by contacting it with a catalyst (e.g., a vanadium pentoxide (N 2 O 5 ) catalyst) at a temperature effective for catalytic oxidation of SO 2 (e.g., at least about 410 to about 420°C for a N 2 O 5 catalyst) in the presence of molecular oxygen.
  • a catalyst e.g., a vanadium pentoxide (N 2 O 5 ) catalyst
  • a temperature effective for catalytic oxidation of SO 2 e.g., at least about 410 to about 420°C for a N 2 O 5 catalyst
  • This reaction is often conducted in a catalytic converter which comprises a plurality of catalyst beds in series (conventionally, 4 or more catalyst beds).
  • a catalytic converter which comprises a plurality of catalyst beds in series (conventionally, 4 or more catalyst beds).
  • the reaction conditions be controlled so that the heat evolved from the oxidation reaction does not overheat the catalyst to a temperature which may lead to thermal damage and premature deactivation of the catalyst and/or adversely affect the reaction equilibrium.
  • the oxidation reaction can be controlled, for example, by limiting the concentration of SO 2 or oxygen fed into the catalytic converter, or by using a converter comprising a tube-in-shell device such as that disclosed by Daley et al. in U.S. Patent No. 4,643,887 wherein the catalyst is cooled by indirect heat exchange with a cooling medium(e.g., air or molten salts).
  • a cooling medium e.g., air or molten salts
  • sulfuric acid and/or oleum is normally conducted in an absorption zone within an SO 3 absorption tower, in which the conversion gas containing the SO 3 is contacted with water, or, more typically, a concentrated solution of sulfuric acid (e.g., a solution containing about 98.5 weight% sulfuric acid) to form sulfuric acid and/or oleum.
  • Water is normally less preferred because it tends to form an acid mist of H 2 SO 4 that is difficult to condense.
  • Tail-gas scrubbers have been particularly useful in conjunction with low-conversion, single-stage SO 3 absorption plants.
  • a number of SO 2 tail-gas scrubbing processes are available, many of which use non-regenerable scrubbing mediums such as ammonia, sodium hydroxide, or hydrogen peroxide.
  • Such techniques have various disadvantages. For example, they require expensive equipment (e.g., a separate scrubbing tower). Such equipment takes up valuable space and produces an additional pressure drop in the overall gas system, which decreases the gas handling capacity of the system.
  • the scrubbing processes using a base often produce a by-product which must be properly disposed of (e.g., when ammonia is used to scrub the tail gas, a side stream of ammonium sulfate is produced; and when sodium hydroxide is used, a side stream of sodium sulfate is produced).
  • ammonium salt scrubbing solutions typically results in the formation of submicron aerosol fumes which must be removed using sophisticated and expensive mist eliminators.
  • Sulfuric acid plants have also controlled SO 2 emissions by using a dual SO 3 absorption process.
  • an SO 3 absorption tower containing an intermediate SO 3 absorption zone is positioned between two of the catalyst beds of the converter.
  • gas exiting the second or third catalyst bed is passed through an intermediate SO 3 absorption zone wherein the gas is contacted with a concentrated solution of H 2 SO 4 to form product acid.
  • Gas exiting the intermediate SO 3 absorption zone is returned to the next bed of the converter.
  • This invention provides for an improved process for making SO 3 which comprises oxidizing SO 2 in a catalytic converter. More particularly, this invention provides for a process for making SO 3 which can be implemented with relatively low capital and operating costs; a process for making SO 3 which allows for a minimal volume of gas to be handled upstream of the catalytic converter, thus allowing for smaller equipment (i.e., equipment having lower capital and operating costs) to be used upstream of the converter; a process for making SO 3 from an SO 2 source gas that has a relatively low SO 2 gas strength; a process for making SO 3 wherein the catalytic converter can be operated without the use of an extraneous energy source to bring the
  • the present invention is directed to a process for making SO 3 wherein the catalytic converter may be operated autothermally even when a weak source gas (e.g., a source gas having an SO 2 concentration of less than about 5 mole%) is used; a process for making SO 3 from spent sulfuric acid; a process for making SO 3 from sulfidic metal oxidation off gases; and a process for the production of sulfuric acid and/or ileum wherein the recovery of heat energy is enhanced.
  • a weak source gas e.g., a source gas having an SO 2 concentration of less than about 5 mole%
  • the process comprises contacting the source gas with a liquid SO 2 absorption solvent in an SO 2 absorption zone to selectively transfer SO 2 from the source gas to the SO 2 absorption solvent and form an SO 2 -depleted gas and an SO 2 -enriched solvent.
  • Sulfuric dioxide is then stripped from the SO 2 -enriched solvent in an SO 2 stripping zone to form an SO 2 - depleted absorption solvent and an SO 2 -enriched stripper gas having an SO 2 gas strength greater than the SO 2 gas strength of the source gas.
  • a reaction gas comprising a first portion of the SO 2 -enriched stripper gas is then formed.
  • An oxidation product gas (comprising SO 3 and residual SO 2 ), in turn, is formed by a process comprising passing the reaction gas through a plurality of catalyst beds in series (this plurality comprises at least 2 and no greater than 4 catalyst beds which contain a catalyst effective for oxidizing SO 2 into SO 3 ).
  • a second portion of the SO 2 -enriched gas is introduced into at least one catalyst bed downstream of the most upstream catalyst bed to increase the amount of SO 2 being fed into the downstream bed.
  • the process comprises contacting the source gas with a liquid SO 2 absorption solvent in an SO 2 absorption zone to selectively transfer SO 2 from the source gas to the SO 2 absorption solvent and form an SO 2 -depleted gas and an SO 2 -enriched solvent.
  • Sulfur dioxide is then stripped from the SO 2 -enriched solvent in an SO 2 stripping zone to form an SO 2 -depleted absorption solvent and an SO 2 -enriched stripper gas having an SO 2 gas strength greater than the SO 2 gas strength of the source gas.
  • a reaction gas is then formed which comprises a first portion of the SO 2 -enriched stripper gas (this first portion comprises at least about 30% of the SO 2 in the SO 2 -enriched stripper gas).
  • an oxidation product gas (comprising SO 3 and residual SO 2 ) is formed by a process comprising passing the reaction gas through a plurality of catalyst beds in series (this plurality comprises at least 2 catalyst beds which contain a catalyst effective for oxidizing SO 2 into SO 3 ).
  • this plurality comprises at least 2 catalyst beds which contain a catalyst effective for oxidizing SO 2 into SO 3 ).
  • a second portion of the SO 2 -enriched gas is introduced into at least one catalyst bed downstream of the most upstream catalyst bed to increase the amount of SO 2 being fed into the downstream bed.
  • the process comprises contacting the source gas with a liquid SO 2 absorption solvent in an SO 2 absorption zone to selectively transfer SO 2 from the source gas to the SO 2 absorption solvent and form an SO 2 -depleted gas and an SO 2 -enriched solvent.
  • Sulfur dioxide is then stripped from the SO 2 -enriched solvent in an SO 2 stripping zone to form an SO 2 -depleted absorption solvent and an SO 2 -enriched stripper gas having an SO 2 gas strength greater than the SO 2 gas strength of the source gas.
  • a reaction gas is then formed which comprises a first portion of the SO 2 -enriched stripper gas.
  • an oxidation product gas (comprising SO 3 and residual SO 2 ) is formed by a process comprising passing the reaction gas through a plurality of catalyst beds in series (this plurality comprises at least two catalyst beds which comprise a catalyst effective for oxidizing SO 2 into SO 3 ).
  • this plurality comprises at least two catalyst beds which comprise a catalyst effective for oxidizing SO 2 into SO 3 ).
  • a second portion of the SO 2 -enriched gas is introduced into at least one catalyst bed downstream of the most upstream catalyst bed to increase the amount of SO 2 being fed into the downstream bed.
  • the molar ratio of O 2 to SO 2 is greater than about 0.2:1 in the gas entering each of the catalyst beds in the plurality.
  • the process comprises contacting the source gas with a liquid SO 2 absorption solvent in an SO 2 absorption zone to selectively transfer SO 2 from the source gas to the SO 2 absorption solvent and form an SO 2 -depleted gas and an SO 2 -enriched solvent.
  • Sulfur dioxide is then stripped from the SO 2 -enriched solvent in an SO 2 stripping zone to form an SO 2 -depleted absorption solvent and an SO 2 -enriched stripper gas.
  • a converter feed gas is formed which comprises a first portion of the SO 2 -enriched stripper gas. This converter feed gas is divided into a first portion and a second portion.
  • a first partial conversion gas and a second partial conversion gas are then formed by passing the first portion of the converter feed gas through a catalyst bed, and passing the second portion through a different catalyst bed in parallel with the catalyst bed through which the first portion of the converter feed gas is passed (both catalyst beds comprise an oxidation catalyst effective for oxidizing SO 2 to SO 3 ).
  • a first portion of the remainder of the SO 2 - enriched stripper gas is then combined with the first partial conversion gas to fortify the SO 2 gas strength of the first partial conversion gas.
  • a second portion of the remainder of the SO 2 -enriched stripper gas is combined with the second partial conversion gas to fortify the SO 2 gas strength of the second partial conversion gas.
  • the fortified first partial conversion gas and the fortified second partial conversion gas are then passed through at least one further catalyst bed (also comprising an oxidation catalyst effective for oxidizing SO 2 to SO 3 ), thereby oxidizing additional SO 2 to SO 3 and forming a conversion gas comprising SO 3 and SO 2 .
  • This invention also provides for an improved process for making sulfuric acid and/or oleum. More particularly, this invention provides for a process for making sulfuric acid and/or oleum which meets SO 2 emissions standards; a process for making sulfuric acid and/or oleum having greater SO 2 oxidation capacity than typical conventional sulfuric acid plants without having greater SO 2 emissions; a process for making sulfuric acid and/or oleum in which SO 2 emissions are confined to a single purge stream for simple control and monitoring; a process for making sulfuric acid and/or oleum which achieves at least about 99.7% recovery of SO 2 , even at low single pass SO 2 conversions (e.g., SO 2 single-pass conversions of as low as about 75% or lower); a process for making sulfuric acid and/or oleum which can be implemented with relatively low capital and operating costs; a process for making sulfuric acid and/or oleum which allows for a lesser volume of gas to be handled upstream of the catalytic converter than typical conventional sulfuric acid contact
  • the present invention is directed to a process for making sulfuric acid and/or oleum from a source gas comprising SO 2 .
  • the process comprises contacting at least a portion of the source gas with a liquid SO 2 absorption solvent in an SO 2 absorption zone to selectively transfer SO 2 from the portion of the source gas to the SO 2 absorption solvent and form an SO 2 -depleted gas and an SO 2 -enriched solvent.
  • Sulfur dioxide is then stripped from the SO 2 -enriched solvent in an SO 2 stripping zone to form an SO 2 -depleted absorption solvent and an SO 2 -enriched stripper gas having an SO 2 gas strength greater than the SO 2 gas strength of the source gas.
  • An oxidation product gas (comprising SO 3 and residual SO 2 ) is then formed by a process comprising passing the SO 2 -enriched stripper gas through a plurality of catalyst beds in series (each comprising an oxidation catalyst effective for oxidizing SO 2 to SO 3 ).
  • the oxidation product gas is combined with water vapor to form an acid product gas comprising: (a) sulfuric acid formed by a gas phase reaction between SO 3 from the oxidation product gas and water vapor, thereby generating the heat of formation of sulfuric acid in the gas phase; (b) SO 3 ; and (c) SO 2 .
  • Heat energy from the gas phase heat of formation of sulfuric acid is recovered by transfer of heat from the acid product gas to steam or feed water in an indirect heat exchanger.
  • the process comprises forming an oxidation product gas (comprising SO 3 and residual SO 2 ) by a process comprising passing a first portion of the source gas through a plurality of catalyst beds in series (this plurality comprises at least 2 catalyst beds which contain a catalyst effective for oxidizing SO 2 into SO 3 ).
  • a second portion of the source gas is introduced into at least one catalyst bed downstream of the most upstream catalyst bed to increase the amount of SO 2 being fed into the downstream bed.
  • the oxidation product gas is combined with water vapor to form an acid product gas comprising: (a) sulfuric acid formed by a gas phase reaction between SO 3 from the oxidation product gas and water vapor, thereby generating the heat of formation of sulfuric acid in the gas phase; (b) SO 3 ; and (c) SO 2 .
  • Heat energy from the gas phase heat of formation of sulfuric acid is recovered by transfer of heat from the acid product gas to steam or feed water in an indirect heat exchanger.
  • the cooled acid product gas is then contacted with liquid sulfuric acid in an SO 3 absorption zone to form additional sulfuric acid and/or oleum and an SO 3 - depleted gas comprising SO 2 .
  • Another embodiment of this invention is directed to an improved process for making sulfuric acid and/or oleum from a source gas comprising SO 2 and water vapor.
  • This process comprises forming a reaction gas comprising SO 2 , and then forming an oxidation product gas (comprising SO 3 and residual SO 2 ) by a process comprising passing the reaction gas through a plurality of catalyst beds in series (each catalyst bed comprises an oxidation catalyst effective for oxidizing SO 2 into SO 3 ).
  • the oxidation product gas is combined with water vapor to form an acid product gas comprising: (a) sulfuric acid formed by a gas phase reaction between SO 3 from the oxidation product gas and water vapor, thereby generating the heat of formation of sulfuric acid in the gas phase; (b) SO 3 ; and (c) SO 2 .
  • Heat energy is recovered from the gas phase heat of formation of sulfuric acid by transferring heat from the acid product gas to steam or feed water in an indirect heat exchanger.
  • the cooled acid product gas is contacted with a solution comprising sulfuric acid in an SO 3 absorption zone to form additional sulfuric acid and/or oleum and an SO 3 -depleted gas comprising SO 2 .
  • the improvement in this process comprises combining at least a portion of the source gas with the oxidation product gas to form the acid product gas, and forming the reaction gas from the SO 3 -depleted gas.
  • Fig. 1 is a schematic flow sheet illustrating various features of one embodiment of the process of the present invention.
  • Fig. 2 is a schematic flow sheet showing a 4 bed catalytic converter of a contact sulfuric acid plant modified in accordance with the present invention.
  • Fig. 3 is a schematic flow sheet illustrating various features of another embodiment of the process of the present invention for use with a wet SO 2 source gas.
  • Fig. 4 is a schematic flow sheet illustrating an embodiment of the process of the present invention described in the Example below.
  • a source gas 3 containing SO 2 is formed from a sulfur- containing raw material 6.
  • a decomposable sulfate is often suitable.
  • Such a sulfate may include, for example, calcium sulfate, ammonium sulfate, or spent H 2 SO 4 (i.e., contaminated or diluted H 2 SO 4 ).
  • the sulfate is typically injected as a liquid spray into a combustion zone 9, along with a carbonaceous material (i.e., a fuel) and an oxygen source 12 (normally air).
  • a gas 3 is formed which typically contains sulfurous acid (H 2 SO 3 ), SO 2 , O 2 , CO 2 , N 2 , and water vapor.
  • the sulfur containing raw material 6 is an oxidizable material, such as elemental sulfur, hydrogen sulfide (H 2 S), or iron pyrite (FeS 2 ) or another sulfide-containing metal ore.
  • the sulfur-containing material 6 is typically burned with an oxygen source 12 in a kiln or other suitable thermal combustion zone 9 to produce a source gas 3 containing SO 2 .
  • the most economically practical oxygen source 12 is normally air, which, when burned with the oxidizable sulfur material 6, produces a source gas 3 containing SO 2 , O 2 , and N 2 (and water vapor if, for example, the air and/or the raw sulfur material contains water, or the sulfur-containing raw material is H 2 S).
  • the process of this invention may be practiced with a wide range of SO 2 concentration in the source gas 3 (i.e., the source gas 3 may contain from about 0.1 to about 100 mole% SO 2 ).
  • the process is used in conjunction with other manufacturing processes which either need to reduce or eliminate the sulfur content in a particular material, or need to reduce or eliminate a sulfur-containing material in a waste stream.
  • this process provides, for example, a practical way to utilize the SO 2 which is produced as an off-gas when a metal ore is roasted or smelted during a metal recovery operation.
  • This process also, for example, provides a practical way to utilize spent H 2 SO 4 .
  • the SO 2 concentration in the source gas 3 is typically less than about 11 mole%, and more typically from about 0.1 to about 5 mole%. Because there are greater operational and capital costs associated with larger process equipment, it is often preferable to minimize the volume of the SO 2 source gas 3, while also increasing the concentration of SO 2 in the source gas 3. In a particularly preferred embodiment, this is achieved by using elemental sulfur as the raw material 6. When elemental sulfur is burned in air, for example, SO 2 concentrations of from about 11 to about 21 mole% (and more typically, from about 15 to about 20 mole%) may be obtained.
  • the amount of the oxygen source 12 fed into the combustion zone 9 of the sulfur burner preferably is the amount necessary to maintain the molar ratio of O 2 to elemental sulfur at slightly greater than about 1.0, more preferably from about 1.05 to about 1.3, and most preferably about 1.05 to about 1.1. In most embodiments, it is prefe ⁇ ed for the O 2 concentration in the source gas 3 to be from about 0.5 to about 5 mole%, more preferably from about 0.5 to about 3 mole%, and most preferably from about 0.5 to about 2 mole%.
  • the SO 2 -containing source gas 3 is preferably introduced into an SO 2 absorption/stripping zone to remove and recover SO 2 in the form of an SO 2 -enriched gas (i.e., a gas having an increased SO 2 content relative to the source gas 3).
  • an SO 2 -enriched gas i.e., a gas having an increased SO 2 content relative to the source gas 3.
  • the source gas 3 is at an elevated temperature (i.e., greater than about 50°C) and/or contains entrained particulate impurities, it is generally prefe ⁇ ed to first condition the source gas 3 to cool the gas 3 and remove particulates from the gas 3 before introducing it into the SO 2 absorption/desorption zone.
  • elevated temperature i.e., greater than about 50°C
  • the source gas 3 is a combustion gas exiting a sulfur burner, its temperature is typically from about 900 to about 1600°C, and more typically from about 1050 to about 1600°C.
  • This gas 3 may, for example, be cooled by: (a) passing the gas 3 through an indirect heat exchanger where heat from the gas 3 is used, for example, to the preheat the oxygen source 12 (e.g., air) being used in the combustion chamber 9, thereby reducing fuel costs in heating the oxygen source 12 with an external source; (b) by passing the gas 3 through a waste heat boiler where it is cooled by generation of high pressure steam (i.e., steam having a pressure of at least about 27 bar (gauge)); and/or (c) passing the gas 3 through a humidifying tower and one or more indirect heat exchangers, where it is further cooled with, for example, cooling tower water.
  • an indirect heat exchanger where heat from the gas 3 is used, for example, to the preheat the oxygen source 12 (e.g., air) being used in the combustion chamber 9, thereby reducing fuel costs in heating the oxygen source 12 with an external source
  • a waste heat boiler where it is cooled by generation of high pressure steam (i.e., steam having a
  • the source gas 3 is formed from spent sulfuric acid or is the off-gas from a metal roasting or smelting operation
  • an electrostatic precipitator is often used to remove particulates from the gas after it is cooled.
  • a gas 3 may be conditioned by passing the gas 3 through one or more reverse jet scrubbers of the type, for example, sold by Monsanto Enviro-Chem Systems, Inc. (St. Louis, MO, USA) under the trademark "DYNAWAVE". It should be noted that a portion (e.g., 5-10%) of the conditioned source gas 3 may be recycled back to the combustion zone 9 (particularly a sulfur burner) to control the temperature in the combustion zone 9 below a desired maximum temperature.
  • the SO 2 - containing source gas 3 is contacted with a liquid SO 2 absorption solvent 15 in an SO 2 absorption zone 18.
  • the liquid SO 2 absorption solvent 15 selectively absorbs SO 2 from the source gas 3, thereby transferring SO 2 from the source gas 3 to the SO 2 absorption solvent 15 and producing an SO 2 -depleted exhaust stripper gas 21 (from which the SO 2 has been substantially removed) and an SO 2 -enriched absorption solvent 24.
  • the liquid SO 2 abso ⁇ tion solvent 15 may be either a physical or a chemical solvent. Physical solvents, however, are generally more prefe ⁇ ed.
  • Suitable absorbents include various organic absorbents (e.g., tetraethylene glycol dimethyl ether), and aqueous solutions of alkali metals (e.g., a sodium sulfite/bisulfite solution).
  • a suitable physical sulfur dioxide abso ⁇ tion solvent is one comprising tetra ethylene glycol diethel ether such as that disclosed and utilized in the sulfur dioxide recovery processes described in U.S. Patent No. 4,659,553 (Line) and U.S. Patent No. 4,795,553 (Hensel et al.), the entire disclosures of which are inco ⁇ orated herein by reference.
  • the liquid sulfur dioxide absorbent preferably contains more than 50% by weight tetra ethylene glycol diethel ether.
  • Such a liquid sulfur dioxide absorbent suitably comprises, on a dry weight basis, from about 60% to about 80% tetra ethylene glycol diethel ether, from about 15% to about 25% triethylene glycol diethel ether, from about 2.5% to about 7.5% pentaethylene glycol diethel ether and from about 2.5% to about 7.5% mono ethers.
  • the circulating tetra ethylene glycol diethel ether-containing absorbent may contain water, for example, up to about 10% by weight.
  • Suitable SO 2 abso ⁇ tion solvents include aqueous solutions of various amines.
  • Exemplary amine absorbing agents include, for example, aniline derivatives (e.g., dimethylaniline), alkanolamines (e.g., diethanolamine, triethanolamine, tripropanolamine, and tributanolamine), tetrahydroxyethylalkylenediamines (e.g * ., tetrahydroxymethylenediamine, tetrahydroxyethylethylenediamine, tetrahydroxyethyl- 1 ,3-propylenediamine, tetrahydroxyethyl- 1 ,2-propylenediamine, tetrahydroxyethyl- 1 ,5- pentylpentylenediamine), and heterocyclic diamines (e.g., piperazine; dimethylpiperazine; N,N'-bis(2-hydroxyethyl)piperazine; -methylpy ⁇ oli
  • An even more prefe ⁇ ed traditionally used absorbing agent is a half salt of a diamine having the following formula (I):
  • R 1 , R 2 , R 3 , and R 4 may be the same or different, and can be hydrogen, alkyl (preferably having from 1 to about 8 carbon atoms, and including cycloalkyls), hydroxyalkyl (preferably having from 2 to about 8 carbon atoms), aralkyl (preferably having from about 7 to about 20 carbon atoms), aryl (preferably monocyclic or bicyclic), or alkylaryl (preferably having from about 7 to about 20 carbon atoms). It should be noted that any of R 1 , R 2 , R 3 , and R 4 may together form cyclic structures.
  • the free nitrogen of the half salt preferably has a pKa of from about 4.5 to about 7.3.
  • particularly prefe ⁇ ed diamines include the sulfite half salts of N,N' ,N' -(trimethyl)-N(2-hydroxyethyl)ethylenediamine; N,N,N' ,N' -tetramethylethylenediamine; N,N,N',N' -tetrakis(2- hydroxyethyl)efhylenediamine; N-(2-hydroxyethyl)ethylenediamine; N,N' - dimethylpiperazine; N,N,N',N'-tetrakis(2-hydroxyethyl)-l ,3-diaminopropane; and N,N'-dimethyl-N,N-bis(2-hydroxyethyl)ethylenediamine.
  • These half-salt diamine absorbents are described by Hakka in U.S. Patent No. 5,0
  • the abso ⁇ tion solvent 15 preferably comprises an aqueous solution containing from about 20 to about 40 weight% of the absorbing agent on an amine (rather than an amine salt) basis.
  • the SO 2 -enriched abso ⁇ tion solvent 24, in turn, preferably has an SO 2 /amine-absorbing-agent weight ratio of from about 0.1 :1 to about 0.25:1.
  • the above-listed traditional SO 2 absorbents are often hampered by one or more shortcomings. These shortcomings include, for example, relatively low SO 2 abso ⁇ tion capacity and the tendency to absorb substantial quantities of water vapor from the source gas 3. Abso ⁇ tion of substantial quantities of water, in turn, can lead to a significant reduction in the SO 2 abso ⁇ tion capacity of the SO 2 abso ⁇ tion solvent 15, thereby requiring a greater flow of the SO 2 abso ⁇ tion solvent 15. Such water abso ⁇ tion can also lead to excessive co ⁇ osion of the equipment used in the SO 2 abso ⁇ tion/stripping process.
  • the SO 2 abso ⁇ tion solvent 15 comprises an organic phosphorous compound, as described in U.S. Patent No. 5,851,265 (Burmaster et al.) which the entire disclosure is herein inco ⁇ orated by reference).
  • the SO 2 abso ⁇ tion solvent 15 preferably comprises a phosphate triester, phosphonate diester, phosphinate monoester, or a mixture thereof.
  • the substituents bonded to the phosphorous atom, as well as the organic radicals of the ester functionality, in the compounds are preferably independently aryl or C, to C g alkyl (i.e., an alkyl group containing from 1 to 8 carbon atoms).
  • suitable phosphate triesters include: tributyl phosphate, tripentyl phosphate, trihexyl phosphate, and triphenyl phosphate.
  • suitable phosphinate monoesters include: butyl dibutyl phosphinate, pentyl dipentyl phosphinate, hexyl dihexyl phosphinate, and phenyl diphenyl phosphinate.
  • the SO 2 abso ⁇ tion solvent 15 comprises at least one substantially water- immiscible organic phosphonate diester having formula (II)
  • R 1 , R 2 , and R 3 are independently aryl or C, to C 8 alkyl, with R 1 , R 2 , and R 3 being selected such that (1) the organic phosphonate diester has a vapor pressure of less than about 1 Pa at 25° C, and (2) the solubility of water in the organic phosphonate diester is less than about 10 weight% at 25° C.
  • the organic phosphonate is a dialkyl alkyl phosphonate, and R 1 , R 2 , and R 3 are independently C, to C 6 alkyl.
  • R 1 , R 2 , and R 3 are identical, with each containing at least 4 carbon atoms.
  • suitable organic phosphonate deters for use in the practice of the present invention include dibutyl butyl phosphonate, dipentyl pantile phosphonate, dihexyl hassle phosphonate and diphenyl phenyl phosphonate.
  • the SO 2 abso ⁇ tion solvent 15 comprises dibutyl butyl phosphonate.
  • Dibutyl butyl phosphonate is a neutral diester of phosphonic acid, and is a clear, colorless liquid with a relatively low viscosity and very mild odor.
  • Dibutyl butyl phosphonate has a molecular weight of 250.3 and a vapor pressure of about 0.1 Pa at 25 ° C.
  • the solubility of water in dibutyl butyl phosphonate is about 5.5 weight% at 25° C.
  • An SO 2 abso ⁇ tion solvent comprising at least one organic phosphonate diester as defined above tends to be more prefe ⁇ ed because such a solvent typically possesses a combination of characteristics which renders it particularly useful in an SO 2 abso ⁇ tion/deso ⁇ tion process, including: (1) increased SO 2 solubility, especially at low partial pressures of SO 2 in the source gas 3; (2) high heats of solution, which reduce the amount of energy required for stripping SO 2 from the SO 2 -enriched abso ⁇ tion solvent 24; (3) low melting points, so that the solvent 15 will remain a liquid over a wide range of process temperatures; (4) low viscosity, which allows the size of both thermal and abso ⁇ tion/stripping equipment to be reduced; (5) low vapor pressure, which reduces solvent 15 losses; (6) decreased tendency to react with water and undergo hydrolysis; and (7) being substantially water immiscible (i.e., non- hygroscopic) such that the solubility of water in the solvent 15 is decreased.
  • the SO 2 abso ⁇ tion zone 18 preferably comprises a means for promoting mass transfer between the gas and liquid phases, and more preferably comprises a bed of random packings such as saddles or rings in a vertical tower.
  • the source gas 3 is contacted countercu ⁇ ently with the SO 2 abso ⁇ tion solvent 15.
  • the source gas 3 is preferably introduced through an inlet near the bottom of the SO 2 abso ⁇ tion zone 18, and the SO 2 abso ⁇ tion solvent 15 is introduced through an inlet near the top of the SO 2 abso ⁇ tion zone 18 and distributed over the packing.
  • the SO 2 -enriched abso ⁇ tion solvent 24 is then withdrawn from an outlet near the bottom of the SO 2 abso ⁇ tion zone 18, and the exhaust gas substantially free of SO 2 (i.e., the SO 2 -depleted gas 21) is removed from an outlet near the top of the SO 2 abso ⁇ tion zone 18.
  • the SO 2 abso ⁇ tion zone 18 may comprise a conventional, randomly packed tower, those skilled in the art will appreciate that other configurations may be suitably used as well.
  • the tower may contain structured packing or comprise a tray tower, in either of which the process streams preferably flow countercu ⁇ ently.
  • the SO 2 abso ⁇ tion zone 18 preferably is operated at an average temperature of from about 10 to about 60°C (more preferably from about 10 to about 50°C, and most preferably from about 30 to about 40°C), and a pressure of from about 50 to about 150 kPa (absolute). It should be recognized that although pressure increases the amount of SO 2 that the SO 2 abso ⁇ tion solvent 15 can absorb, the abso ⁇ tion can alternatively be carried out at a relatively low pressure, thereby reducing equipment costs.
  • the temperature of the solvent 15 introduced into the abso ⁇ tion zone 18 preferably is above the dew point temperature of the source gas 3 fed into the abso ⁇ tion zone 18.
  • the mass flow rate ratio (L/G) of the SO 2 abso ⁇ tion solvent 15 and the source gas 3 necessary to achieve substantial transfer of SO 2 from the source gas 3 to the SO 2 abso ⁇ tion solvent 15 in the abso ⁇ tion zone 18 may be determined by conventional design practice.
  • the SO 2 abso ⁇ tion zone 18 is designed and operated such that the SO 2 content of the SO 2 -depleted gas 21 is less than about 400 ppmv, more preferably less than about 200 ppmv, and most preferably less than about 150 ppmv.
  • the SO 2 -depleted gas 21 may be passed through a mist eliminator for recovery of entrained liquid before being discharged through a stack.
  • Use of the highly efficient organic phosphorous solvents discussed above allows the concentration of the SO 2 in the SO 2 -enriched stripper gas 30 exiting the stripper zone 27 to be significantly greater than the concentration of the SO 2 in the source gas 3 fed to the system.
  • the process of the present invention may be operated such that the ratio of the SO 2 molar concentration in the in the SO 2 -enriched stripper gas 30 to the SO 2 molar concentration in the source gas 3 is greater than about 1.1:1, preferably at least about 2.75 : 1 , more preferably at least about 4:1, even more preferably at least about 7:1, and most preferably at least about 10:1. It should be recognized that even greater ratios may often be achieved, depending on the SO 2 concentration of the source gas 3.
  • At least 67 mole% (more preferably at least about 75 mole%, still more preferably at least about 85 mole%, and most preferably at least about 90 mole%) of the SO 2 -enriched stripper gas 30 consist of SO 2 .
  • SO 2 may be stripped by contacting the SO 2 -enriched abso ⁇ tion solvent 24 with a non-condensable, oxygen-containing stripping gas 36 such that SO 2 is transfe ⁇ ed from the SO 2 -enriched abso ⁇ tion solvent 24 to the stripping gas 36 to produce the SO 2 -enriched stripper gas 30 and the SO 2 -depleted abso ⁇ tion solvent 33.
  • the non-condensable, oxygen-containing stripping gas 36 comprises air.
  • solvents comprising organic phosphorous compounds are their inherent flame retarding property and resistance to oxidation.
  • organic solvents used in conventional SO 2 abso ⁇ tion/deso ⁇ tion cycles e.g., tetraethylene glycol dimethyl ether
  • the organic solvents utilized in the present invention can be readily stripped of SO 2 using an oxygen-containing stripping gas with minimal risk of solvent degradation or explosion.
  • the SO 2 stripper zone 27 preferably comprises a means for promoting mass transfer between the gas and liquid phases. Like the SO 2 abso ⁇ tion zone 18, the SO 2 stripper zone 27 preferably comprises a bed of conventional random packing in a vertical tower. To maximize transfer of SO 2 , the SO 2 -enriched abso ⁇ tion solvent 24 is preferably contacted countercu ⁇ ently with the SO 2 stripping gas 36. In this embodiment, a non-condensable, oxygen-containing SO 2 stripping gas 36 preferably is introduced through an inlet near the bottom of the SO 2 stripper zone 27, and the SO 2 -enriched abso ⁇ tion solvent 24 is introduced through a liquid inlet near the top of the SO 2 stripper zone 27 and distributed over the packing material.
  • the SO 2 -depleted abso ⁇ tion solvent 33 is then preferably withdrawn from an outlet near the bottom of the SO 2 stripper zone 27, and the SO 2 -enriched stripper gas 30 is removed from an outlet near the top of the SO 2 stripper zone 27.
  • the SO 2 -depleted abso ⁇ tion solvent 33 is recycled back to the solvent inlet near the top of the SO 2 abso ⁇ tion zone 18, thereby serving as the SO 2 abso ⁇ tion solvent 15 for further abso ⁇ tion of SO 2 from the source gas 3.
  • a conventional packed tower is typically prefe ⁇ ed, those skilled in the art will appreciate that the SO 2 stripper zone 27, like the SO 2 abso ⁇ tion zone 18, may have other suitable configurations, including structured packing or a tray tower.
  • the mass flow rate ratio (L/G) of the SO 2 -enriched abso ⁇ tion solvent 24 to the stripping gas 36 necessary to achieve substantial transfer of SO 2 from the SO 2 - enriched abso ⁇ tion solvent 24 to the stripper gas 36 may be determined by conventional design practice. Preferably, essentially all (i.e., at least about 90%, and more preferably at least about 95%) of the SO 2 contained in the SO 2 -enriched abso ⁇ tion solvent 24 is transfe ⁇ ed to the stripper gas 36.
  • the SO 2 -enriched stripper gas 30 exiting the top of the SO 2 stripper zone 27 is preferably passed to an overhead condenser, and a portion of any water vapor contained in the SO 2 -enriched stripper gas 30 is condensed by transfer of heat in the SO 2 -enriched stripper gas 30 to cooling water.
  • This condensate and the remainder of the SO 2 -enriched stripper gas 30 are then preferably transfe ⁇ ed to liquid/gas phase separator.
  • Solvent that may have been vaporized in the SO 2 stripper zone 27 may also be condensed in the overhead condenser and form part of the refluxed condensate. However, to avoid formation of two liquid phases in the separator, it is prefe ⁇ ed to operate the condenser such that the condensate refluxed to the stripper consists essentially of water vapor condensed from the SO 2 -enriched stripper gas 30.
  • the SO 2 -enriched abso ⁇ tion solvent 24 can be stripped by steam distillation (i.e., contacting the SO 2 -enriched abso ⁇ tion solvent 24 with live steam introduced into the bottom of the SO 2 stripper zone 27) to recover the SO 2 from the SO 2 -enriched abso ⁇ tion solvent 24.
  • the SO 2 preferably is stripped from the SO 2 -enriched abso ⁇ tion solvent 24 under non-reducing conditions.
  • the SO 2 stripper zone 27 preferably is operated at an average temperature of from about 80 to about 120°C, and more preferably from about 90 to about 110°C.
  • the prefe ⁇ ed operating pressure in the SO 2 stripper zone 27 is from about 20 to about 150 kPa (absolute).
  • Temperature control within the SO 2 abso ⁇ tion zone 18 and SO 2 stripper zone 27 may be achieved by controlling the temperature of the various process streams fed to these apparatus.
  • the temperature in the SO 2 stripper zone 27 is maintained within the desired range by controlling only the temperature of the SO 2 - enriched abso ⁇ tion solvent 24, while air is introduced at from about 20 to about
  • the SO 2 -enriched abso ⁇ tion solvent 24 exiting the SO 2 abso ⁇ tion zone 18 preferably is at a temperature of from about 10 to about 60 °C, more preferably from about 10 to about 50°C, and most preferably from about 30 to about 40°C.
  • This SO 2 -enriched abso ⁇ tion solvent 24 is preferably passed through a solvent heat interchanger 39 where it is preheated by indirect transfer of heat from the SO 2 -depleted solvent 33 being recycled from the SO 2 stripper zone 27 to the SO 2 abso ⁇ tion zone 18 (this, in turn, cools the SO 2 -depleted solvent 33 exiting the SO 2 stripper zone 27, which is typically at a temperature from about 80 to about 120°C). If further heating is required to achieve the desired temperature in the SO 2 stripper zone 27, the preheated SO 2 -enriched abso ⁇ tion solvent 24 leaving the interchanger 39 may be passed through a solvent heater, where it is further heated by indirect heat exchange with steam.
  • the SO 2 -depleted solvent 33 leaving the interchanger 39 may be passed through a solvent cooler where it is further cooled by indirect heat exchange with cooling tower water. It should be recognized that the use of a solvent interchanger 39 reduces the energy demands of the solvent heater, and reduces the cooling water required in the solvent cooler.
  • a purge stream may be periodically or continuously removed from the SO 2 -depleted solvent 33 and directed to a solvent purification vessel.
  • An aqueous wash stream such as water or a mildly alkaline aqueous solution (e.g., a sodium bicarbonate solution), is also introduced into the purification vessel and contacted with the purge stream.
  • the resulting two-phase mixture is then decanted to separate the aqueous phase containing the inorganic salt contaminants from the organic phase comprising SO 2 -depleted solvent 33 having a reduced contaminant concentration.
  • a waste stream comprising the aqueous waste is discharged from the purification vessel, while a liquid stream comprising the purified SO 2 abso ⁇ tion solvent is returned to the remaining SO 2 -depleted solvent 33 routed back to the SO 2 abso ⁇ tion zone 18.
  • the quantity of solvent 33 treated in this manner preferably is sufficient to maintain the contaminant concentration in the circulating solvent 33 at a level low enough to provide low process equipment co ⁇ osion rates and not materially compromise SO 2 abso ⁇ tion efficiency. It should be understood that the washing of the SO 2 -depleted solvent 33 may be carried out in a batch or a continuous fashion. If the SO 2 -depleted solvent 33 is washed continuously, a suitable liquid-liquid phase separator (e.g., a centrifugal contactor) may be used to separate the aqueous waste and purified organic phases.
  • a suitable liquid-liquid phase separator e.g., a centrifugal contactor
  • the SO 2 abso ⁇ tion/stripping zones are particularly useful when the source gas 3 has a relatively weak SO 2 concentration (i.e., from about 0.1 to about 11 mole%, and even more so at from about 0.1 to about 5 mole%) because they can be used to remove the inert gases (most notably, N 2 ) from the source gas 3 and thereby significantly increase the SO 2 concentration.
  • a relatively weak SO 2 concentration i.e., from about 0.1 to about 11 mole%, and even more so at from about 0.1 to about 5 mole%
  • One advantage of having a greater SO 2 concentration is that it allows for a smaller volume of gas to be handled during the process, thereby permitting the use of smaller equipment (which has cheaper capital and operational costs).
  • the oxygen concentration in the gas 30 can be increased without necessarily increasing the total volume of the SO 2 -containing gas.
  • This process also provides a mechanism for delaying the introduction of the oxygen needed for the SO 2 oxidation until the oxygen is actually needed (i.e., in the catalytic converter 45). This is particularly advantageous because, under such a scheme, only the amount of oxygen needed for producing the SO 2 has to be introduced into combustion zone 9.
  • the combustion zone 9 and other equipment upstream of the converter 45 does not have to be sized to handle the oxygen-containing gas which is required for the SO 2 oxidation. Because smaller equipment can be used upstream of the catalytic converter 45, significant capital and operational expenses can be avoided. Because the SO 2 abso ⁇ tion/stripping zones may be used to remove water from the source gas 3, they are particularly useful in embodiments where it is desirable to remove water vapor from the source gas 3 so that the SO 2 -containing gas fed to the catalytic converter 45 contains essentially no water vapor. Such embodiments include, for example, embodiments where the converter 45 and/or equipment downstream of the converter 45 are made of material which is vulnerable to co ⁇ osion caused by sulfuric acid formed by the vapor phase reaction of water vapor with SO 3 .
  • the SO 2 abso ⁇ tion/stripping zones are also particularly useful in embodiments where the H 2 O/SO 2 molar ratio in the source gas 3 is greater than the molar ratio of H 2 O/SO 3 in the desired acid product 51 (this situation may especially occur when the source gas 3 is prepared from spent acid, the off-gas of a metal roasting or smelting operation, or H 2 S).
  • the desired product acid concentration is 98.5 weight%
  • the H 2 O/SO 3 molar ratio in the conversion gas 54 fed to the SO 3 abso ⁇ tion zone 57 cannot exceed about 1.08.
  • the H 2 O/SO 2 molar ratio in the source gas 3 also preferably does not exceed about 1.08.
  • the SO 2 abso ⁇ tion/stripping zone may be used (alone or together with, for example, a drying tower and/or a cooling tower(s) which condenses liquid out of the source gas 3) to ensure that the H 2 O/SO 3 molar ratio is maintained below this value.
  • the SO 2 -enriched stripper gas 30 is preferably combined with a source of molecular oxygen 42 to form a converter feed gas 48, which is then passed through a catalytic converter 45 to oxidize the SO 2 to form a conversion gas 54 containing SO 3 .
  • the oxygen source 42 may be any oxygen-containing gas.
  • an "oxygen-containing gas” is a gas comprising molecular oxygen (O 2 ), which optionally may also comprise one or more diluents which are non-reactive with O 2 , SO 2 , SO 3 , and sulfuric acid under the reaction conditions. Examples of such gases are air, pure molecular oxygen, or molecular oxygen diluted with nitrogen and/or another inert gas(es).
  • the oxygen source 42 preferably is air or essentially pure molecular oxygen, with air being most prefe ⁇ ed. It should be recognized that the stripper gas 36 advantageously may provide part (or, in some instances, all) of the oxygen required in the converter feed gas 48 if the stripper gas 36 is air or another O 2 - containing gas.
  • the converter feed gas 48 contains essentially no water vapor, thereby reducing the risk of co ⁇ osion to process equipment downstream.
  • the SO 2 -enriched stripper gas 30 is wet, it preferably is dried, such as by being contacted with concentrated sulfuric acid in a drying tower before being introduced into the catalytic converter 45.
  • the SO 2 abso ⁇ tion solvent 15 is an organic phosphorous solvent as described above and dry air is used to strip the SO 2 from the SO 2 -enriched abso ⁇ tion solvent 24, the SO 2 -enriched stripper gas 30 often does not need to be dried before being routed to the converter 45.
  • the catalytic converter typically comprises at least two catalyst beds in series through which the converter feed gas 48 passes.
  • the catalyst in each of the catalyst beds may generally be any material which catalyzes the oxidation reaction of SO 2 to SO 3 .
  • Conventionally used catalysts include, for example, various vanadium compounds, platinum compounds (e.g., platinized asbestos), silver compounds, ferric oxide, chromium oxide, etc.
  • the catalyst comprises vanadium or a combination of vanadium and cesium.
  • the catalyst comprises vanadium pentoxide (V 2 O 5 ).
  • the SO 2 - enriched stripper gas 30 is normally at a temperature of no greater than about 120°C upon exiting the SO 2 stripper zone 27. And this temperature is typically decreased when the SO 2 -enriched stripper gas 30 is combined with the oxygen source 42, which is often near ambient temperature.
  • the more prefe ⁇ ed oxidation catalysts however, have an activation temperature which is significantly greater than 120°C.
  • the converter feed gas 48 is often preferably heated before being introduced into the first catalyst bed 60 of the converter 45.
  • the reaction is also preferably controlled so that the temperature of the catalyst bed 60 does not increase so much as to deactivate the catalyst and/or shift the reaction equilibrium to favor the reverse reaction.
  • a vanadium-containing catalyst e.g., V 2 O 5
  • V 2 O 5 vanadium-containing catalyst
  • it is typically prefe ⁇ ed for the converter feed gas 48 and partial conversion gas 69 and 72 entering catalyst beds 60, 63 and 66, respectively, to have a temperature of from about 410 to about 450 °C (even more preferably from about 415 to about 435 °C), and then to control the temperature in each bed so that the gas temperature approaches, but does not exceed, about 650 °C (more preferably about 630°C).
  • Temperature control in the converter 45 is preferably accomplished by maintaining the SO 2 strength (i.e., the SO 2 concentration) in the converter feed gas 48 and partial conversion gas 69 and 72 introduced into catalyst beds 60, 63 and 66, respectively, at no greater than about 15 mole%, more preferably no greater than about 13.5 mole%, and still more preferably no greater than about 12 mole%.
  • SO 2 strength i.e., the SO 2 concentration
  • the amount of the oxygen source 42 combined with the SO 2 -enriched stripper gas 30 be such that the molar ratio of O 2 to SO 2 in the gas converter feed gas 48 and partial conversion gas 69 and 72 introduced into catalyst beds 60, 63 and 66, respectively, be greater than about 0.2:1, more preferably at least about 0.5:1, even more preferably at least about 0.7:1, still even more preferably from about 0.7:1 to about 1.4:1, and most preferably from about 0.9:1 to about 1.2:1.
  • the SO 2 oxidation reaction is exothermic, it is often advantageous to use an indirect heat exchanger(s) 75 and 78 to heat the converter feed gas 48 with the partial conversion gas 81 and 84 exiting the catalyst beds 60 and 63 of the catalytic converter 45.
  • the converter feed gas 48 contains at least about 5 mole% SO 2 (and particularly at least about 8 mole%) and an excess amount of O 2 , the oxidation reaction can evolve sufficient heat for increasing the temperature of the converter feed gas 48 to the activation temperature of the oxidation catalyst, thus avoiding the need for any extraneous heat source for heating the converter feed gas 48 after startup (i.e., making the converter 45 energy self-sustaining or "autothermal").
  • the converter feed gas 48 preferably has an SO 2 concentration of from about 7 to about 15 mole%, more preferably from about 7 to about 13.5 mole%, even more preferably from about 7 to about 12 mole%, still even more preferably from about 10 to about 12 mole%, and most preferably about 11.5 mole%.
  • the converter feed gas 48 preferably is preheated using two indirect heat exchangers in series: first, a cold heat exchanger 78 in which the converter feed gas 48 is preheated by transfer of heat from the partial conversion gas 84 leaving the second bed 63 of the converter 45; and, second, a hot heat exchanger 75 in which the converter feed gas 48 is further heated by transfer of heat from the partial conversion gas 81 leaving the first catalyst bed 60 of the converter 45.
  • the converter feed gas 48 may also (or alternatively), for example, be heated by passing it through an indirect heat exchanger to transfer heat from the source gas 3 to the converter feed gas 48.
  • the SO 2 -enriched stripper gas 30 is split into at least two streams.
  • a portion, preferably at least about 30% (more preferably at least about 40%, and even more preferably at least about 50%) of the SO 2 -enriched stripper gas 30 is combined with the oxygen source 42 (either before or after being preheated, and preferably before) to form the converter feed gas 48, which, in turn, is introduced into the first catalyst bed 60 of the converter 45 wherein a portion of the SO 2 content of the gas 48 is oxidized to SO 3 to form a partial conversion gas 81 containing SO 3 and residual SO 2 .
  • the cooled partial conversion gas exiting indirect heat exchanger 75 is then combined with a second portion 31 of the SO 2 -enriched stripper gas 30 to fortify the SO 2 concentration in the partial conversion gas.
  • the fortified partial conversion gas 69 is then passed through at least one additional catalyst bed (63 and 66 in Fig. 1) to oxidize further SO 2 in the gas 69.
  • Fortifying the SO 2 gas strength of the partial conversion gas is advantageous because it significantly increases the capacity of the converter 45.
  • the maximum SO 2 concentration of the gas fed into the first catalyst bed 60 is normally limited (in the presence of excess oxygen) to about 15 mole% (more typically about 13.5 mole%, and even more typically about 12 mole%) because greater SO 2 concentrations will typically cause too much heat to be released during the oxidation reaction, thereby causing the catalyst to deactivate and/or the reaction equilibrium to shift unfavorably.
  • the amount of SO 2 added to the partial conversion gas increases the SO 2 concentration to no greater about 15 mole%, more preferably from about 7 to about 13.5 mole%, even more preferably from about 7 to about 12 mole%, still even more preferably from about 10 to about 12 mole%, and most preferably about 11.5 mole%.
  • the SO 2 -enriched stripper gas 30 is split into more than 2 portions and subsequently used to fortify the feed gas to more than one catalyst bed of the converter.
  • the SO 2 -enriched stripper gas may, for example, be split into three portions.
  • the first portion of the SO 2 -enriched stripper gas is combined with the oxygen source to form the converter feed gas, which, in turn, is introduced into the first catalyst bed of the converter where SO 2 in the gas is oxidized to form a partial conversion gas.
  • This partial conversion gas is then combined with the second portion of the SO 2 -enriched stripper gas to fortify the SO 2 strength in the partial conversion gas.
  • the fortified partial conversion gas is then passed through the second catalyst bed to oxidize further SO 2 and form a second partial conversion gas.
  • This second partial conversion gas is then combined with the third portion of the SO 2 - enriched stripper gas to fortify the SO 2 strength in the second partial conversion gas.
  • This fortified second partial conversion gas is then passed through the third catalyst bed to oxidize still further SO 2 and form a third partial conversion gas.
  • This third partial conversion gas is then passed through the fourth (i.e., the final) catalyst bed to oxidize at least a portion of any remaining SO 2 .
  • the conversion gas 54 exiting the catalytic converter 45 preferably is contacted with water or, more preferably, concentrated sulfuric acid 87 (preferably an aqueous solution containing from about 96 to about 99.5 weight% H 2 SO 4 , more preferably from about 98.5 to about 99.5 weight%, and most preferably from about 99 to about 99.5 weight%) in an SO 3 abso ⁇ tion zone 57 to absorb SO 3 from the conversion gas 54, thereby forming an SO 3 -depleted gas 90 and additional sulfuric acid and/or oleum 51.
  • This heat recovery zone 93 preferably recovers energy from the heat of abso ⁇ tion of the SO 3 in the SO 3 abso ⁇ tion zone 57.
  • Sulfur trioxide abso ⁇ tion zones and heat recovery zones suitable for use in accordance with this invention are well-known in the art. See, e.g., McAlister et al., U.S. Patent Nos. 4,670,242 and 4,576,813 (both inco ⁇ orated herein by reference).
  • the conversion gas 54 is cooled in an economizer to a temperature which is above the dew point of the conversion gas 54, and then introduced into the lower portion of a vertical tower comprising the SO 3 abso ⁇ tion zone 57.
  • the SO 3 abso ⁇ tion zone 57 preferably comprises a bed of random packing (although the SO 3 abso ⁇ tion zone 57 may alternatively comprise another gas-liquid contacting device, such as a tray tower).
  • the cooled conversion gas 54 flows upward through the SO 3 abso ⁇ tion zone 57.
  • This concentrated sulfuric acid 87 preferably has a temperature of greater than about 120°C. Such conditions tend to reduce sulfuric acid co ⁇ osiveness to alloys used in many conventional abso ⁇ tion towers, while providing a high degree of SO 3 abso ⁇ tion.
  • the sulfuric acid concentration in the sulfuric acid solution 96 is preferably greater than about 98 weight% (more preferably greater than about 98.5 weight%, even more preferably greater than about 99 weight%, and most preferably from about 99 to about 100 weight%). It should be recognized that these prefe ⁇ ed concentrations can be greater if the SO 3 abso ⁇ tion zone 57 is operated at pressure significantly greater than atmospheric pressure.
  • the temperature of the liquid sulfuric acid increases as the liquid sulfuric acid becomes more concentrated while passing through the abso ⁇ tion zone 57.
  • the temperature of the concentrated sulfuric acid preferably increases to a temperature of up to about 250°C (this prefe ⁇ ed maximum temperature is greater at absorber pressures greater than atmospheric pressure).
  • the liquid sulfuric acid 96 preferably is passed through a heat recovery zone 93 (which may either be physically inside or outside of the abso ⁇ tion zone 57, and most preferably comprises an indirect heat exchanger outside the abso ⁇ tion zone 57) to remove the heat of abso ⁇ tion of the SO 3 .
  • This heat may, in turn, be used, for example, to generate low to medium pressure steam (typically up to about 10.5 bar (gauge)) for use within the manufacturing complex su ⁇ ounding the sulfuric acid plant or to generate electricity.
  • the liquid sulfuric acid concentration preferably is at least about 99 weight% throughout the course of the heat transfer. It is also prefe ⁇ ed that the temperature of the liquid sulfuric acid 96 throughout the heat exchanger be greater than about 130°C (more preferably greater than about 140°C, and most preferably greater than about 150°C) where low pressure steam is desired (i.e., up to about 3.5 bar (gauge)), and be greater than about 150°C (more preferably greater than about 175°C, and most preferably greater than about 200°C) where medium pressure steam is desired (i.e., from about 6.5 to about 10.5 bar (gauge)).
  • a portion of the sulfuric acid stream 96 preferably is recovered as product 51.
  • the remainder 99 preferably is diluted with water 102 (in either liquid or vapor form) or dilute sulfuric acid, and reticulated to the top of the SO 3 abso ⁇ tion zone 57 to again be passed through the SO 3 abso ⁇ tion zone 57.
  • the gas 90 may optionally be passed through a second SO 3 abso ⁇ tion zone which may be a second stage of the tower containing the first SO 3 abso ⁇ tion zone 57, or may be located in a separate tower.
  • the pu ⁇ ose of such a second stage or tower is to remove any residual SO 3 that remains in the SO 3 -depleted gas 90. It should be recognized, however, that in many instances, essentially all the SO 3 is absorbed in the primary SO 3 abso ⁇ tion zone 57, rendering a second stage or a second tower unnecessary.
  • the process of this invention may comprise more than one SO 3 abso ⁇ tion zone such that partial conversion exiting an intermediate catalyst bed of the converter is contacted with water or a liquid comprising sulfuric acid to absorb SO 3 from the gas before the gas is passed through one or more subsequent catalyst beds of the converter (i.e., the process may be used with a system comprising an intermediate SO 3 absorber).
  • the partial conversion gas leaving the second or third bed may be passed through an intermediate SO 3 abso ⁇ tion zone (i.e., an inte ⁇ ass abso ⁇ tion zone) for removal of SO 3 in the form of product acid and/or oleum.
  • Gas exiting the intermediate abso ⁇ tion zone is then returned to the next downstream catalyst bed of the converter. Because the conversion of SO 2 to SO 3 is an equilibrium reaction, removal of SO 3 in the inte ⁇ ass abso ⁇ tion zone helps drive the reaction forward in the succeeding bed or beds of the converter to achieve higher conversions.
  • Use of an intermediate SO 3 abso ⁇ tion zone is normally less prefe ⁇ ed in the practice of the present invention because it substantially adds to the capital and operating costs.
  • At least a portion of the SO 3 -depleted gas 90 exiting the SO 3 abso ⁇ tion zone 57 is recycled back to the SO 2 abso ⁇ tion zone 18 and contacted with the SO 2 abso ⁇ tion solvent 15 along with the source gas 3.
  • unconverted SO 2 in the tail gas 90 is thereby recaptured in the SO 3 -enriched abso ⁇ tion solvent 24 exiting the SO 2 abso ⁇ tion zone 18, stripped from the SO 3 -enriched abso ⁇ tion solvent 24 in the SO 2 stripper zone 27, and returned to the catalytic converter 45 as part of the SO 2 -enriched stripper gas 30 for ultimate recovery as product acid 51.
  • at least a substantial portion of the inert gases and excess O 2 in the recycled tail gas 90 will be purged from the process in the SO 2 -depleted gas 21 exiting the SO 2 abso ⁇ tion zone 18.
  • emission standards may be met by recycling less than all of the tail gas 90 from the SO 3 abso ⁇ tion zone 57.
  • target emissions may be met by recycling 90%, 75%, or even 50% of the tail gas 90, with some resultant savings in energy costs for gas compression. It is ordinarily prefe ⁇ ed, however, that substantially all the tail gas 90 be recycled.
  • the SO 2 conversion per single pass through the entire converter 45 i.e., the total amount of SO 2 consumed during a single pass through the entire converter ⁇ total amount of SO 2 fed into the converter x 100%
  • the SO 2 conversion per single pass through the entire converter 45 be at least 75%, more preferably at least about 85%, even more preferably at least about 90%, and most preferably at least about 95%.
  • the process of this invention may be implemented using only two (or, more preferably, three (as shown in Fig. 1)) catalyst beds in the catalytic converter 45.
  • this process may also be implemented using a double SO 3 abso ⁇ tion plant and/or 4 or more catalyst beds in the catalytic converter.
  • an existing contact acid plant having, for example, a 4-catalyst-bed converter
  • an already-existing contact sulfuric acid production plant including a catalytic converter with 4 catalyst beds in series and at least two associated indirect heat enchanters for cooling the partial conversion gas passing between catalyst beds is modified (i.e., retrofitted) so that the converter comprises 2 parallel sets of 2 catalyst beds in series.
  • the flow scheme for such a retrofitted catalytic converter is schematically illustrated in Fig. 2.
  • the parallel sets of catalyst beds are typically contained within the single vessel which housed the serial catalyst beds of the original converter. However, it should be understood that the parallel sets of catalyst beds could be housed in separate vessels.
  • the SO 2 -enriched stripper gas 30 is preferably ultimately divided into 4 portions.
  • a first portion of the SO 2 - enriched stripper gas 30 is combined with an oxygen source 42 to form a converter feed gas 48, which is subsequently divided to form a first converter feed gas 48A and a second converter feed gas 48B.
  • the first converter feed gas 48A is heated in indirect heat exchanger 75 and passed through the first catalyst bed 60 of the first set of catalyst beds to form a first partial conversion gas 81 A
  • the second converter feed gas 48B is simultaneously heated in indirect heat exchanger 78 and passed through the first catalyst bed 65 of the second set of catalyst beds to form a second partial conversion gas 84A.
  • the remainder of the SO 2 -enriched stripper gas 31 is divided and a first portion 31A is combined with the cooled first partial conversion gas exiting indirect heat exchanger 75 to fortify the SO 2 concentration in the first partial conversion gas and produce a fortified first partial conversion gas 69 A.
  • the second portion 31B of the remainder of the SO2-enriched stripper gas 31 is likewise combined with the cooled second partial conversion gas exiting indirect heat exchanger 78 to fortify the SO 2 concentration in the second partial conversion gas and produce a fortified second partial conversion gas 72A.
  • the first fortified partial conversion gas 69 A is passed through the second catalyst bed 63 of the first set of catalyst beds to form a first conversion gas 54A
  • the second fortified partial conversion gas 72 A is passed through the second catalyst bed 66 of the second set of catalyst beds to form a second conversion gas 54B.
  • the first and second conversion gases 54 A and 54B may then be combined to form conversion gas 54 and introduced into a single SO 3 abso ⁇ tion zone.
  • the first conversion gas 54A preferably is introduced into one of the SO 3 abso ⁇ tion zones, while the second conversion gas 54B is introduced into the other SO 3 abso ⁇ tion zone (i.e., the two SO 3 abso ⁇ tion zones are operated in parallel). In either case, it is particularly prefe ⁇ ed to recycle the SO 3 - depleted tail gas exiting the SO 3 abso ⁇ tion zone (or zones) to the SO 2 abso ⁇ tion zone.
  • Water vapor may be introduced into the conversion gas 54 exiting the catalytic converter 45.
  • the water vapor reacts with the SO 3 in the conversion gas 54 to produce gaseous sulfuric acid.
  • a portion of the energy from the heat of formation of the gaseous sulfuric acid may, in turn, be recovered by, for example, passing the resulting gas through a heat exchanger.
  • Substantial additional energy may be recovered by also (or alternatively) passing the gas through a condensing economizer.
  • the source of the water vapor may, for example, be low pressure steam (i.e., up to about 6.5 bar (gauge), more preferably up to about 3.5 bar (gauge), and most preferably from about 0.2 to about 1 bar (gauge)).
  • This low pressure steam may be obtained from a variety of sources at a sulfuric plant, such as, for example, a low pressure port on a steam turbine for an electrical generator, steam generated from low temperature sulfuric acid, etc.
  • a wet SO 2 source gas 3 is used, and at least a portion (often preferably all) of the source gas 3 is combined with the conversion gas 54 to supply at least a portion (preferably all) of the water vapor.
  • An example of such an embodiment is illustrated in Fig. 3.
  • the water vapor in the wet source gas 1003 reacts with the SO 3 in the conversion gas 1006 to produce gaseous sulfuric acid.
  • the vapor phase formation of gaseous sulfuric acid generates heat which preferably is recovered as energy by, for example, transferring the heat to steam or feed water in an indirect heat exchanger 1012.
  • more energy is preferably recovered by condensing at least a portion of the gaseous sulfuric acid into liquid sulfuric acid in a condensing economizer 1015.
  • the gas 1018 exiting the condensing economizer 1015 is then preferably passed through an SO 3 abso ⁇ tion zone 1021 (which preferably is associated with a heat recovery means 1024 which recovers the energy from the heat of abso ⁇ tion produced in the SO 3 abso ⁇ tion zone 1021) where SO 3 , water vapor, and any additional gaseous sulfuric acid is separated from the gas 1018 to form a dry SO 3 -depleted gas 1066.
  • This dry SO 3 -depleted gas 1066 is a/the source of SO 2 for the converter feed gas 1030.
  • the condensing economizer 1015 preferably comprises an indirect heat exchanger in which heat is transfe ⁇ ed to a heat transfer fluid (e.g., boiler feed water).
  • a heat transfer fluid e.g., boiler feed water
  • This indirect heat exchanger preferably comprises heat transfer wall means (e.g., the tubes of a shell and tube type heat exchanger), preferably constructed of an alloy (e.g., an Fe/Cr or Fe/Cr/Ni alloy) which is resistant to co ⁇ osion by condensing sulfuric acid.
  • heat transfer wall means e.g., the tubes of a shell and tube type heat exchanger
  • an alloy e.g., an Fe/Cr or Fe/Cr/Ni alloy
  • at least a portion of the wall means on the gas stream side of the exchanger is at a temperature which is less than the dew point of the gas stream in the exchanger.
  • the condensing economizer 1015 may be operated to condense as sulfuric acid as much as from about 5 to about 20% of the SO 3 generated in the catalytic converter 1009.
  • Table 1 shows the heat evolved when SO 3 and water react to form sulfuric acid under various phase conditions.
  • Table 1 Sulfuric Acid Heat of Reaction from Standard Heat of Formation (25°C) No. Reaction Conditions Heat of Reaction
  • the gas phase reaction (Equation 2) produces 74% of the heat produced by the normal liquid phase reaction (Equation 1). Transferring the heat from condensing sulfuric acid to boiler feed water results in the ultimate recovery of both the heat of formation and heat of condensation of the sulfuric acid.
  • the boiler feed water may be further heated with the source gas 1003 as the source gas 1003 exits the SO 2 - producing combustion zone 1002 to form high grade energy, i.e., steam at a pressure of at least about 30 bar (gauge), and more preferably from about 40 to about 60 bar (gauge). This steam may be further heated by, for example, the conversion- gas/source-gas mixture 1039 in the indirect heat exchanger 1012.
  • the conversion of SO 3 to sulfuric acid in the vapor phase increases as the temperature of the vapor phase decreases.
  • the condensing economizer 1015 can be operated to extract a maximum amount of the vapor phase energy of formation of sulfuric acid without the necessity for close control of the fluid flow rates or wall temperatures within the economizer 1015.
  • the concentration of acid in the condensate 1033 varies only slightly with the H 2 O/SO 3 molar ratio in the gas 1036, and consequently does not vary significantly with either the temperature to which the gas 1036 is cooled or the wall temperature of the condensing economizer 1015. Thus, it is not necessary to closely control the operation of the condensing economizer 1015 to avoid co ⁇ osive conditions therein. And, variations in inlet air humidity, or excursions in sulfur flow rate, do not materially affect the concentration of the acid condensate 1033 on the tube walls of the condensing economizer 1015. As much as 140% of the stoichiometric amount of water vapor may be present in the gas 1036 without reducing the concentration of the condensing acid condensate 1033 to less than 98%.
  • the energy equivalent of from about 40 to about 70% (most typically about 60%) of the heat of formation of sulfuric acid vapor may be recovered by cooling the gas 1039 before it enters the SO 3 abso ⁇ tion zone 1021.
  • an initial heat exchanger 1012 and a condensing economizer 1015 typically from about 70% to about 90% (and more typically about 75%) of the recovered heat of formation is recovered in the condensing economizer 1015.
  • the boiler feed water enters the condensing economizer at a temperature of from about 110 to about 180°C
  • the gas 1036 enters the condensing economizer 1015 at a temperature of from about 320 to about 470°C, and with an H 2 O/SO 3 mole ratio of from about 0.2 to about 1.05.
  • the gas 1018 leaving the condensing economizer 1015 preferably has a temperature of from about 240 to about 300°C.
  • a substantial portion of the vapor phase heat of formation of sulfuric acid can be extracted without condensation in the economizer 1015.
  • recovery of a substantial fraction of the heat of formation may be achieved without condensation by transferring heat from the gas 1036 to boiler feed water in a co-cu ⁇ ent heat exchanger.
  • the wet gas 1018 leaving the condensing economizer 1015 preferably is directed to an SO 3 abso ⁇ tion zone 1021 where it is contacted countercu ⁇ ently with a concentrated solution of sulfuric acid 1048.
  • the SO 3 abso ⁇ tion zone 1021 comprises a means in a vertical tower for promoting mass transfer and heat transfer between the gas and liquid phases within the tower (preferably a bed of random packings such as saddles or rings, although it should be understood that other gas liquid contacting devices, e.g., a countercu ⁇ ent tray tower or a co-cu ⁇ ent venturi absorber, may be used in lieu of random packing).
  • the inlet gas 1018 to the abso ⁇ tion zone 1021 comprises SO 3 and sulfuric acid vapor. Contact of the gas 1018 with the liquid sulfuric acid 1048 causes abso ⁇ tion of SO 3 , condensation and abso ⁇ tion of any water vapor, and condensation and abso ⁇ tion of sulfuric acid vapor into the sulfuric acid solution.
  • heat of abso ⁇ tion and “energy of abso ⁇ tion” include all these various heat effects, and may also include energy of formation of sulfuric acid in the vapor phase that has not been recovered in condensing economizer 1015.
  • the use of hot acid for SO 3 abso ⁇ tion provides at least two advantages.
  • the heat of abso ⁇ tion is generated at relatively high temperature which allows subsequent recovery of this energy at high temperature.
  • the use of high temperature acid avoids shock cooling of the gas 1018 and consequently minimizes the formation of acid mist in the wet gas.
  • the temperature of the acid 1051 at the exit of the abso ⁇ tion zone 1021 is no greater than about 40°C less than (and more preferably no greater than about 20°C less than) the dew point of the inlet gas 1018.
  • the gas 1018 can typically be at a temperature of up to about 300°C as it enters the SO 3 abso ⁇ tion zone 1021, thereby allowing recovery of the maximum amount of the energy of vapor phase formation and condensation of sulfuric acid in the form of high pressure steam as a result of the transfer of this heat to the high pressure boiler feed water for waste heat boiler.
  • the concentrated sulfuric acid contact solution 1048 is introduced at an inlet near the top of the SO 3 abso ⁇ tion zone 1021, while the gas 1018 is introduced at an inlet near the lower end of the SO 3 abso ⁇ tion zone 1021.
  • the acid solution 1048 at the acid inlet preferably has a temperature of from about 170 to about 220°C, and a sulfuric acid concentration of from about 98.5 to about 99.5%, and more preferably from about 99 to about 99.5%.
  • the gas 1018 at the gas inlet preferably has a temperature of from about 240 to about 300°C, and an H 2 O/SO 3 molar ratio which preferably is less than the H 2 O/SO 3 molar ratio in the acid solution 1048 , and equals from about 0.2 to about 1.05 (more preferably from about 0.7 to about 1.0).
  • the H 2 O/SO 3 molar ratio in the gas 1018 entering the SO 3 abso ⁇ tion zone 1021 preferably is reduced by either partially drying the source gas
  • the acid solution 1051 preferably is discharged from the SO 3 abso ⁇ tion zone 1021 at a temperature of at least about 190°C, more preferably from about 190 to about 250°C, and even more preferably from about 210 to about 250°C. At least a major portion of this solution preferably flows to a circulating pump, and passed through an indirect heat exchanger 1024 where the energy of abso ⁇ tion is recovered by transfer of heat to another fluid.
  • the indirect heat exchanger 1024 comprises a heat recovery system boiler, and the heat energy is ultimately recovered in the form of low to medium pressure (i.e., up to about 10.5 bar (gauge)).
  • the acid solution 1054 from the indirect heat exchanger 1024 is preferably recirculated back to the SO 3 abso ⁇ tion zone 1021.
  • a portion 1057 of the acid solution 1054 preferably is removed as product before the acid solution 1054 is recirculated (additional heat energy may be recovered from this acid product by, for example, passing it through one or more additional indirect heat exchangers).
  • An equal amount of water 1060 is then added to the remaining sulfuric acid solution 1063. This water 1060 may, for example, be added in liquid or vapor form, or in the form of diluted sulfuric acid.
  • the SO 3 -depleted gas 1066 exiting the top of this zone 1021 is relatively hot. This, in turn, often results in the stripping of sulfuric acid from the acid stream into the gas stream.
  • the abso ⁇ tion efficiency of the SO 3 abso ⁇ tion zone 1021 is at least about 90%
  • high temperature operation of the abso ⁇ tion zone 1021 also typically results in some unabsorbed SO 3 passing through the abso ⁇ tion zone 1021.
  • Gas 1066 exiting the top of the SO 3 abso ⁇ tion zone 1021 is therefore preferably directed to a condensing stage for abso ⁇ tion of residual SO 3 and condensation of sulfuric acid vapor.
  • This condensing stage preferably contains means for promoting gas/liquid contact and mass transfer and heat transfer.
  • this stage comprises a countercu ⁇ ent packed section wherein relatively cool acid having a concentration of about 98.5% is fed to the top of this stage and gas 1066 leaving the main SO 3 abso ⁇ tion zone 1021 (which is typically at a temperature of from about 170 to about 230°C) enters the bottom of the condensing stage.
  • the acid entering the condensing stage preferably is at a temperature of less than about 120°C, most preferably from about 60 to about 80°C.
  • the gas 1066 On passage through the condensing stage, the gas 1066 preferably is cooled to a temperature of from about 75 to about 140°C, and more preferably from about 80 to about 120°C. Gas leaving the condensing stage is then preferably passed through a mist eliminator.
  • the acid flow rate in the condensing stage preferably is maintained at a rate low enough so that the acid leaves the stage at a temperature which approaches the temperature of the acid entering the main SO 3 abso ⁇ tion packed bed.
  • the gas 1066 exiting the SO 3 abso ⁇ tion zone 1021 i.e., the SO 3 -depleted gas
  • any portion 1045 of the source gas 1003 that is not combined with the conversion gas 1006 is preferably used to form the converter feed gas 1030.
  • the SO 3 -depleted gas 1066 (along with any portion 1045 of the source gas 1003 which is not combined with the conversion gas 1006) is first passed through the SO 2 abso ⁇ tion/stripper zones described previously. This removes the excess inert gases, and can be used to enhance the SO 2 concentration in the SO 3 -depleted gas 1066 where the SO 2 concentration in the SO 3 -depleted gas 1066 is less than the desired concentration.
  • the gas exiting the SO 2 abso ⁇ tion/stripper zones (i.e., the SO 2 -enriched stripper gas 1069) is then preferably combined with a dry oxygen source 1072 if the stripper gas 1075 does not supply the desired level of oxygen.
  • the SO 2 -enriched stripper gas 1069 may also be divided into 2 or more portions in the same manner as described above wherein one portion of the SO 2 -enriched stripper gas 1069 is combined with the dry oxygen source 1072 and fed into the first catalyst bed 1078 of converter 1009, and a second portion 1070 is used to fortify the SO 2 concentration of the partial conversion gas 1081 exiting the first catalyst bed 1078.
  • any oxygen source 1072 combined with the SO 2 -enriched stripper gas 1069 preferably is dried beforehand.
  • the SO 2 abso ⁇ tion solvent 1084 consist essentially of a composition that transfers little or no water to the SO 2 -enriched stripper gas 1069.
  • the organic phosphorus solvents discussed above are generally suitable for this pu ⁇ ose, particularly where the stripper gas 1075 is dry air.
  • the wet-source-gas embodiment described above is advantageous because it produces a dry SO 2 gas 1030 for the converter 1009 without having to first pass the entire source gas 1003 through a drying tower, thereby avoiding the capital and operational expenses associated with such a tower (and associated equipment, e.g., a pump, piping, a pump tank, and a cooler) as to the portion 1042 of the source gas 1003 that is combined with the conversion gas 1006 (as noted above, it is most often prefe ⁇ ed that this portion 1042 be the entire source gas 1003).
  • this process is advantageous because it produces heat (i.e., the heat of formation of gaseous sulfuric acid, the heat of condensation of gaseous sulfuric acid, and the heat of condensation of water vapor) which may be transfe ⁇ ed and used elsewhere as energy.
  • heat i.e., the heat of formation of gaseous sulfuric acid, the heat of condensation of gaseous sulfuric acid, and the heat of condensation of water vapor
  • wet SO 3 -containing gas can be handled in carbon steel equipment, although the gas temperature in such equipment preferably is kept above the dew point to avoid the condensation of gaseous sulfuric acid formed from the water vapor and SO 3 .
  • the dew point is generally high and much of the equipment (particularly the condensing economizer) is operated at a temperature below the dew point.
  • This equipment therefore, preferably is made of a material that is resistant to sulfuric acid co ⁇ osion under the conditions of this invention.
  • Alloy performance may be characterized by a co ⁇ osion index (Cl), which is defined in terms of alloy composition by the following relationship:
  • [Cr] is the weight percent of chromium in the alloy
  • [Ni] is the weight percent of nickel in the alloy
  • [Mo] is the weight percent of molybdenum in the alloy. Alloys which work best in high temperature strong sulfuric acid service have been found to have a co ⁇ osion index of greater than 7, and particularly greater than 8.
  • the alloys most likely to exhibit low co ⁇ osion rates are those with the highest co ⁇ osion index.
  • high chromium is desirable, and it is preferable to avoid alloys which have both high nickel and high molybdenum. It should be recognized, however, that alloys which contain high nickel and very low molybdenum, or low nickel and moderate amounts of molybdenum, are often acceptable.
  • Particular alloys found suitable for use in contact with liquid phase sulfuric acid at high temperature include, for example, those having UNS designations S30403, S30908, S31008, S44627, S32304, and S44800.
  • a source gas 2003 containing about 19 mole% SO 2 , about 2 mole % O 2 , and about 79 mole% N 2 is formed in a sulfur burner 2006 by burning sulfur 2009 in the presence of dry air 2012.
  • This source gas 2003 (initially at a temperature of about 1538°C upon exiting the sulfur burner 2006) is cooled to about 548°C in a waste heat boiler 2002.
  • the source gas 2004 is further cooled to about 337°C in an indirect heat exchanger 2015 (i.e., MonplexTM, Monsanto Environ-Chem Systems, Inc., St. Louis, MO, USA) by transferring heat from the source gas 2004 to the gas 2018 being fed into the SO 2 oxidation catalytic converter 2021.
  • the source gas 2024 is cooled even further to about 204°C in yet another indirect heat exchanger 2023 which uses heat in the source gas 2024 to form steam.
  • the cooled source gas 2025 is split into two portions: one portion 2026 (being about 6.6 volume% of the cooled source gas 2025) is fed back into the sulfur burner 2006 to maintain the desired temperature in the burner 2006, and the remaining portion 2028 (being about 94% of the cooled source gas 2025) is introduced into the SO 2 abso ⁇ tion/stripping zones (i.e., a Claus MasterTM, Monsanto Environ-Chem Systems, Inc., St. Louis, MO, USA).
  • the source gas 2028 is passed through a packed SO 2 abso ⁇ tion column 2027, where it is contacted with a liquid SO 2 abso ⁇ tion solvent comprising dibutyl butyl phosphonate 2030 flowing countercu ⁇ ently to the source gas 2028.
  • the dibutyl butyl phosphonate 2030 selectively absorbs SO 2 to form an SO 2 -enriched abso ⁇ tion solvent 2033 and an SO 2 - depleted gas 2036 (the SO 2 -depleted gas 2036 containing substantially all the residual O 2 and inert gases (mostly N 2 ) from the source gas 2028).
  • the SO 2 -depleted gas 2036 is discharged from the system, and the SO 2 -enriched abso ⁇ tion solvent 2033 is introduced into a packed SO 2 stripper column 2039, where the SO 2 -enriched abso ⁇ tion solvent 2033 is contacted with a countercu ⁇ ent flow of dry air 2042 (the dry stripper air 2042 entering the column 2039 has a temperature of about 110°C) to form an SO 2 -enriched stripping gas 2045 (containing about 90 mole% SO 2 , with the remaining being air) and an SO 2 -depleted abso ⁇ tion solvent 2048 (which is recycled back to the SO 2 abso ⁇ tion column 2027 to be used again as the SO 2 abso ⁇ tion solvent 2030).
  • Both the SO 2 abso ⁇ tion column 2027 and the SO 2 stripper column 2039 are operated at nearly atmospheric pressure.
  • the SO 2 -enriched stripping gas 2045 is divided into two portions: one portion (i.e., the primary SO 2 gas 2051) being about 54 volume% of the SO 2 -enriched stripping gas 2045, and the other portion (i.e., the bypass SO 2 gas 2054) being about 46 volume% SO 2 -enriched stripping gas 2045 (both portions having the same composition, i.e., 90 mole% SO 2 , with the remaining being air).
  • the primary SO 2 gas 2051 is combined with dry air 2057 (the dry air 2057 having a temperature of about 66°C) to form a converter feed gas 2018 containing about 12 mole% SO 2 and having a temperature of about 130°C. This converter feed gas 2018 is heated to a temperature of about 410°C by the gas 2004 coming from the sulfur burner 2006 using the
  • the converter feed gas 2060 is passed through a first catalyst bed 2063 containing V 2 O 5 which converts (i.e., oxidizes) about 67% of the SO 2 in the converter feed gas 2060 into SO 3 , thereby forming a partial conversion gas 2066 having a temperature of about 637°C, and containing about 4 mole% SO 2 and about 8 mole% SO 3 .
  • the V 2 O 5 catalyst in the first catalyst bed 2063 is a potassium-promoted catalyst coated on a silica support, and is in the shape of rings having an outer diameter of 12.5 mm, an inner diameter of 5 mm, and an average length of 14 mm (Cat. No.
  • the first catalyst bed 2063 has a diameter of about 26.25 feet and contains about 50,000 liters of the catalyst.
  • the total flowrate of the converter feed gas 2060 into the first catalyst bed 2063 is about 50,767 scfin (i.e., standard cubic feet per minute (defined at 70°F and 1 arm)).
  • the partial conversion gas 2066 exiting the first catalyst bed 2063 is cooled to about 420°C by transferring heat to feed water in an indirect heat exchanger 2069.
  • the cooled partial conversion gas 2072 is then combined with the bypass SO 2 gas
  • the SO 2 -fortified partial conversion gas 2075 (having a temperature of about 423°C) is then passed through a second catalyst bed 2078 containing V 2 O 5 to oxidize more SO 2 to form a second partial conversion gas 2081 having a temperature of about 607°C, and containing about 15.4 mole% SO 3 and about 6.1 mole% un-oxidized SO 2 .
  • the second catalyst bed 2078 has the same dimensions, the same V 2 O 5 catalyst, and the same volume of catalyst as the first catalyst bed 2063.
  • the total flowrate of gas entering the second catalyst bed 2078 is about 54,366 scfin.
  • the second partial conversion gas 2081 is cooled to about 420°C by transferring heat to feed water in a second indirect heat exchanger 2084, and then passed through a third catalyst bed 2087 containing V 2 O 5 to oxidize still more SO 2 and form a final conversion gas 2090 having a temperature of about 519°C, and containing about 20.0 mole% SO 3 and about 2.1 mole% un-oxidized residual SO 2 .
  • the V 2 O 5 catalyst in the third catalyst bed 2087 is a potassium-promoted catalyst coated on a silica support and is in the shape of rings having an outer diameter of 9.5 mm, an inner diameter of 4 mm, and an average length of 13 mm (Cat. No.
  • the third catalyst bed 2087 has a diameter of about 26.25 feet and contains about 80,000 liters of the catalyst.
  • the total flowrate of gas entering the third catalyst bed 2087 is about 52,406 scfin.
  • the final conversion gas 2090 is cooled to a temperature of about 166°C in an indirect heat exchanger 2091, and then contacted in a packed SO 3 abso ⁇ tion column 2093 (having a diameter of about 12 feet and a height of about 40 feet) with a countercu ⁇ ent flow of an aqueous solution 2096 containing about 98.5 weight% H 2 SO 4 to form a more concentrated sulfuric acid solution 2097 having a sulfuric acid concentration of about 99.5 weight%.
  • the flowrate of the conversion gas 2090 is about 51,349 scfin, while the flowrate of the aqueous sulfuric acid solution 2096 is about 1,700 gallons per minute.
  • the temperature of the aqueous sulfuric acid solution 2096 entering the column 2093 is about 82°C, and the temperature of the sulfuric acid solution 2097 exiting the SO 3 abso ⁇ tion column 2093 is about 110°C.
  • the gas 2102 exiting the SO 3 abso ⁇ tion column 2093 (i.e., "the SO 3 -depleted gas” or “tail gas") is split into 2 portions: one portion 2103 (being about 80 volume% of the SO 3 -depleted gas 2102) is combined with the source gas stream 2028, and thereby routed to the SO 2 abso ⁇ tion column 2027.
  • the other portion 2104 (being about 20 volume% of the SO 3 -depleted gas 2102) is combined with the dry air 2057 being combined with the primary SO 2 gas 2051, thereby maintaining a smaller volume of total gas being fed into the SO 2 abso ⁇ tion column 2027.
  • both portions of the SO 3 -depleted gas 2102 are ultimately recycled back to the converter 2021 so that substantially all the residual SO 2 in the SO 3 -depleted gas 2102 can eventually be converted into sulfuric acid.
  • the single pass SO 2 conversion for the whole converter 2021 is about 90.4%.
  • the overall conversion of the SO 2 in the source gas 2003 is about 99.87%.

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  • Organic Chemistry (AREA)
  • Inorganic Chemistry (AREA)
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  • Treating Waste Gases (AREA)
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Abstract

L'invention concerne la formation d'un gaz d'alimentation transformé contenant une première partie de gaz décapant enrichi de SO2. On forme un gaz de conversion contenant du SO3 et du SO2 résiduel en faisant passer le gaz d'alimentation transformé à travers plusieurs lits catalytiques en série, soit au moins deux et au maximum quatre lits catalytiques. On introduit une seconde partie de gaz enrichi de SO2 dans au moins un lit catalytique en aval du lit catalytique le plus en amont afin de renforcer la teneur en SO2 dans le gaz amené vers le lit en aval. L'invention concerne également un procédé de fabrication d'acide sulfurique et/ou d'oléum à partir d'un gaz source contenant du SO2. On forme un gaz de conversion contenant du SO3 et du SO2 résiduel en faisant passer le gaz décapant enrichi de SO2 à travers plusieurs lits catalytiques en série. Le gaz de conversion est combiné à la vapeur d'eau pour former un gaz-produit acide contenant: (a) de l'acide sulfurique formé par une réaction en phase gazeuse entre le SO3 du gaz de conversion et la vapeur d'eau, ce qui génère la chaleur de la formation de l'acide sulfurique dans la phase gazeuse; (b) du SO3; et (c) du SO2. L'énergie thermique dégagé par la chaleur en phase gazeuse de la formation de l'acide sulfurique est récupérée par transfert thermique du gaz-produit acide en vapeur ou en eau d'alimentation dans un échangeur thermique indirect. Le gaz-produit acide refroidi est ensuite mis en contact avec l'acide sulfurique liquide dans une zone d'absorption de SO3 pour former de l'acide sulfurique et/ou de l'oléum supplémentaire et un gaz SO3 appauvri contenant du SO2.
PCT/US2000/030095 1999-11-01 2000-11-01 Procede de fabrication du trioxyde de soufre, de l'acide sulfurique, et de l'oleum a partir du dioxyde de soufre WO2001036324A1 (fr)

Priority Applications (8)

Application Number Priority Date Filing Date Title
EA200200516A EA200200516A1 (ru) 1999-11-01 2000-11-01 Способ получения триоксида серы, серной кислоты и олеума из диоксида серы
BR0015265-0A BR0015265A (pt) 1999-11-01 2000-11-01 Processo para preparar trióxido de enxofre, ácido sulfúrico e oleum de dióxido de enxofre
EP00975531A EP1230150A1 (fr) 1999-11-01 2000-11-01 Procede de fabrication du trioxyde de soufre, de l'acide sulfurique, et de l'oleum a partir du dioxyde de soufre
JP2001538282A JP2003517419A (ja) 1999-11-01 2000-11-01 二酸化硫黄から三酸化硫黄、硫酸およびオレウムを製造する方法
CA002387988A CA2387988A1 (fr) 1999-11-01 2000-11-01 Procede de fabrication du trioxyde de soufre, de l'acide sulfurique, et de l'oleum a partir du dioxyde de soufre
AU13573/01A AU1357301A (en) 1999-11-01 2000-11-01 Method for making sulfur trioxide, sulfuric acid, and oleum from sulfur dioxide
MXPA02004408A MXPA02004408A (es) 1999-11-01 2000-11-01 Metodo para preparar trioxido de azufre, acido sulfurico y oleum a partir de dioxido de azufre.
KR1020027005602A KR20020049001A (ko) 1999-11-01 2000-11-01 이산화황으로부터 삼산화황, 황산 및 발연 황산의 제조 방법

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US60/163,061 1999-11-01

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JP (1) JP2003517419A (fr)
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CN (1) CN1384805A (fr)
AU (1) AU1357301A (fr)
BR (1) BR0015265A (fr)
CA (1) CA2387988A1 (fr)
EA (1) EA200200516A1 (fr)
MX (1) MXPA02004408A (fr)
WO (1) WO2001036324A1 (fr)
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Cited By (5)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
EP1918248A3 (fr) * 2006-10-29 2010-06-09 Silicon Fire AG Préparation d'H2O2 à partir d'acide sulfurique, obtenu par combustion de matières combustibles fossiles contenant de résidus de soufre, et utilisation de H2O2 en tant que source d'énergie
WO2011067046A1 (fr) * 2009-12-01 2011-06-09 Bayer Technology Services Gmbh Procédé de retraitement d'effluents gazeux sulfurés
US8679447B2 (en) 2011-01-11 2014-03-25 Albemarle Corporation Process for producing sulfur dioxide and sulfur trioxide
CN111533092A (zh) * 2020-05-09 2020-08-14 山东合生固废处置工程有限公司 一种含硫磷有机废弃物的处置和再利用的方法
CN111747382A (zh) * 2020-05-25 2020-10-09 惠州宇新化工有限责任公司 一种利用烷基化废酸回收装置进行稀酸提浓回收利用的方法及系统

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DE102005008109A1 (de) * 2005-02-21 2006-08-24 Outokumpu Technology Oy Verfahren und Anlage zur Herstellung von Schwefelsäure
DE102007027841B4 (de) * 2007-06-13 2012-02-16 Outotec Oyj Verfahren und Vorrichtung zur Mischung von Gasen
JOP20200123A1 (ar) * 2010-01-20 2017-06-16 Mecs Inc استرجاع الطاقة في تصنيع حمض السلفريك
WO2011147431A1 (fr) 2010-05-27 2011-12-01 Haldor Topsoe A/S Procédé et appareil de production d'acide sulfurique
CN102351156B (zh) * 2011-07-28 2013-03-27 湖南科技大学 一种硫酸厂的无尾气生产装置及工艺方法
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CN103405997A (zh) * 2013-07-29 2013-11-27 中国恩菲工程技术有限公司 制备硫酸的设备
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AU2016349302B2 (en) * 2015-11-06 2021-05-20 Haldor Topsoe A/S Method and plant design for reduction of start-up sulfur oxide emissions in sulfuric acid production
WO2018015138A1 (fr) * 2016-07-21 2018-01-25 Haldor Topsøe A/S Procédé de production d'acide sulfurique à partir de charges contenant du soufre avec trempe au gaz
CN106512980B (zh) * 2016-08-31 2018-11-13 上海奥威日化有限公司 由二氧化硫氧化生产三氧化硫的催化剂
CN106744714A (zh) * 2016-11-18 2017-05-31 上海东化环境工程有限公司 一种从烟气中回收二氧化硫制硫酸的工艺
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CN109603429B (zh) * 2018-11-14 2021-03-05 惠州宇新新材料有限公司 一种原料空气预处理方法
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CN111533092A (zh) * 2020-05-09 2020-08-14 山东合生固废处置工程有限公司 一种含硫磷有机废弃物的处置和再利用的方法
CN111747382A (zh) * 2020-05-25 2020-10-09 惠州宇新化工有限责任公司 一种利用烷基化废酸回收装置进行稀酸提浓回收利用的方法及系统
CN111747382B (zh) * 2020-05-25 2022-04-15 惠州宇新化工有限责任公司 一种利用烷基化废酸回收装置进行稀酸提浓回收利用的方法及系统

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CN1384805A (zh) 2002-12-11
CA2387988A1 (fr) 2001-05-25
EP1230150A1 (fr) 2002-08-14
AU1357301A (en) 2001-05-30
EA200200516A1 (ru) 2002-10-31
MXPA02004408A (es) 2002-09-02
KR20020049001A (ko) 2002-06-24

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