CA2765673C - Acidic gas capture by diamines - Google Patents

Abstract

Compositions and methods related to the removal of acidic gas. In particular, the present disclosure relates to a composition and method for the removal of acidic gas from a gas mixture using a solvent comprising a thermally stable amine and carbon dioxide. One example of a method may involve a method for removing acidic gas comprising contacting a gas mixture having an acidic gas with a solvent, wherein the solvent comprises a thermally stable amine in an amount from about 3 to about 20 moles/kg of water, and carbon dioxide in an amount of from about 0.1 to about 0.6 moles per mole of a thermally stable amine.

Description

ACIDIC GAS CAPTURE BY DIAMINES
BACKGROUND
As concerns of global climate changes spark initiatives to reduce carbon dioxide emissions, its economic removal from gas streams is becoming increasingly important.
Removal by absorbtion/stripping is a commercially promising technology, as it is well suited to sequester carbon dioxide (C02). Such carbon dioxide emissions may be produced by a variety of different processes, such as the gas stream produced by coal-fired power plants.
The removal of CO2 can be an expensive process, potentially increasing the cost of electricity by 50% or more. Therefore, technology improvements to reduce the costs associated with the removal are highly desirable.
The removal of CO2 from fuel gas and flue gas by absorption/stripping with aqueous amines is a disclosed and commercially practiced technology. A typical flowsheet for such a process is give by Kohl and Nielsen (1997) (Figure 1). The gas at 30 to 50 C
containing CO2 and inerts such as methane, hydrogen, or nitrogen is contacted countercurrently in a trayed or packed column with lean aqueous solvent entering at 30 to 50 C. The aqueous rich solvent containing 3 to 6 molar amine is heated by cross exchange with the hot lean solvent. The approach temperature for this exchanger has historically been 10 to 30 C with a lean solution loading of 0.01 to 0.25 moles C02/mole amine. CO is removed from the solvent at 1.5 -2 atm and 90-130 C in a countercurrent reboiled stripper with trays or packing.
Commercially used amines that are used by themselves in water include monoethanolamine, diethanolamine, methyldiethanolamine, diglycolamine, diisopropanolamine, some hindered amines, and others (Kohl and Nielsen (1997)). These amines are soluble or miscible with water at ambient temperature at high concentrations that are used in the process to maximize capacity and reduce sensible heat requirements. Other amines, including piperazine, are used in combination with methyldiethanolamine and other primary amines.
A number of mono- and polyamines, including piperazine, are identified as potentially useful solvent components but have not been used because they are insufficiently soluble in water when used by themselves. Piperazine is a diamine that has previously been studied as a promoter for amine systems to improve kinetics. In water at 25 C, solid piperazine has a solubility less than 2 M, so it cannot be used in traditional systems at HOU03:1200893.3 concentrations that give adequate CO2 capacity for good energy performance.
BASF has disclosed the used of piperazine in combination with other amines (such as alkanolamines) or highly water soluble organics (such as triethyleneglycol) to promote the water solubility of piperazine.
It has also been claimed that number of potentially useful amines such as piperazine would be too volatile if used in high concentrations in aqueous solvents. The boiling point of piperazine (146.5 C) is lower that that of monoethanolamine (170 C), so the use of Raoult's law would suggest that it would have a greater volatility at the top of the absorber.

DRAWINGS
Some specific example embodiments of the disclosure may be understood by referring, in part, to the following description and the accompanying drawings.
Figure 1 shows a typical flowsheet for aqueous amine absorption/stripping for removal.
Figure 2 shows an exemplary double-matrix stripper configuration.
Figure 3 shows an exemplary internal exchange stripper configuration.
Figure 4 shows an exemplary multipressure stripper configuration with a split feed.
Figure 5 shows an exemplary flashing feed stripper configuration.
Figure 6 shows an exemplary double-matrix stripper configuration with exemplary operating parameters.
Figure 7 shows an exemplary internal exchange stripper configuration with exemplary operating parameters.
Figure 8 shows an exemplary multipressure stripper configuration with a split feed with exemplary operating parameters.
Figure 9 shows an exemplary flashing feed stripper configuration with exemplary operating parameters.
Figure 10 shows an exemplary multistage flash stripper configuration.
Figure 11 shows an exemplary generalized flowsheet for multistage stripping.
Figure 12 shows an exemplary generalized flowsheet for multistage stripping.
Figure 13 shows a solid-liquid transition temperature for aqueous PZ.

Figure 14 shows a comparison of solid solubility for aqueous PZ solutions.
HOU03:1200893.3 Figure 15 shows viscosity of amine solutions at typical rich loading and 40 C.
Figure 16 shows CO2 solubility in aqueous PZ solutions ranging from 0.9 to 8 m PZ
and from 40 to 100 C.

Figure 17 shows a comparison of mass transfer coefficients in 8 m PZ and 7 m MEA
from 40 to 100 C.

Figure 18 shows a comparison of PZ and MEA volatility normalized to amine concentration.
Figure 19 shows and exemplary flowsheet of a three stage flash.

Figure 20 shows equivalent work for stripping with 5 C approach and rich P*CO2 of 5 kPa for 8 m PZ.
While the present disclosure is susceptible to various modifications and alternative forms, specific example embodiments have been shown in the figures and are described in more detail below. It should be understood, however, that the description of specific example embodiments is not intended to limit the invention to the particular forms disclosed, but on the contrary, this disclosure is to cover all modifications and equivalents as illustrated, in part, by the appended claims.

DESCRIPTION
The present disclosure, according to certain embodiments, generally relates to compositions, systems, and methods for the removal of acidic gas. In particular, the present disclosure relates to compositions, systems, and methods for the removal of acidic gas from a gas mixture using a solvent comprising a thermally stable amine (e.g., piperazine) and carbon dioxide. Thermally stable amines generally refers to amines that are functional at elevated temperatures. For example, thermally stable amines may be stable up to about 130 C, 140 C, 150 C, and 170 C. Examples of suitable thermally stable amines include, but are not limited to, piperazine (PZ) and various substituted piperazines (e.g., methylpiperazine, dimethylpiperazine, ethylpiperazine, and diethylpiperazine), morpholine, 5-amino-l-pentanol, 2-amino-2-methyl-l-propanol (AMP), diglycolamine (DGA ), 4-amino-l-butanol, 3-amino-l-propanol, hydroxyethylpiperazine (HEP), 1-amino-2-propanol, methyldiethanolamine (MDEA), 2-amino-l-propanol.

HOU03:1200893.3 The present disclosure is based in part on the discovery of optimum stripper process configurations and operating conditions that result in unexpectedly high lean loading of CO2.
Such process configurations may include the matrix, internal exchange, flashing feed, multipressure, and stripper processes described herein. Such operating conditions may include an unexpectedly low exchanger approach temperature. In some embodiments, such exchanger approach temperatures may be approximately 5 C.
The present disclosure is also based in part on the discovery that thermally stable amines may be less volatile in an aqueous solution than expected from Raoult's law. In certain embodiments, the activity coefficient of a thermally stable amine (e.g., piperazine) at infinite dilution in water may be about 0.05, whereas monoethanolamine (1VIEA) has an activity coefficient of about 0.16.
The present disclosure is also based in part on the discovery that when thermally stable amine solutions are loaded with about 0.1 to about 0.6 moles carbon dioxide per amine equivalent, the volatility of the thermally stable amine may be further reduced. As used herein, loading refers to moles C02/mole alkalinity where monoamine have one mole alkalinity per mole of amine and diamines have two moles of alkalinity per mole amine. In certain embodiments, the loading may be 0.25 to 0.45 moles carbon dioxide per amine equivalent. Such a reduction may occur at least in part because of the formation of carbamate ions. Such a reduction may result in the ability to produce concentrated solutions of thermally stable amine loaded with CO2 which have a volatility acceptable for use in the methods of the present disclosure.
The present disclosure is also based in part on the discovery that the total solubility of a solid thermally stable amine may be enhanced in solutions loaded with CO2.
In certain embodiments, the present disclosure provides solutions comprising from about 3 in to about 20 in (moles thermally stable amine/kg water) total thermally stable amine when said solutions are loaded with from about 0.1 to about 0.6 moles CO2 per amine equivalent. This increase in solubility may be due in part to the formation of carbamate ions.
In certain embodiments, solutions comprising from about 4 in to about 12 in (moles thermally stable amine/kg water) total thermally stable amine. In certain embodiments, the solutions are loaded with from about 0.25 to about 0.45 moles CO2 per amine equivalent.

HOU03:1200893.3 The present disclosure is also based in part on the discovery that concentrated aqueous thermally stable amines may be more stable to oxidative and/or thermal degradation as compared to conventional solutions, such as MEA. In certain embodiments, the presence of dissolved iron may catalyze the degradation of MEA at a higher rate than the degradation of thermally stable amine. In certain embodiments, solutions of thermally stable amine loaded with CO2 may not degrade significantly even at temperatures as high as 150 C, whereas MEA may undergo significant degradation (up to about 50%) at 120 C. Thus, in certain embodiments, the present disclosure provides solutions comprising a thermally stable amine which may be used advantageously at higher pressures and/or temperatures. For example, the solutions comprising a thermally stable amine may be used at temperatures less than 175 C.
Such an ability to operate at higher pressures and/or temperatures may, among other things, reduce the amount of energy necessary to perform the methods of the present disclosure. In certain embodiments, such a reduction of the amount of energy may range from about 10% to about 30%. Additionally, solutions comprising thermally stable amine may absorb CO2 at faster rates. In certain embodiments, the use of solutions comprising a thermally stable amine may result in increased in CO2 absorption rates ranging from about 20% to about 100%. Such increased CO2 absorption rates may, among other things, enable absorber configurations which require less packing and pressure drop.
When used in the methods of the present invention, the thermally stable amine may be recovered following absorption of CO2. In certain embodiments, such recovery may occur through an evaporation process using a thermal reclaimer.
In certain embodiments, the present disclosure provides a method for the removal of acidic gases from a gas mixture comprising contacting the gas mixture with a solvent comprising a thermally stable amine in an amount from about 0.1 to about 0.6 moles carbon dioxide per amine equivalent.
While the present disclosure primarily discusses removal of C02, any acidic gas capable of removal by the methods of the present invention is contemplated by the present disclosure. Such acidic gases may include, but are not limited to, hydrogen sulfide (1-12S) or carbonyl sulfide (COS), CS2, and mercaptans.

The gas mixture may be any gas mixture comprising CO2 for which CO2 removal is desired and which is compatible with (i.e. will not be adversely affected by, or will not HOU03:1200893.3 adversely react with) the methods of the present disclosure. In certain embodiments, the gas mixture may comprise any gas mixture produced as the byproduct of a chemical process.
Suitable gas mixtures may comprise one or more of natural gas and hydrogen.
Process Configurations In certain embodiments, the present disclosure provides several process configurations that may be useful in the methods of the present disclosure.
The choice of process configuration may depend upon a number of factors, including, but not limited to, the composition of the gas mixture, the desired amount of CO2 removal, the concentration of thermally stable amine to be used, and resource or environmental considerations.
One type of process configuration that may be useful in the methods of the present invention is a matrix stripper configuration. In certain embodiments, such a matrix stripper configuration may be a two-stage matrix, such as the configuration shown in Figure 2. In such a two-stage matrix configuration, the temperature change across the stripper may be reduced without the inefficiencies that may be associated with mechanical compression. The rich solution from the absorber may be split into two streams. The first stream may be sent to the first stripper at a higher pressure, which may result in a slightly superheated feed. Heat may be applied via reboiler steam. The lean solution from the first column may be the semirich feed to the middle of the second column, which may operate at a lower pressure.
The other rich stream may be fed to the top of the second stripper. The second column may produce a semilean stream and a lean stream. The semilean stream may be crossexchanged with the rich feed to the second column, while the lean solution may be crossexchanged with the rich solution to the first stripper. The water vapor from the overhead of the second column may be condensed, and the CO2 may be sent to the first stage of the compression train. The water vapor in the overhead from the first column may be condensed, and the C02 may be sent to the second stage in the compression train. The compression work in this configuration may be reduced due at least in part to recovery of a portion of the C02 at a higher pressure, which may reduce the need for compression downstream. In certain embodiments, the lower pressure column may be set to 160 kPa for normal pressure operations. In certain embodiments, the lower pressure column may be set to 30 kPa for vacuum operations. The pressure of the higher-pressure column and the flow into the flash section may be optimized to minimize the total equivalent work of the system.
Even though a HOU03:1200893.3 two-stage matrix is described in the present disclosure, a three-stage matrix may also be used with reduced energy requirement.
Another type of process configuration that may be useful in the methods of the present invention is multistage flash stripper configuration. In certain embodiments, such multistage flash strippers may be configured as a multistage flash with a multistage intercooled compressor as illustrated in Figure 3. Cold rich solvent from the absorber is heated by cross exchange with hot lean solution from the last stage. At stage n rich solution is heated then flashed to a lower pressure (Pn) to release CO2 with some water vapor. The vapor from the flash tank is combined with vapor from the next stage (n+l), intercooled to condense water and compressed to the pressure (Pi-1) of the previous stage.
Lean solution from the last stage is returned to the absorber through the cross exchanger.
The process may be optimized to select a number of stages from 1 to 6, a pressure ratio (Põ
/Põ+1) from stage to stage of 1.2 to 10, and a heat rate at each stage from 0 to 200 kJ/mol CO2.
The temperature of the flash tank may practically vary from 80 to 175 C. This configuration will be especially attractive with flash tank temperature from 120 to 170 C when used with thermally stable amines such as piperazine that do not degrade at the elevated temperature. The most attractive configuration with concentrated piperazine solution might use 3 stages, each at 140 to 150 C, with the about the same heat rate, and with approximately equal pressure ratios.
Another type of process configuration that may be useful in the methods of the present invention is an exchange stripper configuration. In certain embodiments, such an exchange stripper configuration may be an internal exchange stripper, such as the configuration shown in Figure 4. Among other things, this configuration integrates the stripping process with heat transfer. In certain embodiments, this configuration may approach the theoretical limit of adding and removing material and energy streams along the entire column. Similar configurations have been described previously by Leites et al.
and Mitsubishi. In certain embodiments, this configuration may alleviate the temperature drop across the stripper by exchanging the hot lean solution with the solution in the stripper. In certain embodiments, the configuration may comprise a continuous heat exchange surface, which may allow for countercurrent heat exchange of the hot-lean solution with the solution passing through the stripper. In certain embodiments, a large overall heat transfer capability of 41.84 W/K-mol solvent per segment may be used. Such a heat transfer capability may HOU03:1200893.3 result in a typical AT of about 1.2 K and about 3 K in the internal exchanger for the vacuum operation, and for operation at normal pressure, respectively.
Another type of process configuration that may be useful in the methods of the present invention is a multipressure configuration. In certain embodiments, such a multipressure configuration may be a multipressure configuration with a split feed, such as the configuration shown in Figure 5. Similar multipressure configurations have been described in our previous work. In certain embodiments, this configuration may take a 10%
split feed from the liquid flowing from the middle to the lowest pressure level in a multipressure stripper, and it may send this stream to an appropriate point in the absorber. In certain embodiments, the temperatures at the bottom of the stripper pressure sections may be equal, and heat may be added to each stripper pressure section to achieve isothermal operation in each section. Such operating conditions, among other things, may reduce irreversibilities and work loss. Among other things, this configuration may take advantage of the favorable characteristics of the multipressure configuration and the split flow configurations. In certain embodiments, the middle pressure may be configured to be approximately the geometric mean of the top pressure and the bottom pressure.
Another type of process configuration that may be useful in the methods of the present invention is a flashing feed configuration. An example of such a configuration is shown in Figure 6. In certain embodiments, this configuration may comprise special configurations of the split flow concept described by Leites et al. and Aroonwilas. In certain embodiments, at least a fraction of the rich stream may be sent to the middle of the stripper, where, after stripping, a lean solution may exit at the bottom. The rich solution may be cross-exchanged with the lean solution exiting the stripper bottom. In certain embodiments, the vapor leaving the stripper may then be contacted with the absorber rich flow in a five-staged upper section where the latent heat of water vapor may be used to strip the C02 in the `cold feed' and a semilean stream may be produced. In certain embodiments, the semilean product may be cross-exchanged with the rich solution fed to the upper section. In certain embodiments, the reboiler duty may remain substantially unchanged, and `free stripping' may be achieved in the upper section. In certain embodiments, the split ratio of the rich streams into the middle and upper sections may be optimized to minimize equivalent work.

HOU03:1200893.3 The choice of operating conditions for each of these process configurations may depend upon a number of factors, including, but not limited to, the composition of the gas mixture, the desired amount of CO2 removal, the concentration of piperazine to be used, and resource or environmental considerations. Examples of suitable operating conditions are shown in Figures 7, 8, 9, and 10 for the double matrix, internal exchange, multipressure with split feed, and flashing feed stripper configurations, respectively.
Another type of process configuration that may be useful in the methods of the present disclosure is a multistage stripper configuration and methods for multistage stripping that may be used at temperatures from about 120 C to about 160 C with thermally stable amines. The process and configuration also may be used at lower temperatures.
An example of such a configuration is shown in Figure 11.
Figure 11 provides a generalized flowsheet example for such an embodiment. In operation, rich solution is heated by exchange with hot lean solution. In each of N stages the hot lean solution, LL_,, is preheated with steam or another convenient source of heat. The rich solution is then distributed at the top of a gas/liquid contacting section in the stage j stripper, usually a packed column. The stripper is reboiled with heat provided by steam or another convenient source to the maximum temperature, T. The hot semilean solution, LL, is then sent on to stage J+1.
The vapor from the stripper at pressure, Pj, is sent to the intercooler of stage J of the compressor. The intercooler may be cooled by cooling water or it may serve as a source of useful heat. Because the high temperature stripper produces vapor as hot as 150 C, useful heat can obtained at 150 to 80 C from both the sensible heat and latent heat of water vapor as the stream is cooled. The recovered heat could include boiler feedwater preheating. The recovered heat could also be used as in multieffect evaporation to heat a similar generalized multistage stripper at a lower temperature, such as 100 to 120 C. Condensed water is separated from the cooled vapor. The CO2 vapor is compressed to the pressure of the previous stage, Pi-1.
Any number of stages may be used. With only one stage, the system is very much like a conventional simple stripper. Two or three stages may be optimal for many applications. After stage N the hot lean solution, LN, is cooled in the exchanger before being returned to the absorber.

HOU03:1200893.3 In certain embodiments, elements of the generalized flowsheet described above can be deleted to provide simpler effective flowsheets. Usually a useful flowsheet will use a preheater without a reboiler or a reboiler without a preheater. One version of a simple two stage heated flash would delete the packing and reboiler in both stages. Both of the stages would operate at the same temperature, from 80 to 160 C. The preferred temperature with an amine that is resistant to thermal degradation, such as piperazine, is 130 to 160 C. A second version of the two stage heated flash would delete the packing and preheater in both stages.
Another useful two-stage configuration would delete the packing and reboiler from stage 1 and the preheater from stage 2.
In the most likely configuration the heat for all of the preheaters and reboilers will be provided by steam at the same temperature. However, if steam or recovered heat is available at multiple temperatures the optimum configuration may used unequal temperature in preheaters or reboilers at the same or different stages.
In certain embodiments, the present disclosure provides a multistage stripper configuration and methods for multistage stripping that may be used at temperatures from about 120 C to about 160 C with thermally stable amines with integrated heat recovery useing four compressor stages and two exchangers. The process and configuration also may be used at lower temperatures. An example of such a configuration is shown in Figure 12. In 70 to 90% of the rich solution would be fed to Exchanger,. The heated rich solution, L1 is fed to Preheater, heated by steam to 150 C. Without using packing or reboilerl, the solution is flashed in Stripper 1 at 16 atm. The semilean solution, L2, is heated in Preheater2 to 150 C
with steam and flashed without packing or reboiler2 in Stripper2 at 8 atm. The hot lean solution is returned through Exchanger, to the absorber. 10 to 30% of the rich solution from the absorber is fed through Exchanger2 to Preheater3 and heated by Heat Recovery, and/or Heat Recovery2 to 110 C. Preheater3 and Heat Recovery, may be the same heat exchanger.
It is then flashed in Stripper3 at 4 atm without packing or reboiler3. The semilean solution from Stripper3 is fed through Preheater3 and heated to 110 C at 2 atm.
Preheater4 is heated by and may be the same heat exchanger as Heat Recovery, and/or Heat Recovery2.
Preheater4 may also use heat from high temperature intercooling of other compressor stages or from other sources such as hot flue gas before the flue gas desulfurization system.
The Semilean HOU03:1200893.3 solution is then flashed at 2 atm in Stripper4 without packing or Reboiler4.
The hot lean solution, L4, is returned to the absorber through Exchanger2.
Therefore, the present invention is well adapted to attain the ends and advantages mentioned as well as those that are inherent therein. While numerous changes may be made by those skilled in the art, such changes are encompassed within the spirit of this invention as illustrated, in part, by the appended claims.
EXAMPLES
Thermal Stability of Amines.
Example amines were screened for thermal stability based on loss of all amines after 4 weeks at 135 C with a loading of 0.4 mol C02/mol alkalinity. In this example, and under these conditions, thermally stable amines were those that demonstrated less than 37%
degradation.
Table 1: Thermal degradation screening for loss of all amines.

Amine Initial Concentration Loss of Amine (m) (%) Piperazine (PZ) 3.5 0 Morpholine 7 0 5-amino-1-pentanol 7 7

2-amino-2-methyl-1-propanol (AMP) 7 9 Diglycolamine (DGA ) 7 9 4-amino-1-butanol 7 10

3 -amino-1-propanol 7 13 Hydroxyethylpiperazine (HEP) 3.5 13 1 -amino-2-propanol 7 20 Methyldiethanolamine (MDEA) 8.4 (50 wt%) 33 2-amino-1-propanol 7 33 Monoethanolamine (MEA) 7 37 Aminoethylpiperazine (AEP) 2.33 37 Ethylenediamine (EDA) 3.5 45 6-amino-1-propanol 7 51 2-piperidine methanol (2PD) 7 73 Diethylenetriamine (DETA) 2.33 94 Hydroxyethylethylenediamine (HEEDA) 3.5 98 HOU03:1200893.3 Carbon dioxide capture with concentrated, aqueous piperazine.
Concentrated, aqueous piperazine (PZ) was investigated as a novel amine solvent for carbon dioxide (C02) absorption. The CO2 absorption rate with aqueous PZ is more than double that of 7 m MEA and volatility at 40 C ranges from 7 to 20 ppm. Thermal degradation is negligible in concentrated PZ solutions up to a temperature of 150 C, a significant advantage over MEA systems. Oxidative degradation of concentrated PZ
solutions is appreciable in the presence of copper (4 mM), but negligible in the presence of chromium (0.6 mM), nickel (0.25 mM), iron (0.25 mM), and vanadium (0.1 mM).
Initial system modeling suggests that 8 m PZ will use 10 to 20% less energy than 7 m 1VIEA. The fast kinetics and low degradation rates suggest that concentrated PZ has the potential to be a preferred solvent for CO2 capture.
Materials and methods Solution preparation. Aqueous piperazine solutions were created by heating anhydrous piperazine (99% pure, Fluka) with water until the solid crystals melted into a solution. The warm solution was transferred to a glass cylinder with a CO2 gas sparger and the cylinder was placed on a scale. The scale was used to gravimetrically add CO2 to achieve the desired loading.
CO2 loading through total inorganic carbon (TIC). The concentration of CO2 in solution was determined by total inorganic carbon analysis (Hilliard, 2008).
The sample is diluted and then acidified in 30 wt% phosphoric acid to release aqueous C02, carbamate, and bicarbonate species as gaseous CO2. The CO2 is carried in a nitrogen stream to an infrared analyzer which detects and records changes in voltage. The resulting voltage peaks are integrated and correlated to CO2 concentrations using a 1000 ppm inorganic carbon standard made from a mixture of potassium carbonate and potassium bicarbonate. CO2 loading is reported as moles CO2 per mole alkalinity or moles CO2 per equivalence of PZ, where two moles of alkalinity per mole PZ is the conversion factor.
Amine titration. The concentration of piperazine in solution was determined using acid titration (Hilliard, 2008). An automatic Titrando series titrator with automatic equivalence point detection was used (Metrohm, USA). A 300X diluted sample was titrated with 0.1 N H2SO4 to a pH of 2.4. The amount of acid needed to reach the equivalence point at a pH of 3.9 was used to calculate the total amine concentration in solution. This HOU03:1200893.3 equivalence point represents the addition of two protons to the PZ molecule creating a diprotonated PZ molecule. Additional equivalence points seen prior to 3.9 were not used in the analysis.

Viscosity measurements. Viscosity was measured using a Physica MCR 300 cone and plate rheometer (Anton Paar GmbH, Graz, Austria). The apparatus allows for precise temperature control for measuring viscosity at temperatures ranging from 20 to 70 C. To determine viscosity, the angular speed of the top disk (cone) is increased from 100 to 1000 s-1 over a period of 100 seconds and the shear stress exerted by the solution is measured every seconds. Reported viscosities are averages of these 10 individual measurements.
10 Oxidative degradation. Oxidative degradation experiments were performed in a low gas flow agitated reactor with 100 mL/min of a saturated 98%/2% 02/CO2 gas mixture fed into the headspace (Sexton, 2008). The reactor is a 500-mL jacketed reactor is filled with 350 mL of solvent. The jacket contains circulated water maintained at 55 C.
The reactor is agitated at 1400 rpm to increase the mass transfer of oxygen into the solution. The reactor is operated continuously for 3-5 weeks, depending on the experiment. Liquid samples are taken every two days and water is added to maintain the water balance on the reactor contents. The liquid samples were analyzed for PZ concentration, CO2 loading, and degradation products by acid titration, TIC, and cation and anion chromatography, respectively.
Vapor-liquid equilibrium. CO2 solubility and amine volatility were measured in a batch equilibrium cell with gas recycle through a hot gas FTIR (Hilliard, 2008). The cell was a jacketed, glass reactor where temperature is controlled within 1 C. The inlet gas is sparged from the bottom of the reactor and there is additional mechanical agitation to enhance mass transfer. The gas in the headspace of the reactor is continuously sampled by an FT-IR. The gas leaves the reactor and passes through a mist eliminator and into a sample line heated to 180 C. The heated gas stream is then analyzed by the multi-component FTIR
analyzer and recycled to the reactor as the inlet gas stream.
Thermal degradation. Thermal bombs were constructed from 1/4 or 3/8-inch stainless steel tubing with two Swagelok end caps (Davis, 2008). Bombs were filled with 2 or 10 mL
of PZ solution, sealed, and placed in forced convention ovens at multiple different temperatures. Individual bombs were removed from the ovens each week and the contents were analyzed for degradation products, remaining amine concentration, and CO2 loading.
HOU03:1200893.3 Amine losses are reported as the percent of amine lost compared to the initial amine concentration as analyzed using cation chromatography.
Wetted-wall column operation. The wetted wall column counter-currently contacts an aqueous piperazine solution with a saturated N2/CO2 stream on the surface of a stainless steel rod with a known surface area (Cullinane and Rochelle, 2006; Dugas, 2008). The wetted wall column can either perform absorption or desorption of CO2 depending on the inlet CO2 partial pressure of gas phase. By bracketing CO2 partial pressures that result in absorption and desorption, the equilibrium partial pressure of the solution can be determined.
The gas flow rate entering the wetted wall column is controlled via mass flow controllers. Inlet and outlet CO2 concentrations are measured by Horiba CO2 analyzers. As Equation 1 shows, the calculated CO2 flux divided by the CO2 partial pressure driving force provides an overall mass transfer coefficient for the experiment (KG). The overall mass transfer coefficient is related to the liquid and gas phase mass transfer coefficients via a series resistance relationship shown in Equation 2.

Flux = KG (PC02,bulk - P*C02) Eqn. 1 - T - Eqn. 2 I$ k, kQ

The gas phase mass transfer coefficient, kg, is correlated to experimental conditions and is a strong function of the geometry of the apparatus. The liquid film mass transfer coefficient, kg', quantifies how fast the solution will absorb or desorb CO2.
Results.
Solid solubility. The solid solubility of PZ was studied over a range of PZ
concentration, CO2 loading, and temperature. Solutions were prepared to cover the desired solution properties and were allowed to equilibrate at each condition with stirring before solubility observations were made. The transition temperature of 8 and 10 in PZ solutions over a range of CO2 loading is shown in Figure 13. The transition temperature is the temperature at which a liquid solution will first precipitate when cooled slowly. The approximate temperature ramp for all transitions was 1 C every 5 minutes. The two dashed lines at rich loadings in Figure 13 represent soluble PZ solutions indicating that the solubility envelope extends at least this far. The transition temperature of unloaded PZ
solutions HOU03:1200893.3 ranging from 1.0 to 40 m PZ is shown in Figure 14 (The Dow Chemical Company, 2001;
Bishnoi, 2000; Hilliard, 2008).
The data from this study shows a eutectic point around 60 wt% PZ that was observed in the other data sources shown as well. For 8 m PZ, a CO2 loading of approximately 0.25 mole CO2 per mole of alkalinity is required to maintain a liquid solution without precipitation at room temperature (20 C). In addition, the solubility of anhydrous PZ at 20 C is 14 wt%
PZ, which corresponds to 1.9 m PZ.
Viscosity. The viscosity of aqueous PZ solutions has been measured from 0.20 to 0.45 mole C02 per mole alkalinity, 2 m PZ to 20 m PZ, and 25 C to 60 C. The viscosity of 8 and 10 m PZ is compared with other amines in Figure 15 (Huntsman Chemical, 2005;
Closmann, 2008). The amine concentration is plotted in units of moles alkalinity per kilogram of water in order to compare mono- and diamines on a similar basis. All of the viscosities shown in Figure 15 are at 40 C and at the rich loading of the system (0.3 mole C02 per mole alkalinity for MDEA and MDEA/PZ blend; 0.4 mole CO2 per mole alkalinity for PZ and DGA;
0.5 mole C02 per mole alkalinity for MEA).
Comparison of the viscosity on this basis shows how the amine basic group affects overall viscosity. As the concentration of basic groups increases in a molecule, the viscosity increases in a linear direction. The viscosity of 8 m PZ is higher than that of 7 m 1VIEA, but as compared to 60 wt % DGA , the viscosity of PZ is lower for a higher alkalinity.
Therefore, PZ has the advantage of having two amine functional groups without suffering an increase in viscosity over DGA . DGA solutions at 60 wt % are successfully used in natural gas treating (Al-Juaied, 2004).
Oxidative degradation. Heavy metals are known to catalyze the oxidative degradation of amines (Goff and Rochelle, 2004). The results of oxidative degradation of concentrated PZ in the presence of several dissolved metals are shown in Table 1. The experiments simulated four scenarios: (1) leaching of stainless steel metals (iron, chromium, and nickel), (2) addition of a copper-based corrosion inhibitor, (3) addition of a vanadium-based corrosion inhibitor (low concentration), and (4) addition of a copper-based corrosion inhibitor and proprietary inhibitor "A".

Oxidative degradation of concentrated PZ was found to be four times slower than that of WA in the presence of stainless steel metals (Fe2+, Cr3+, and Ni2+) and a low HOU03:1200893.3 concentration of vanadium. As with 1VIEA solutions, PZ was determined to be highly susceptible to oxidative degradation in the presence of Cu2+ (Goff and Rochelle, 2006). The primary degradation products were found to be ethylenediamine (EDA), formate, oxalate, and N-formylpiperazine, the amide of formate and PZ (denoted as Formamide in the table). The N-formylpiperazine concentration was not measured directly, but inferred from formate production through the basic reversal of the N-formylpiperazine formation reaction. Also, as with WA, Inhibitor "A" was able to vastly reduce this degradation to levels comparable with the stainless steel and vanadium cases (Goff and Rochelle, 2006).

Table 1: Oxidative Degradation of PZ and WA at 55 C (100 ml per min of 98%
02/2%
C02, 350 mL solution) Case Solution Additives Rate of Formation (mM/hr) (m) (mM) Formate Formamide EDA Amine - 7 MEA 1.0 Fe 0.29 0.35 - -3.8 1 10 PZ 0.6 Fe2+, 0.25 Cr3+, 0.25 Ni2+ 0.005 0.007 0 -1.1 2 10 PZ 4.0 Cu2+ 0.14 0.24 0.43 -3.0 3 8 PZ 0.1 Fe2+, 0.1 V4+ 0.006 0.013 0 -0.8

4 8 PZ 4.0 Cu2+, 0.1 Fe2+, 100 "A" 0.011 0.016 0.009 -1.1 Thermal degradation. Thermal degradation was investigated in PZ solutions at slightly above stripper temperature (135 C) and much higher than stripper temperatures (150 C and 175 C). The thermal degradation results are shown in Table 2 and are reported as the percent of amine lost per week as compared with the initial amine concentration.
Experiments ranged from 4 to 18 weeks in length.
PZ thermal degradation was determined to be negligible at 135 and 150 C as compared to 7 m WA. At 175 C, PZ thermal degradation was observed as a loss of 32% of the initial PZ in 4 weeks. EDA was observed as a thermal degradation product at 175 C but not at lower temperatures. Addition of 5.0 mM Cu2+/0.1 mM Fe2+, 5.0 mM
Cu2+/0.1 mM
Fe2+/100 mM Inhibitor "A", and 0.6 mM Cr3+/0.25 mM Fe2+/0.25 mM Ni2+ did not affect degradation rates at 175 C.

HOU03:1200893.3 Table 2: Comparison of Thermal Degradation for PZ and MEA
Temperature Solvent Loading Amine Loss ( C) vent (mol/mol alkalinity) (% per week) 135 7 m MEA 0.4 5.3 m PZ 0.3 0.25 7 m MEA 0.4 11 150 10 m PZ 0.3 0.80 8 m PZ 0.3 0.44 175 8 m PZ 0.3 8.0

5 C02 solubility. The measured solubility of CO2 in 2 m to 8 m PZ solutions ranging from 40 to 100 C is in given in Figure 16 and compared to previous studies (Dugas, 2008;
Ermatchkov et al., 2006; Hilliard, 2008). The CO2 solubility data for PZ was regressed to yield the solid lines shown at on the figure at the various temperatures indicated. The regression of the data is the equilibrium partial pressure of CO2 in terms of temperature, T, in 10 Kelvin, CO2 loading, a, in mole CO2 per mole alkalinity, and the universal gas constant, R, in kJ per mole-K, as shown in Equation 3.

Lli(P,,,) 36.1 Eqn. 3 The CO2 solubility of concentrated, aqueous PZ solutions follows the trends found previously for lower concentration PZ solutions at 40 and 60 C. CO2 solubility is known to not be a strong function of amine concentration and this is confirmed for high concentration PZ solutions (Hilliard, 2008). At 40 C, 8 m PZ provides a working capacity of 0.73 mole per kg (PZ+H20), which is calculated based on a change in the equilibrium C02 partial pressure from 7.5 kPa (loading of 0.415 mole CO2 per mole alkalinity) to 0.75 kPa (0.33 mole CO2 per mole alkalinity). For 7 m WA at 40 C, the working capacity is 0.43 mole CO2 per kg (MEA+H20) based on a change in the equilibrium partial pressure of C02 from 5 kPa (0.53 mole CO2 per mole alkalinity) to 0.5 kPa (0.45 mole CO2 per mole alkalinity).
The selected range of CO2 loading for the 8 m PZ solution falls within the solubility envelope established in Figures 13 and 14.
Kinetics of C02 absorption in PZ solutions. The kinetics of the CO2 absorption into concentrated aqueous PZ was studied in a wetted wall column. The measured liquid-side HOU03:1200893.3

6 PCT/US2009/045075 mass transfer coefficient based on a gas side driving force, kg', for 8 in PZ
is shown compared to 7 in 1VIEA in Figure 17 for 40, 60, 80, and 100 C (Dugas, 2008).
The rate data at 60, 80 and 100 C are plotted as function of the equilibrium partial pressure of CO2 of the solution at 40 C.
As demonstrated in Figure 17, this normalized flux, kg', for 8 in PZ is 2 to 3 times greater than for 7 in WA. For example, at 40 C and an equilibrium C02 partial pressure of 500 Pa, the kg' for 8 in PZ and 7 in WA are 1.98 x 10-6 and 7.66 x 10-7 mol/s-Pa-m2, respectively. This demonstrates that the kinetic rate of concentrated PZ is over twice as fast as WA at 40 C. The same trend is observed for the data at 60 C. At 80 and 100 C, the performance improvement of PZ over WA is nearly double, although not quite as apparent as the lower temperatures.
Volatility of PZ solutions. The volatility of PZ was measured in an equilibrium cell with hot gas FTIR. The volatility of 8 in PZ solutions is compared to that of 5 in PZ and 7 in MEA in Figure 18. The volatility of each solution is normalized by the amine concentration for comparison purposes.
At 40 C, the normalized volatility of PZ solutions is in the same range as the normalized volatility of WA solutions. It was anticipated that PZ would have a higher volatility than WA because the boiling point of PZ, 146 C, is lower than that of WA, 170 C. However, the volatility of both 5 and 8 in PZ is slightly lower at 40 C. Modeling of PZ systems demonstrates this effect as a greatly reduced activity coefficient for PZ due to the solution's non-ideality (Hilliard, 2008). At 40 C, PZ volatility varies from 7 to 20 ppm at atmospheric pressure.
Estimated energy requirement. The thermodynamic model for PZ developed by Hilliard (2008) was modified to represent the new data for concentrated PZ.
The stripper of a system for C02 removal was simulated for 8 in PZ and compared with 7 in WA.
One set of these simulations included a simple stripper with C02 compression to 15 MPa (150 atm), a 5 C cold side temperature approach for the cross heat exchanger, and a 10 C
approach for the reboiler. The columns were simulated using the AspenPlus RateSep tool that calculated heat and mass transfer rates but assumed reactions reached equilibrium. In each simulation, 15 meters of CMR NO-2P packing and an 80% approach to flood were used.

HOU03:1200893.3 A second set of simulations was performed in AspenPlus using two and three stage flash configurations. The flowsheet for the three stage flash is shown in Figure 19. The two stage flash is analogous with one less flash tank. In a multi-stage flash, hot, rich amine leaving the cross exchanger enters a series of flash tanks that are either heated or adiabatic.
The figure shows the design used for these simulations, where each stage is shown heated with steam. In each tank, C02 flashes off and is sent to a multi-stage compressor. A multi-stage flash collects C02 at multiple pressure levels, therefore reducing compression work.
There is a potential opportunity for heat recovery from the water vapor leaving each of the flash tanks. One option is to use this heat to pre-heat the boiler feed water used in the coal-fired power plant (Gibbins and Crane, 2004).
The rich stream for each case assumed a P*C02 of 5 kPa at the absorber temperature of 40 C. Equivalent work, Weq, is calculated as shown in equation 3 using the C02 removal rate, nCO2, stripper reboiler duty, Q, reboiler temperature, Treboiler, cooling water temperature of 40 C, Tsink, total pumping work, Wpump, and total CO2 compression work to achieve 15 MPa, Wc mp.

7~ Egn.3 Each system was optimized for lean loading and the equivalent work as a function of lean loading as shown in Figure 20. The baseline system, 7 m MEA, had an equivalent work of 40.3 kJ per mole C02. The 8 m PZ simple stripper system had a minimum equivalent work of 36.5 kJ per mole C02. The two and three stage flashes using 8 m PZ had minimum equivalent works of 34.1 and 33.8 kJ per mole C02, respectively.
The increased capacity of PZ improved its performance in all cases over the baseline 7 m WA case, despite a lower AHabs. For the PZ cases, the lowest equivalent work was achieved in the three stage flash simulation, demonstrating the advantages of multistage compression and heat recovery that can be achieved using a solvent that is resistant to thermal degradation.
Degradation of Concentrated Piperazine in Pilot Plant.
A long term thermal degradation experiment demonstrated the thermal resistance of concentrated PZ. After 18 weeks at 150 C, only 8.0% of the initial PZ was lost. This HOU03:1200893.3 amounts to a loss of only 0.44% of the original PZ per week. The most prevalent degradation products were EDA (1.2 mM/wk), formate (0.9 mM /wk), and N-formyl amides (2.3 mM/wk). As demonstrated by the low weekly loss of PZ, 8 m PZ has enhanced resistance to thermal degradation as compared with 1VIEA, DGA, and MDEA/PZ blends.
Table 3: Comparison of Thermal Degradation Rates of Amines Temperature Solvent System CO2 Loading Amine Loss ( C) (mol/mol alkalinity) (/o/week) 7 m MEA 0.4 6.0 135 7 m MDEA/2 m PZ 0.1 3.7 7 m DGA 0.4 1.8 m PZ 0.3 0.3 7 m MEA 0.4 11 150 7 m MDEA/2 m PZ 0.1 6.4 10 m PZ 0.3 0.80

8 m PZ 0.3 0.44 175 8 m PZ 0.3 8.0 Conclusions.
Concentrated, aqueous solutions of PZ have shown promise for improved solvent performance in absorption/stripping systems used for CO2 capture. For 8 m PZ, a CO2 10 loading of approximately 0.25 mole CO2 per mole alkalinity is required to maintain a liquid solution without precipitation at room temperature (20 C). Additionally, the solubility of PZ
at 20 C is approximately 14 wt% PZ, or 1.9 m PZ. The volatility of 8 m PZ
systems was found to be between 7.3 and 20.2 ppm PZ at 40 C, which is comparable to 7 m WA
solutions.
Oxidative degradation of concentrated PZ has been shown to be four times slower than 7 m MEA in the presence of the combination of Fe2+/Cr3+/Ni2+ and Fe2+/V4+. In the presence of copper-based corrosion inhibitors, oxidative degradation is an issue but can be drastically reduced with the use of Inhibitor "A". Concentrated PZ is resistant to thermal degradation up to 150 C but does degrade at 175 C, losing 32% of the PZ over 4 weeks. The resistance of PZ to thermal degradation allows for the possibility of higher pressure strippers to improve energy performance.

Kinetic measurements have shown that the rate of CO2 absorption into 8 m PZ is more than twice that of 7 m WA at 40 C and nearly double at 60 C. The working capacity HOU03:1200893.3 of an 8 m PZ solution is 0.73 mole CO2 per kg (PZ + H20), nearly double that of 7 m MEA.
Initial modeling of a simple stripper section indicate that the equivalent work required for stripping of an 8 m PZ solution will be approximately 10-20% lower than that of 7 m MEA.
The use of a multi-stage flash also has demonstrated advantages for a high temperature operation that is feasible with the thermally stable 8 m PZ solution.
The rapid rate of CO2 absorption, low degradation rate, and low predicted equivalent work indicate that 8 m PZ solutions are an attractive option for CO2 capture in absorption/stripping systems.
Therefore, the present invention is well adapted to attain the ends and advantages mentioned as well as those that are inherent therein. While numerous changes may be made by those skilled in the art, such changes are encompassed within the spirit of this invention as illustrated, in part, by the appended claims.
References:
S. Bishnoi, Carbon Dioxide Absorption and Solution Equilibrium in Piperazine Activated Methyldiethanolamine. The University of Texas at Austin, Austin, TX, 2000.
M.D. Hilliard, A Predictive Thermodynamic Model for an Aqueous Blend of Potassium Carbonate, Piperazine, and Monoethanolamine for Carbon Dioxide Capture from Flue Gas. The University of Texas at Austin, Austin, TX, 2008.
J.T. Cullinane and G.T. Rochelle, "Thermodynamics of aqueous potassium carbonate, piperazine, and carbon dioxide." Fluid Phase Equilibria. 227(2) (2005) 197-213.
A. Sexton, "Catalysts and inhibitors for MEA oxidation." Presentation at GHGT-

9, Washington D.C., 2008.
J. Davis, "Thermal degradation of monoethanolamine at stripper conditions."
Presentation at GHGT-9, Washington D.C., 2008.
R. Dugas, "Absorption and desorption rates of carbon dioxide with monoethanolamine and piperazine." Presentation at GHGT-9, Washington D.C., 2008.
Brochure, Dow Chemical Company, Ethyleneamines; August, 2001 p 48.
Brochure, Diglycolamine Agent - Product Information, Diglycolamine Agent -Product Information; 2005 p 60.
F. Closmann, "MDEA/piperazine as a solvent for CO2 capture." Presentation at GHGT-9, Washington D.C., 2008.

HOU03:1200893.3 M.A. Al-Juaied, Carbon Dioxide Removal from Natural Gas by Membranes in the Presence of Heavy Hydrocarbons and by Aqueous Diglycolamine%Morpholine. The University of Texas at Austin, Austin, TX, 2002.
G.S. Goff and G.T. Rochelle, "Monoethanolamine degradation: 02 mass transfer effects under CO2 capture conditions." Ind. Eng. Chem. Res. 43(20) (2004) 6400-6408.
G.S. Goff and G.T. Rochelle, "Oxidation inhibitors for copper and iron catalyzed degradation of monoethanolamine in CO2 capture processes." Ind. Eng. Chem.
Res. 45(8) (2006) 2513-2521.
V. Ermatchkov, A.P.S. Kamps, D. Speyer, and G. Maurer, "Solubility of carbon dioxide in aqueous solutions of piperazine in the low gas loading region." J
Chem. Eng. Data 51(5) (2006) 1788-1796.

HOU03:1200893.3

Claims (17)

What is claimed is:

1. A low-volatility aqueous composition consisting essentially of a thermally stable amine from 4 to 12 moles/kg of water loaded with carbon dioxide in an amount from 0.25 to 0.45 moles carbon dioxide per amine equivalent of the thermally stable amine, wherein the composition is at a temperature of between 120°C and 160°C.

2. The composition of claim 1 wherein the thermally stable amine is piperazine, or methyldiethanolamine, or wherein the thermally stable amine is piperazine, methylpiperazine, dimethylpiperazine, ethylpiperazine, diethylpiperazine, morpholine, 5-amino-1-pentanol, 2-amino-2-methyl-1-propanol , diglycolamine, 4-amino-1-butanol, 3-amino-1-propanol, hydroxyethylpiperazine, 1-amino-2-propanol, 2-amino-1-propanol or mixtures thereof.

3. A method comprising contacting a gas mixture having an acidic gas with a solvent, wherein the solvent comprises a thermally stable amine in an amount from 4 to 12 moles/kg of water loaded with carbon dioxide in an amount from 0.25 to 0.45 moles carbon dioxide per amine equivalent of the thermally stable amine,wherein the gas mixture is in the form of a gaseous stream and wherein the acid gas is allowed to transfer from the gaseous stream to the solvent, further comprising forming a purified gaseous stream and a rich solvent stream, and routing the rich solvent stream through a stripper, wherein the stripper is operated at a temperature between 120°C and 160°C.

4. The method of claim 3 wherein the stripper is a simple stripper, matrix stripper, a multistage flash stripper, an exchange stripper, a multipressure stripper, a flashing feed stripper, a multistage stripper or mixtures thereof.

5. The method of claim 3 or 4 wherein the gaseous stream is a flue gas, a natural gas, a hydrogen gas, a synthesis gas or mixtures thereof.

6. The method of claim 3 or 4 wherein the acidic gas is CO2, H2S, COS, CS2, mercaptans or mixtures thereof.

7. The method of claim 3 or 4 further comprising recycling a solvent stream exiting the stripper.

8. The method of claim 3 wherein the stripper is a multistage flash stripper operated at a temperature of 120 to 160°C.

9. A method according to claim 3, wherein the contact between the said gas mixture and said solvent occurs in an absorber, further comprising allowing the acid gas to transfer from the gaseous stream to the solvent; forming a purified gaseous stream and a rich solvent stream; routing the rich solvent stream through a multistage flash stripper or a multistage stripper.

10. A method for reducing the volatility of a thermally stable amine comprising providing a solvent that comprises a thermally stable amine in an amount from 4 to 12 moles/kg of water and adding carbon dioxide to the solvent in an amount from 0.25 to 0.45 moles per amine equivalent of thermally stable amine wherein the solvent is subjected to a temperature between 120°C and 160°C.

11. The method of claim 10 wherein carbamate ions form and reduce the volatility.

12. A method for increasing the solubility of a solid thermally stable amine comprising providing a thermally stable amine to a solution consisting essentially of 0.25 to 0.45 moles carbon dioxide per amine equivalent of the thermally stable amine, wherein the thermally stable amine is from 4 to 12 moles/kg of water.

13. The method of claim 12 wherein carbamate ions form.

14. Use of a low-volatility aqueous composition consisting essentially of a thermally stable amine from 4 to 12 moles/kg of water loaded with carbon dioxide in an amount from 0.25 to 0.45 moles carbon dioxide per amine equivalent of the thermally stable amine, in a stripper operated at a temperature between 120°C and 160°C.

15. The use of claim 14, wherein the carbon dioxide is in an amount of 0.3 to 0.4 moles.

2)

16. The method of claim 3, 10 or 12, wherein the thermally stable amine is piperazine, or methyldiethanolamine, or wherein the thermally stable amine is piperazine, methylpiperazine, dimethylpiperazine, ethylpiperazine, diethylpiperazine, morpholine, 5-amino-1-pentanol, 2-amino-2-methyl-1-propanol , diglycolamine, 4-amino-1-butanol, 3-amino-1-propanol, hydroxyethylpiperazine, 1-amino-2-propanol, 2-amino-1-propanol or mixtures thereof.

17. The use of claim 14 or 15, wherein the thermally stable amine is piperazine, or methyldiethanolamine, or wherein the thermally stable amine is piperazine, methylpiperazine, dimethylpiperazine, ethylpiperazine, diethylpiperazine, morpholine, 5-amino-1-pentanol, 2-amino-2-methyl-1-propanol , diglycolamine, 4-amino-1-butanol, 3-amino-1-propanol, hydroxyethylpiperazine, 1-amino-2-propanol, 2-amino-1-propanol or mixtures thereof.