Absorption Systems
Absorption processes are characterized by washing the synthesis gas with a liquid solvent, which selectively removes the acid components (mainly H2S and C02) from the gas. The laden solvent is regenerated, releasing the acid components and recirculated to the absorber. The washing or absorption process takes place in a column, which is usually fitted with (dumped or structured) packing or trays.
The absorption characteristics of a solvent depend either on simple physical absorption or on a chemical bond with the solvent itself. This provides the basis for the classification of AGR systems into physical or chemical washes, which have distinctly different loading characteristics.
The loading capacity of a physical wash depends primarily on Henry’s law and is therefore practically proportional to the partial pressure of the component to be
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Figure 8-4. Equilibrium of Physical and Chemical Absorption |
removed (Figure 8-4). This leads to the fact that the solution rate for any particular operating pressure is approximately proportional to the volume of raw gas to be processed.
In contrast, the loading capacity of a chemical wash is limited by the quantity of the active component of the solution. Once a saturation level is reached only a minor additional loading can be achieved by physical absorption in the solution. The solution rate is approximately proportional to the volume of acid gas removed.
Some mixed solvents have been developed using both effects. These are known as physical-chemical washes.
Generally, solvent regeneration is achieved by one of or a combination of flashing, stripping, and reboiling. Both flashing and stripping reduce the partial pressure of the acid component. In physical washes, reboiling raises the temperature and reduces the acid gas solubility. In chemical washes the increased temperature serves to break the chemical bond. In such systems the acid components are released in the same chemical form in which they were absorbed (Figure 8-5).
An additional class of washing systems, oxidative washes, regenerate the chemically absorbed sulfur by oxidizing the active component in the solvent and recovering the sulfur in elemental form.
Amines. Solutions of amines in water have been used for acid gas removal for over 50 years. The principle amines used for synthesis gas treatment are mono — and diethanolamine (MEA and DEA), methyldiethanolamine (MDEA), and di-isopropa — nolamine (DIPA), the latter particularly as a component of the Sulfinol solvent. Others amines used in natural gas applications, such as diglycolamine (DGA) or triethanolamine (TEA), have not been able to make any significant impact in syngas applications.
MDEA is the most widely used amine today. It is more selective than primary (e. g. MEA) or secondary (e. g. DEA) amines, due to the fact that C02 is absorbed more slowly than H2S.
A number of proprietary formulations have been developed to address specific issues. For example, Ucarsol was developed to reduce corrosion with high C02 loading. BASF’s aMDEA includes an activator to accelerate C02 absorption, where selectivity is not a requirement. Variation in the degree of promotion can influence the energy requirement for regeneration. Exxon developed the Flexsorb family of hindered amines specifically for high selectivity.
Typical performance data of different amine washes may be seen in Table 8-1. The flowsheet of a typical MDEA wash is shown in Figure 8-6.
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Table 8-1 Properties of Amine Solvents |
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Standard MEA |
Inhibited MEA |
DEA |
MDEA |
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Molecular weight |
61 |
105 |
119 |
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C02 partial pressure, bar |
<100 |
<100 |
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Gas purity C02 ppmv |
20-50 |
20-50 |
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Solution strength, wt% |
10-20 |
30 |
25-35 |
30-50 |
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Solution loading, mol/mol |
0.25-0.45 |
0.4-0.8 |
0.8 |
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Energy demand, MJ/kmol C02 |
210 |
140 |
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Notes: |
selective |
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ACID GAS FLASH ABSORBER VESSEL REGENERATOR Figure 8-6. Typical MDEA Flowchart with Single Flash Stage Physical Washes |
The important characteristics for any successful physical solvent are:
• Good solubility for C02, H2S, and COS in the operating temperature range, preferably with significantly better absorption for H2S and COS compared with C02 if selectivity is an important issue for the application in hand.
• Low viscosity at the lower end of the operating temperature range. Although lowering the operating temperature increases the solubility, the viscosity governs in effect the practical limit to lowering the operating temperature.
• A high boiling point reduces vapor losses when operating at ambient or near ambient temperatures.
Rectisol. The Rectisol process, which uses cold methanol as solvent, was originally developed to provide a treatment for gas from the Lurgi moving-bed gasifier, which in addition to H2S and C02 contains hydrocarbons, ammonia, hydrogen cyanide, and other impurities.
In the typical operating range of -30 to -60°C, the Henry’s law absorption coefficients of methanol are extremely high, and the process can achieve gas purities unmatched by other processes. This has made it a standard solution in chemical applications such as ammonia, methanol, or methanation, where the synthesis catalysts require sulfur removal to less than 0.1 ppmv. This performance has a price, however, in that the refrigeration duty required for operation at these temperatures involves considerable capital and operating expense.
Methanol as a solvent exhibits considerable selectivity, as can be seen in Table 8-2. This allows substantial flexibility in the flowcharting of the Rectisol process and both standard (nonselective) and selective variants of the process are regularly applied according to circumstances.
As a physical wash, which uses at least in part flash regeneration, part of the C02 can be recovered under an intermediate pressure. Typically, with a raw gas pressure of 50 bar, about 60-75% of the C02 would be recoverable at 4-5 bar. Where C02 recovery is desired, whether for urea production in an ammonia application or for sequestration, this can provide significant compression savings.
Figure 8-7 shows the selective Rectisol variant as applied to methanol production. The incoming raw gas is cooled down to about -30°C, the operating temperature of the H2S absorber. Both H2S and COS are washed out with the cold methanol to a residual total sulfur content of less than lOOppbv. The desulfurized gas is then shifted outside the Rectisol unit, the degree of shift being dependent on the final product. Carbon dioxide is then removed from the shifted gas in the C02 absorber to produce a raw hydrogen product. This column is divided into two sections: a bulk C02 removal section using flash regenerated methanol, and a fine C02 removal section in
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Table 8-2 Properties of Physical Solvents |
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Methanol NMP |
DMPEG |
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Chemical Formula |
CH3OH |
CH3N-(H2C)3 CH30(C2H40)3 |
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c=o |
ch3 |
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Mol. Weight Boiling point |
kg/kmol |
32 |
99 |
178 to 442 |
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at 760Torr |
°С |
64 |
202 |
213 to 467 |
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Melting point |
°С |
-94 |
-24.4 |
-20 to -29 |
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Viscosity |
cP |
0.85 at -15°C |
1.65 at 30°C |
4.7 at 30°C |
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1.4 at -30°C |
1.75 at 25°C |
5.8 at 25°С |
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2.4 at -50°C |
2.0 at 15°C |
8.3 at 15°C |
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Specific mass Heat of |
kg/m3 |
790 |
1.027 |
1.031 |
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evaporation Specific heat |
kJ/m3 |
1090 |
533 |
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at 25°C Selectivity |
kJ/kg. K |
0.6 |
0.52 |
0.49 |
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at working temperature |
(H2S:C02) |
1:9.5 |
1:13 |
1:9 |
which hot-regenerated methanol is used. The C02 removal section operates at lower temperatures, typically about -60°C. The permissible C02 slip is dependent on the application. For methanol synthesis gas 1 mol% residual C02 in the raw hydrogen is quite adequate. For hydrogen production based on methanation, typically 100 ppmv would be appropriate. For ammonia where the gas is subsequently treated in a cryogenic nitrogen wash, lOppmv would be typical.
Following the solvent circuit, we see first an intermediate H2S flash from which co-absorbed hydrogen and carbon monoxide are recovered and recompressed back into the raw gas. The flashed methanol is then reheated before entering the hot regenerator. Here the acid gas is driven out of the methanol by reboiling, and a Claus gas with an H2S content of 25-30% (depending on the sulfur content of the feedstock) is recovered. Minor adaptations are possible to increase the H2S content if desired.
The hot-regenerated methanol, which is the purest methanol in the circuit, is used for the fine C02 removal. The methanol from the C02 removal is subjected to flash regeneration in a multistage flash tower. The configuration shown is typical for the methanol applications with only atmospheric flash regeneration. For hydrogen or ammonia applications where better absorption is required, the final flash stage may be under vacuum, or it may use stripping nitrogen from the air separation plant. Finally, the loop is closed with the flash regenerated methanol returning to the H2S absorber.
Water entering the Rectisol unit with the syngas must be removed, and an additional small water-methanol distillation column is included in the process to cope with this.
Typically, the refrigerant is supplied at between -30 and -40°C. Depending on application, different refrigerants can be used. In an ammonia plant, naturally, ammonia is used, and the refrigeration system is integrated with that of the synthesis. In a refinery environment, propane or propylene may be the refrigerant of choice.
The Rectisol technology is capable of removing not only conventional acid gas components but also, for example, HCN and hydrocarbons. Supp (1990, p. 83) describes a typical hydrocarbon prewash system. Mercury capture using Rectisol as a cold trap to condense out metallic mercury is also documented (Koss, Meyer, and Schlichting 2002).
Selexol. The Selexol process was originally developed by Allied Chemical Corporation and is now owned by UOP. It uses dimethyl ethers of polyethylene glycol (DMPEG). The typical operating temperature range is 0-40°C. The ability to operate in this temperature range offers substantially reduced costs by eliminating or minimizing refrigeration duty. On the other hand, for a chemical application such as ammonia, the residual sulfur in the treated gas may be 1 ppmv H2S and COS each (Kubek etal. 2002) which is still more than the synthesis catalysts can tolerate. This is not an issue, however, in power applications where the sulfur slip is less critical. Selexol has a number of references for such plants including the original Cool Water demonstration unit and most recently the 550 MW Sarlux IGCC facility in Italy.
The ratio of absorption coefficients for H2S, COS, and C02 is about 1:4:9 in descending order of solubility (Kubek, Polla, and Wilcher 1997). A plant designed for, say, 1 ppm COS in the clean gas would require about four times the circulation rate of a plant for 1 ppm H2S, together with all the associated capital and operating costs. In a gasification environment it is therefore preferable to convert as much COS to H2S upstream of a Selexol wash. In a plant using raw gas shift for hydrogen or ammonia, this will take place simultaneously on the catalyst with the carbon monoxide shift. Where no CO shift is desired, then COS hydrolysis upstream of the Selexol unit provides a cost-effective solution to the COS issue.
Other characteristics favorable for gasification applications include high solubilities for HCN and NH3 as well as for nickel and iron carbonyls.
The Selexol flowsheet in Figure 8-8 exhibits the typical characteristics of most physical absorption systems. The intermediate flash allows co-absorbed syngas components (H2 and CO) to be recovered and recompressed back into the main stream. For other applications, including H2S concentration in the acid gas or separate C02 recovery, staged flashing techniques not shown here may be applied.
Purisol. NMP or n-methyl-pyrrolidone is the solvent used in Furgi’s Purisol process. The operating range is 15°C to 40°C. The selectivity for H2S/C02 is extremely high and largely independent of the operating temperature (Griinewald 1989). Solvent properties are included in Table 8-2. The characteristics are in many ways comparable with Selexol.
RICH/LEAN
EXCHANGER
Some gas-washing systems exploit the principles of both physical and chemical absorption and are known as physical-chemical washes. They generally use an amine together with organic physical solvent. They can usually accept a higher loading than an aqueous amine solution, thus reducing solvent rates. Furthermore, the organic solvents applied in such systems accelerate the hydrolysis of COS to H2S in the lower sections of the column, thus permitting an improved total sulfur removal performance than a pure amine system. Other aspects, which still need review when considering a physical-chemical system, are the potential for amine degradation, which is generally unchanged compared with the equivalent aqueous amine system. Their effectiveness at absorbing metal carbonyls is not documented and so must be considered as unproven.
Sulfinol. Shell’s Sulfinol solvent in its original form was a mixture of DIP A and Sulfolane (tetrahydrothiophene dioxide). The former provides a chemical solvent and the latter a physical solvent. Meanwhile a modified solvent, known as m-Sulfinol has been developed that uses MDEA as the chemical component. The original Sulfinol formulation has been used successfully downstream of a large number of small
oil gasifiers for the production of oxo-synthesis gas. The AGR at the Buggenum IGCC is an example of a larger m-Sulfinol unit.
AmisoL The Amisol process was developed by Lurgi using a mixture of MEA or DEA with methanol. It has been applied downstream of a number of oil gasification units, but it has not established a wide market. Details can be found in Supp (1990) and Kriebel (1989).
Oxidative washes or liquid redox systems differ from other types of absorption system in that the H2S in the acid gas is oxidized directly to elemental sulfur in the absorption stage. The active agent in the solution is regenerated in a separate oxidizing vessel, which also serves to separate the solid elemental sulfur from the solution. The solvents of oxidative washes absorb essentially only H2S, but not C02 nor COS. This makes them suitable for applications where H2S must be removed from a stream containing large quantities of C02, even if the H2S partial pressure is low.
There is no known existing application in a gasification environment, but such washes exhibit potential as a substitute for a Claus plant, where the gasifier feed has very low sulfur content and the sour gas is unsuitable for treatment in a Claus plant.
Earlier plants, notably the Stretford and Takahax processes, used vanadium-based agents, which undergo a valence change from the pentavalent to the tetravalent state during the absorption stage. Modern processes, of which Lo-Cat and Sulferox are the best known, use chelated iron formulations.
The Lo-Cat process can be arranged in a number of different application-dependant configurations of which that shown in Figure 8-9 is typical. Acid gas enters the absorber, where the H2S is absorbed into the aqueous chelated iron solution. The
Figure 8-9. Lo-Cat Flowsheet (Source: Adapted from
ferric iron oxidizes the HS ion to elemental sulfur according to reaction 8-1. The iron is reduced to the ferrous state.
HS“+2Fe+3 ^ S°+2Fe+2 + H+ (8-1)
In the oxidizer, the sulfur settles out and is transferred to a vacuum filter where it is separated from the solution as a cake. Air is blown into the oxidizer, where oxygen is absorbed into the solution and oxidizes the ferrous iron back to the ferric state (reaction 8-2) for recirculation back to the absorber.
2Fe+2 + 1Л02( )+H20. -» 2Fe+3 + 2 OH“ (8-2)
The raw sulfur from the vacuum filter is typically 65% to 85% sulfur, the remainder being water and dissolved salts including iron. This product requires further treatment to meet generally accepted market quality, melting (to remove the water) and filtering being important process steps. Nonetheless, the usual “bright yellow” color specification for commodity sulfur is not met, and specialized applications need to be located in the marketplace.
Generally, liquid redox systems are applied for small plants, in particular where H2S concentrations are lower than can be handled by the Claus process (see Section 8.4). The Stretford process was regularly applied for Claus tail gas processing as part of the Beavon tail gas treating process. A similar application using an iron chelate process was put into service in 2001 (Nagl 2001).