Patent Application
20150027055
This invention relates to a process for removing acid gases from natural gas and other gas streams at high pressure. In particular, it relates to a process for selectively removing hydrogen sulfide from these gas mixtures in the presence of carbon dioxide.
A number of different technologies are available for removing acid gases such as carbon dioxide, hydrogen sulfide, carbonyl sulfide. These processes include, for example, chemical absorption (amine), physical absorption, cryogenic distillation (Ryan Holmes process), and membrane system separation. Of these, amine separation is a highly developed technology with a number of competing processes in hand using various amine sorbents such as monoethanolamine (MEA), diethanolamine (DEA), triethanolamine (TEA), methyldiethanolamine (MDEA), diisopropylamine (DIPA), diglycolamine (DGA), 2-amino-2-methyl-1-propanol (AMP) and piperazine (PZ). Of these, MEA, DEA, and MDEA are the ones most commonly used. The amine purification process usually contacts the gas mixture countercurrently with an aqueous solution of the amine in an absorber tower. The liquid amine stream is then regenerated by desorption of the absorbed gases in a separate tower with the regenerated amine and the desorbed gases leaving the tower as separate streams. The various gas purification processes which are available are described, for example, in Gas Purification, Fifth Ed., Kohl and Neilsen, Gulf Publishing Company, 1997, ISBN-13: 978-0-88415-220-0.
It is often necessary or desirable to treat acid gas mixtures containing both CO2 and H2S so as to remove the H2S selectively from the mixture while minimizing removal of the CO2. While removal of CO2 may be necessary to avoid corrosion problems and provide the required heating value to the consumer, selective H2S removal may be necessary or desirable. Natural gas pipeline specifications, for example, set more stringent limits on the H2S level than on the CO2 since the H2S is more toxic and corrosive than CO2: common carrier natural gas pipeline specifications typically limit the H2S content to 4 ppmv with a more lenient limitation on the CO2 at 2 vol %. Selective removal of the H2S may enable a more economical treatment plant to be used and selective H2S removal is often desirable to enrich the H2S level in the feed to a sulfur recovery unit.
The reaction kinetics with hindered amine sorbents allow H2S to react more rapidly with the amine groups of the sorbent to form a hydrosulfide salt in aqueous solution but under conditions of extended gas-liquid contact where equilibrium of the absorbed sulfidic species with CO2 is approached, carbon dioxide can displace hydrogen sulfide from the previously absorbed hydrosulfide salt since carbon dioxide is a slightly stronger acid in aqueous solution than hydrogen sulfide (ionization constant for the first ionization step to H+ and HCO3− is approximately 4×10−7 at 25° C. compared to 1×10−7 for the corresponding hydrogen sulfide ionization) so that under near equilibrium conditions, selective H2S removal becomes problematical, presenting a risk of excessive H2S levels in the effluent product gas stream.
An improvement in the basic amine process involves the use of sterically hindered amines. U.S. Pat. No. 4,112,052, for example, describes the use of hindered amines for nearly complete removal of acid gases including CO2 and H2S. U.S. Pat. Nos. 4,405,581; 4,405,583; 4,405,585 and 4,471,138 disclose the use of severely sterically hindered amine compounds for the selective removal of H2S in the presence of CO2. Compared to aqueous MDEA, severely sterically hindered amines lead to much higher selectivity at high H2S loadings. Amines described in these patents include BTEE (bis(tertiary-butylamino)-ethoxy-ethane synthesized from tertiary-butylamine and bis-(2-chloroethoxy)-ethane as well as EEETB (ethoxyethoxyethanol-tertiary-butylamine) synthesized from tertiary-butylamine and chloroethoxyethoxyethanol). U.S. Pat. No. 4,894,178 indicates that a mixture of BTEE and EEETB is particularly effective for the selective separation of H2S from CO2. U.S. 2010/0037775 describes the preparation of alkoxy-substituted etheramines as selective sorbents for separating H2S from CO2.
The use of hydroxyl-substituted amines (alkanolamines) such as those mentioned above has become common since the presence of the hydroxyl groups tends to improve the solubility of the absorbent/acid gas reaction products in the aqueous solvent systems widely used, so facilitating circulation of the solvent through the conventional absorber tower/regeneration tower unit. This preference may, however, present its own problems in certain circumstances. A current business driver is to reduce the cost to regenerate and to recompress acid gases prior to sequestration. For natural gas systems, the separation of the acid gases can occur at pressures of about 4,800-15,000 kPaa (about 700-2,200 psia), more typically from about 7,250-8,250 kPaa (about 1050-1200 psia). While the alkanolamines will effectively remove acid gases at these pressures, the selectivity for H2S removal can be expected to decrease markedly both by direct physisorption of the CO2 in the liquid solvent and by reaction with the hydroxyl groups on the amine compound. Although the CO2 reacts preferentially with the amino nitrogen, higher pressures force reaction with the oxygens and under the higher pressures, the bicarbonate/hemicarbonate/carbonate reaction product formed by the reaction at the hydroxyl site is stabilized with a progressive loss in H2S selectivity with increasing pressure. This effect can be perceived, for example, with MDEA (N-methyl diethanolamine). For example, 5M MDEA in aqueous solution does not absorb carbon dioxide under ambient conditions, but will form a hydrosulfide salt at the nitrogen. However, H2S/CO2 selectivity significantly reduces at high CO2 pressure presumably due to O-carbonation of hydroxyl groups:
A similar trend is observed with the secondary aminoether, ethoxyethoxyethanol-t-butylamine (EEETB): at low pressures, this absorbent offers H2S selectivity over CO2 based on a faster reaction with the hindered secondary amine group although a significant amount of CO2 can be absorbed by the hydroxyl group which has low affinity to H2S. At higher pressures, however, the reaction yield of O-carbonation increases, suppressing the H2S/CO2 selectivity achieved by the hindered secondary amine:
There is therefore a need for an alkanolamine absorbent system that can selectively absorb H2S from gas mixtures that also contain CO2 and that can be regenerated at high pressure and relatively low temperature while maintaining very low CO2 solubility. This can significantly reduce the cost and energy required for regeneration and recompression as well as improving operation of the sulfur recovery plant.
We have now found that it is possible to achieve improved selectivity for the removal of H2S from gas mixtures also containing CO2 at high(er) pressures by the use of capped alkanolamines and more basic sterically hindered secondary and tertiary amines as the absorbent. This effect is particularly useful when treating natural gas streams where reinjection of the carbon dioxide into the subterranean producing formation is to be carried out since CO2 recompression costs are reduced with the separation being operated at the higher pressures requisite for injection back into the formation.
According to the present invention therefore, a process for increasing the selectivity of an alkanolamine/amine absorption process for H2S absorption) from a gas mixture which also contains carbon dioxide (CO2) and possibly other acidic gases such as COS, HCN, CS2 and sulfur derivatives of C1 to C4 hydrocarbons, comprises contacting the gas mixture with a liquid absorbent which is a severely sterically hindered capped alkanolamine or a more basic sterically hindered secondary and tertiary amine; the contacting and regeneration of the absorbent is carried out at high(er) pressure, preferably, at a pressure of at least about 10 bara (about 147 psia) so that selectivity for removal of the H2S relative to the CO2 removal is achieved at a level above that prevailing under ambient pressure (about 1 bara, 14.7 psia). The selectivity described here signifies that the present process is capable of removing H2S in preference to the CO2, that is, the molar proportion of absorbed H2S is greater than the molar proportion of absorbed CO2. This H2S selectivity is achieved according to the kinetics of the respective absorption mechanisms by appropriate control of process conditions notably, the contact time between the gas stream and the liquid absorbent as discussed below.
In its typical mode of application the amine separation process for a natural gas stream which contains both H2S and CO2 achieves improved selectivity for H2S separation (relative to CO2 separation) under higher pressure conditions. It functions by:
When the absorbent is a capped alkanolamine, that is, an alkanolamine in which one or more of the hydroxyl groups have been capped or converted into ether groups, exemplary amine absorbents of this type include, for example, the following:
N-(2-methoxyethyl)-N-methyl-ethanolamine (MDEA-(OMe),
Bis-(2-methoxyethyl)-N-methylamine (MDEA-(OMe)2),
2-amino-prop-1-yl methyl ether (AP-OMe),
2-methyl-2-amino-prop-1-yl methyl ether (AMP-OMe),
2-N-methylamino-prop-1-yl methyl ether (MAP-OMe),
2-N-methylamino-2-methyl-prop-1-yl methyl ether (MAMP-OMe),
2-N-ethylamino-2-dimethyl-prop-1-yl methyl ether, (EAMP-OMe),
2-(N,N-dimethylamino)-ethyl methyl ether (DMAE-OMe),
Methoxyethoxyethoxyethanol-t-butylamine (MEEETB).
When the absorbent is a more basic sterically hindered secondary and tertiary amine, preferred structures include guanidines, amidines, biguanides, piperidines, piperazines, and the like. Tetramethyguanidine, pentamethylguanidine, 1,4-dimethylpiperazine, 1-methylpiperidine, 2-methylpiperidine, 2,6-dimethylpiperidine are examples.
In the accompanying drawings:
The present selective gas separation process is particularly apt for use in the treatment of natural gas which is normally compressed subsequent to gathering from the wellheads for treatment prior to pipelining. Interstate gas transmission lines are usually operated at pressures above 15 bara (about 220 psia) and in most cases in the range of 15 to 100 bara (about 217 to 1450 psia) for economy in transmission by reduction of gas volume. At pressures of this magnitude, the stability and capacity of the H2S/absorbent reaction products is markedly increased as the effect of the pressure is to move equilibrium to the right in the sorption reaction:
R1—O—R2—NHR3+H2S→R1—O—R2—NH2+R3HS−
where R1, R2 and R3 are the groups, usually alkyl or alkylene in the absorbent molecule as described below. The carbonation of the hydroxyl group(s) is no longer permitted by the capping group so that selectivity under these pressure conditions is notably enhanced. At the same time, the regenerability of the absorbent is improved. The absorbed H2S may be released from the hydrosulfide salt formed by reaction at the amino nitrogen amine by a reduction in pressure at a relatively low temperature; significantly lower than the regeneration temperatures conventionally used above about 90° C.; desorption temperatures of from about 40 to 70° C. become usable, with a considerable savings in the energy required in the overall sorption-desorption process. Alternatively, with the stability of the H2S/amine strongly dependent on pressure, the separation process may be operated on a pressure swing cycle with a reduction in pressure to desorb the H2S and regenerate, or partially regenerate the capped amine absorbent.
The separation process may be carried out in a cyclic liquid sorbent gas separation unit as illustrated in
The absorbent solution, which is liberated from most of the absorbed gas while flowing downward through regenerator 12, exits through line 18 at the bottom of the regenerator for transfer to a reboiler 19. Reboiler 19, equipped with an external source of heat (e.g., steam injected through line 20 and the condensate exits through a second line (not shown)), vaporizes a portion of this solution (mainly water) to force the release of more H2S. The H2S and steam driven off are returned via line 21 to the lower section of regenerator 12 and exit through line 13 for entry into the condensation stages of gas treatment. The solution remaining in the reboiler 19, referred to as the “lean” solution, is drawn through line 22, cooled in heat exchanger 9, and introduced by the action of pump 23 (optional if pressure is sufficiently high) through line 5 into the absorber column 2 for re-use.
The stability of the absorbed species generally decreases with increasing temperature so that absorption of the H2S will favored by lower temperatures. With natural gas streams, the temperature will usually be low enough to favor absorption, particularly if the gas has been passed through an expansion before entering the unit. The absorption temperature will typically be at least 10° C. and in most cases at least 15 to 20° C. with the most typical range being about 25° C. to 30° C.; the upper limit on absorption temperature will not normally extend above about 90° C. and will normally not exceed about 50 to 75° C. In most cases, however, a maximum temperature for the sorption will be 75° C. and if operation is feasible at a lower temperature, e.g., with a chilled incoming natural gas, resort may be advantageously made to lower temperatures at this point in the cycle.
The sorption solution may include a variety of additives typically employed in selective gas removal processes, e.g., antifoaming agents, anti-oxidants, corrosion inhibitors. The amount of these additives will typically be in the range that they are effective.
As will be apparent from the following discussion, the selective character of the present absorption process in which the H2S is preferentially absorbed by capped primary and secondary alkanolamines is achieved by the absorption kinetics which initially favor the reaction with the H2S although this reaction is less thermodynamically favored; continued exposure to the carbon dioxide permits displacement of the initial hydrosulfide kinetic reaction product by a carbonate/bicarbonate reaction product formed with the CO2. During the sorption step, therefore, the kinetics favoring H2S absorption are exploited by limiting mass transfer and using short contact times so that the incoming gas mixture does not remain in contact with the absorbent for the CO2 to substantially displace the absorbed H2S. The mass transfer zone designed correctly and the contact time between the incoming gas stream and the absorbent should therefore be monitored and controlled (i.e., of alternate amine inlets) so as to take advantage of the kinetics favoring H2S sorption over the CO2 reaction. Contact times less than 5 minutes and preferably less than 1 minute are effective with H2S selectivity increasing with shorter contact times since opportunities for displacement of absorbed sulfidic species by CO2 are correspondingly reduced. Flow rates in the cyclic operation should therefore be controlled accordingly.
For absorption, the temperature is typically in the range of from about 25° C. to about 90° C., preferably from about 20° C. to about 75° C.; the stability of the H2S/amine species generally decreases with increasing temperature. In most cases, however, a maximum temperature for the sorption will be 75° C. and if operation is feasible at a lower temperature, e.g., with a chilled incoming natural gas or refinery process stream, resort may be advantageously made to lower temperatures at this point in the cycle. Temperatures below 50° C. are likely to be favored for optimal sorption and selectivity.
The minimum pressure is typically about 1.0 bar (absolute) e.g. 1.1 bara, and often above this value, e.g. 10 bara to 15 bara, depending on the handling of the gas stream prior to entering the separation unit. Maximum pressures will not normally exceed about 150 bara and again will vary according to the previous handling of the gas, and in most cases not more than 100 bara or even lower, e.g., 70 bara, 50 bara, 40 bara, 30 bara or 20 bara. The partial pressures of hydrogen sulfide and carbon dioxide in the gas mixture will vary according to the gas composition and the pressure of operation. The gas mixture can be contacted counter currently or co-currently with the absorbent material at a typical gas hourly space velocity (GHSV) of from about 50 (S.T.P.)/hour to about 50,000 (S.T.P.)/hour with the higher velocities favored with aqueous solutions as noted above to disfavor displacement of absorbed H2S by CO2 with longer contact times.
The H2S can be desorbed from the absorbent material by conventional methods. One possibility is to desorb the absorbed H2S by means of stripping with an inert (non-reactive) gas stream such as nitrogen in the regeneration tower. The reduction in the H2S partial pressure which occurs on stripping promotes desorption of the H2S and when this expedient is used, there is no requirement for a significant pressure reduction although the pressure may be reduced for optimal stripping, suitably to the levels used in pressure swing operation.
When carrying out the desorption by inert gas sparging or pressure swing operation, the temperature may be maintained at a value at or close to that used in the sorption step. Desorption, will however, be favored by an increase in temperature, either with or without stripping or a decrease in pressure.
The H2S can be desorbed from the absorbent material by conventional methods including temperature swing, pressure swing and stripping with an inert (non-reactive) gas stream such as nitrogen, CO2, or steam in the regeneration tower. Temperature swing operation is often a choice in conventional cyclic absorption plants. The temperature of the rich solution from the absorption zone is raised in the regeneration tower, e.g., by passage through a heat exchanger at the bottom of the regeneration tower or with steam or other hot gas. Desorption temperatures will be dependent on the vapor/liquid equilibria for the selected system, e.g. alkanolamine, H2S concentration, and will typically be 10° C. or more, and in most cases 15 to 50° C. above the temperature in the absorption zone. Typical temperatures in the regeneration zone will be, for example, from a temperature higher than the temperature of the absorption zone and usually at a temperature from 65 to 100° C.; temperatures above 100° C. are not favored with aqueous systems from the viewpoint of energy consumption as a result of the vaporization of the water in the solvent. Higher temperatures above 100° C. may, however, be used if necessary, for example, to ensure desorption or to drive off any accumulated water from a non-aqueous system; when the preferred regeneration temperature is above 100° C., temperatures up to 120° C. are typically used although temperatures above 120° C. may be preferable to desorb the H2S product at the higher pressures characteristic of this operation. Thermal desorption by passing the rich solution through a hot bath with a head space at controlled pressure (typically above 10 bar) can be a preferred option. Pressure control can be effected by removal of the desorbed gas at an appropriate rate. Pressure swing absorption is likely to be less favored in view of the need for recompression; the pressure drop will be determined by the vapor-liquid equilibria at different pressures.
A slip stream of CO2 may be used for stripping although this may lead to undesirable CO2 remaining in the lean gas stream to the absorption zone although desorption can be favored by heating the CO2 stripping gas. Stripping with steam or an inert (non-reactive) gas is therefore preferred. When carrying out the desorption by inert gas sparging or pressure swing operation, the temperature may be maintained at a value at or close to that used in the sorption step although desorption will be favored by an increase in temperature from the absorption zone to the regeneration zone, either with or without stripping or a decrease in pressure.
In addition to the benefit of improved H2S selectivity with non-aqueous systems, there are other potential advantages in the regeneration of H2S-rich amine streams in non-aqueous systems. In the non-aqueous environment, stripping can be feasible with or without purge gas at relatively lower temperatures. The possibility of desorption at lower temperatures offers the potential for isothermal or near isothermal stripping using a purge gas at a temperature the same as or not much higher than the sorption temperature, for example, at a temperature not more than 30° C. higher than the sorption temperature; in favorable cases, it may be possible to attain a sorption/desorption temperature differential of no more than 20° C. When these factors are taken into consideration the temperature selected for the desorption will typically be in the range of from about 70 to about 120° C., preferably from about 70 to about 100° C., and more preferably no greater than about 90° C.
In non-aqueous systems with water present in the stream to be processed, regeneration may need to be performed at a temperature sufficient to remove the water and prevent build-up in the scrubbing loop. In such a situation, the H2S may be removed at pressures below atmospheric pressure, but above 100° C. For example, the regeneration temperature may be around 90° C., but to remove any water in the sorbent, temperatures in the range of 100 to 120° C. may be required.
For regeneration in non-aqueous systems, stripping with an inert (non-reactive) gas such as nitrogen or a natural gas stream is preferred. Staged heat exchanger systems with intermediate knock out drums in which H2S/water is removed as a pressurized gas stream may be used as one alternative.
Given that the kinetics of the process favor preferential H2S selectivity with short contact times, the present hindered alkanolamine absorbents or more basic sterically hindered secondary and tertiary amine absorbents may advantageously be operated in the kinetic separation mode using the capped alkanolamines as adsorbents in a thin layer on a solid support. Kinetically based separation processes may be operated, as noted in US 2008/0282884, as pressure swing adsorption (PDA), temperature swing adsorption (TSA), partial pressure swing or displacement purge adsorption (PPSA) or as hybrid processes, as noted in U.S. Pat. No. 7,645,324 (Rode/Xebec). These swing adsorption processes can be conducted with rapid cycles, in which case they are referred to as rapid cycle thermal swing adsorption (RCTSA), rapid cycle pressure swing adsorption (RCPSA), and rapid cycle partial pressure swing or displacement purge adsorption (RCPPSA) technologies, with the term “swing adsorption” taken to include all of these processes and combinations of them.
In the kinetically-controlled PSA processes, the adsorption and desorption are more typically caused by cyclic pressure variation, whereas in the case of TSA, PPSA and hybrid processes, adsorption and desorption may be caused by cyclic variations in temperature, partial pressure, or combinations of pressure, temperature and partial pressure, respectively. In the exemplary case of PSA, kinetic-controlled selectivity may be determined primarily by micropore mass transfer resistance (e.g. diffusion within adsorbent particles or crystals) and/or by surface resistance (e.g. narrowed micropore entrances). For successful operation of the process, a relatively and usefully large working uptake (e.g. the amount adsorbed and desorbed during each cycle) of the first component and a relatively small working uptake of the second component may preferably be achieved. Hence, the kinetic-controlled PSA process requires operation at a suitable cyclic frequency, balancing the avoidance of excessively high cycle frequency where the first component cannot achieve a useful working uptake with excessively low frequency where both components approach equilibrium adsorption values.
The faster the beds perform the steps required to complete a cycle, the smaller the beds can be when used to process a given hourly feed gas flow. Several other approaches to reducing cycle time in PSA processes have emerged which use rotary valve technologies as disclosed in U.S. Pat. Nos. 4,801,308; 4,816,121; 4,968,329; 5,082,473; 5,256,172; 6,051,050; 6,063,161; 6,406,523; 6,629,525; 6,651,658 and 6,691,702. A parallel channel (or parallel passage) contactor with a structured adsorbent may be used to allow for efficient mass transfer in these rapid cycle pressure swing adsorption processes. Approaches to constructing parallel passage contactors with structured adsorbents have been disclosed in US20060169142 A1, US20060048648 A1, WO2006074343 A2, WO2006017940 A1, WO2005070518 A1, and WO2005032694 A1.
The use of the hindered capped alkanolamine or more basic sterically hindered secondary and tertiary amines in the form of a film of controlled thickness on the surface of a core which has a low permeability has significant advantages in rapid cycle processes with cycle durations typically less than one minute and often rather less. By using a thin film, heat accumulation and retention is reduced so that exotherms and hot spots in the absorbent bed are minimized and the need for heat sinks such as the aluminum spheres common in conventional beds can be eliminated by suitable choice of the core material; rapid cycling is facilitated by the fast release of heat from the surface coating and the relatively thin layer proximate the surface of the core. A further advantage is secured by the use of low permeability (substantially non-porous) cores which is that largely inhibit entry of the gas into the interior pore structure of the core material is largely inhibited and so that mass and heat transfer takes place more readily in the thin surface layer; and retention of gas within the pore structure is minimized.
Selectivity for H2S sorption will be diminished to a certain extent not only by the relative adsorption characteristics of the selected adsorbent material but also by the physical sorption of CO2 in both liquid and solid systems which becomes more perceivable at higher pressures: the lower the partial pressures of both H2S and CO2, the greater will be the selectivity for H2S. To operate using the capped alkanolamine or more basic sterically hindered secondary and tertiary amine in the solid phase as an adsorbent, the compound is physically or chemically taken up onto on a solid support or carrier material of high surface area. If the basic compound is a solid, it may be dissolved to form a solution which can then be used to impregnate or react with the support material or deposited on it in the form of a thin, wash coat layer of discrete sorbent particles or agglomerates of sorbent particles adhered to the surface of the support. Discrete particles or agglomerates may be adhered effectively by physical interaction at the surface of the support. Porous support materials are generally preferred in view of the greater surface area which they present for the sorption reaction but finely-divided non-porous solids with a sufficiently large surface area may also be used. In either case, the sorbent compound(s) may be physisorbed onto the support material or held onto the surface of the support in the form of a thin, adherent surface layer firmly bonded to the support by physical interaction or alternatively grafted onto the support by chemical reaction.
Porous support materials are frequently used for the catalysts in catalytic processes such as hydrogenation, hydrotreating, hydrodewaxing etc and similar materials may be used for the present sorbents. Common support materials include carbon (activated charcoal) as well as porous solid oxides of metals and metalloids and mixed oxides, including alumina, silica, silica-alumina, magnesia and zeolites. Porous solid polymeric materials are also suitable provided that they are resistant to the environment in which the sorption reaction is conducted. As the components of the gas stream have relatively small molecular dimensions, the minimum pore size of the support is not in itself a severely limiting factor but when the basic nitrogenous compound is impregnated, the entrances to the pore systems of small and intermediate pore size zeolites such as zeolite 4A, erionite, ZSM-5 and ZSM-11 may become occluded by the bulky amine component and for this reason, the smaller pore materials are not preferred, especially with the bases of relatively larger molecular dimensions. Large pore size zeolites with 12-membered ring systems such as ZSM-4, faujasites such as zeolite X and the variants of zeolite Y including Y, REY and USY, may, however, be suitable depending on the dimensions of the basic nitrogenous compound. Amorphous porous solids with a range of different pore sizes are likely to be suitable since at least some of the pores will have openings large enough to accept the basic component and then to leave sufficient access to the components of the gas stream. Supports containing highly acidic reaction sites as with the more highly active zeolites are more likely to be more susceptible to fouling reactions upon reaction with the amino compound and less acidic or non-acidic species are therefore preferred.
A preferred class of solid oxide support is constituted by the mesoporous and macroporous silica materials such as the silica compounds of the M41S series, including MCM-41 (hexagonal) and MCM-48 (cubic) and other mesoporous materials such as SBA-1, SBA-2, SBA-3 and SBA-15 as well as the KIT series of mesoporous materials such as KIT-1. Macroporous silicas and other oxide supports such as the commercial macroporous silicas available as Davisil products are also suitable, e.g. Davisil 634 (6 nm pore size, 480 m2/g pore volume), Davisil 635 (6 nm, 480 m2/g) and Davisil 644 (15 nm, 300 m2/g). According to the IUPAC definition, mesoporous materials are those having a pore size of 2 to 50 nm and the macroporous, those having a pore size of over 50 nm. According to the IUPAC, a mesoporous material can be disordered or ordered in a mesostructure. The preferred mesoporous and macroporous support materials are characterized by a BET surface area of at least 300 and preferably at least 500 m2/g prior to treatment with the base compound. The M41S materials and their synthesis are described in a number of patents of Mobil Oil Corporation including U.S. Pat. Nos. 5,102,643; 5,057,296; 5,098,684 and 5,108,725, to which reference is made for a description of them. They are also described in the literature in “The Discovery of ExxonMobil's M41S Family of Mesoporous Molecular Sieves”, Kresge et al, Studies in Surface Science and Catalysis, 148, Ed. Terasaki, Elsevier bV 2004. SBA-15 is described in “Triblock Copolymer Syntheses of Mesoporous Silica with Periodic 50 to 300 Angstrom Pores”, Dongyuan Zhao, et al. (1998). Science 279 (279). KIT-1 is described in U.S. Pat. No. 5,958,368 and other members of the KIT series are known, for example KIT-5 and KIT-6 (see, e.g. KIT-6 Nanoscale Res Lett. 2009 November; 4(11): 1303-1308). The H2S/CO2 selectivity of the material can be adjusted by the judicious choice of the porous support structure, affording a significant potential for tailoring the selectivity of the adsorbent.
The capped alkanolamine or more basic sterically hindered secondary and tertiary amines may simply be physically absorbed on the support material e.g., by impregnation or bonded with or grafted onto it by chemical reaction with the base itself or a precursor or derivative in which a substituent group provides the site for reaction with the support material in order to anchor the sorbent species onto the support. Chemical bonding is not, however, required for an effective solid phase sorbent material; effective sorbents may be formed by physical interaction when the sorbent is itself strongly adsorbed by the support material. Chemical bonding may be effected by the use of support materials which contain reactive surface groups such as the silanol groups found on zeolites and the M41S silica oxides which are capable or reacting with a silylated derivative of the selected amine compound. The high concentrations of surface silanol groups (SiOH), on silica and ordered siliceous materials such as the zeolites and mesoporous materials, e.g. MCM-41, MCM-48, SBA-15 and related structures, render these materials amenable to surface modification by grafting of the functional amine onto the pore walls of the siliceous support via a reaction between the surface silanol groups of the support and the grafting material according to the conventional technique; see, for example, Huang et al., Ind. Eng. Chem. Res., 2003, 42 (12), 2427-2433. The alkoxy groups e.g., methoxy, ethoxy, present in the alkoxy-capped alkanolamines will be capable of reacting with the —OH groups on the surface of the siliceous material with the release of methanol or ethanol to yield a final grafted structure on the surface of the support with grafting taking place through one or more of the alkoxy groups on the capped alkanolamines.
An alternative method of fixing more volatile adsorbing species on the support is by first impregnating the species into the pores of the support and then cross-linking them in place through a reaction which does not involve the basic nitrogenous groups responsible for the sorption reaction in order to render the sorbing species non-volatile under the selected sorption conditions. Grafting or bonding methods are known in the technical literature. The molecular dimensions of the base sorbent should be selected in accordance with the pore dimensions of the support material since bulky bases or their precursors or derivatives may not be capable of entering pores of limited dimensions. A suitable match of base and support may be determined if necessary by empirical means.
Solid phase adsorbents will normally be operated in fixed beds contained in a suitable vessel and operated in the conventional cyclic manner with two or more beds in a unit with each bed switched between sorption and desorption and, optionally, purging prior to re-entry into the sorption portion of the cycle. Purging may be carried out with a steam of the purified gas mixture, i.e. a stream of the gas from which the H2S has been removed in the sorption process. If operated in temperature swing mode, a cooling step will intervene at some point between desorption and re-entry to sorption; this step will usually constitute a purge after desorption is completed. Alternatively, moving bed systems may be used with particulated solid sorbents or fluidized bed systems with finely-divided solids, e.g. with a particle size up to about 100 μm with the sorbent treated functionally as a liquid circulated between a sorption zone and a desorption/regeneration zone in a manner similar to a fluid catalytic cracking unit; rotating wheel beds are notably useful in rapid cycle sorption systems. All these systems may be operated in their conventional manner when using the present sorbents. Fixed bed systems may be operated with beds of solid porous particulate sorbents, porous monoliths or with layers of the sorbent on a porous or non-porous support For rapid cycle operation it may be possible to operate the separation using thin, adherent wash coats of the sorbent on plate type support elements.
The capped alkanolamine absorbents used in the present separation process comprise sterically hindered alkanolamines having an ether substituent capping all or some of the hydroxy groups which would otherwise be reactive towards the carbon dioxide to diminish H2S selectivity.
The steric hindrance required in the alkanolamine absorbent is provided by the group(s) attached to the amino acyclic or cyclic moieties attached to the amino nitrogen atom(s). The term “severely sterically hindered” signifies that the nitrogen atom of the amino moiety is attached to one or more bulky carbon groupings. Typically, the severely sterically hindered aminoether alcohols have a degree of steric hindrance such that the cumulative Es value (Taft's steric hindrance constant) greater than 1.75 as calculated from the values given for primary amines in Table V in D. F. DeTar, Journal of Organic Chemistry, 45, 5174 (1980), to which reference is made for a description of this parameter.
The 15N nuclear magnetic resonance (NMR) chemical shift provides another means for determining whether a secondary amino compound is “severely sterically hindered”. It has been found that the sterically hindered secondary amino compounds have a 15N NMR chemical shift greater than about δ+40 ppm, when a 90% by wt. amine solution in 10% by wt. D2O at 35° C. is measured by a spectrometer using liquid (neat) ammonia at 25° C. as a zero reference value. Under these conditions, the tertiary amino compound used for comparison, methyldiethanolamine, has a measured 15N NMR chemical shift value of δ 27.4. For example, 2-(2-tertiarybutylamino) propoxyethanol, 3-(tertiarybutylamino)-1-propanol, 2-(2-isopropylamino)-propoxyethanol and tertiarybutylaminoethoxyethanol had measured 15N NMR chemical shift values of δ+74.3, δ+65.9, δ+65.7 and δ+60.5 ppm, respectively, whereas the ordinary sterically hindered amine, secondary-butylaminoethoxyethanol and the non-sterically hindered amine, n-butylaminoethoxyethanol had measured 15N NMR chemical shift values of δ+48.9 and δ 35.8 ppm, respectively. When the cumulative Es values is plotted against the 15N NMR chemical shift values of the amino compounds mentioned above, a straight line is observed. Amino compounds having an 15N NMR chemical shift values greater than δ+50 ppm under these test conditions had a higher H2S selectively than those amino compounds having an 15N NMR chemical shift less than δ+50 ppm.
While hydroxyl-capped secondary and tertiary amines are preferred, capped primary alkanolamines such as monoethanolamine (MEA) are also useful and can be capped in the same way as the other alkanolamines. Aminoethers of this type are conveniently synthesized by amination of a capped alcohol or polyol in which the hydroxyl group(s) is/are replaced by amino group(s). Typically, the polyol will be a glycol; triols and higher polyols may be used for compounds with two or more capped hydroxyl groups but will not normally be preferred for reasons of economy and potential excess viscosity of the H2S reaction (sorption) products.
The preferred capped secondary alkyloxyamines may be made by the amination process described in U.S. 2010/0037775, to which reference is made for a description of the synthesis. In this amination process a capped glycol is reacted with a primary amine to form an aminoether. For example, to produce the preferred capped alkyloxyamines, an alkyloxy glycol is aminated by reaction with a primary amine to form the desired capped secondary aminoether product. Briefly, the amination reaction is carried out in the presence of a hydrogenation catalyst, preferably a nickel under hydrogen pressure at a temperature ranging from about 160 to about 425° C., preferably from about 180 to about 400° C., and most preferably from about 190 to about 250° C. The pressure in the reactor may suitably range from about 50 to about 3000 psig, preferably from about 100 to about 1000 psig, and most preferably from about 150 to about 750 psig.
The hydrogenation catalyst used in the amination process may be platinum, palladium and other noble metals on inert supports such as carbon, silica, alumina or other refractory oxides, Raney nickel, nickel-on-kieselguhr, nickel on inert support, massive nickel or nickel-cobalt or nickel-cobalt-copper coprecipitated with silicate and/or aluminum salts having alumina or kieselguhr supports. Preferred catalysts include coprecipitated nickel, massive nickel, nickel-cobalt, and nickel-cobalt-copper supported on silica, alumina or a mixture thereof. Also preferred is platinum supported on alumina. Further details of the amination catalysts are set out in U.S. Pat. No. 7,442,840 and 2010/0037775 to which reference is made for such details
The initial alkyloxy glycol may conveniently be produced by the Williamson ether synthesis in which an alkoxide (derived in situ from the corresponding alcohol and an alkali metal hydroxide) is reacted with an alkyl halide according to the generalized scheme:
where M is the alkali metal and X is the halide, e.g., Cl, I, Br and R1 and R2 are alkyl and alkylene groups, as above. The same or an alternative ether-forming technique may be used with triols and other polyols to cap the hydroxyls as needed, leaving one or more hydroxyl groups available for amination. One alternative to the Williamson synthesis, reacts the alkanolamine with an alkyl halide, preferably bromide, but the yield tends to be limited and the reaction has the added disadvantage of producing a corrosive hydrogen halide as a by-product. Another alternative is to cap an alkanolamine directly by reaction with an alkali metal hydride although in this case, the amino group of the starting alkanolamine needs to be protected, for example, by reaction with an aldehyde such as p-anisaldehyde, with removal of the protecting group following the methylation step by hydrolysis.
The capping group used to render the hydroxyl of the starting alkoxy glycol or polyol inaccessible to carbonation by the CO2 in the gas mixture is preferably an alkyl group, normally a short chain alkyl of 1 to 4 carbon atoms, methyl, ethyl, n-propyl, i-propyl or butyl (n-, i- or t-) so that the capped alkanolamine is a C1-C4 alkoxy amine.
In general terms, many of the present aminoether H2S absorbents containing secondary amino groups are defined by the formula:
R1—O—R2—NHR3
where R1, R2 and R3 are typically hydrocarbon or substituted hydrocarbon groups, typically alkyl or alkylene groups depending on their position in the molecule, e.g., R1 and R3 are C1-C4 alkyl or C1-C4 substituted alkyl and R2 is C1-C4 alkylene. It is preferred that the substituents should exclude hydroxyl in view of its reactivity with CO2 especially under higher pressure conditions but other, non-CO2 reactive substituent groups are acceptable, especially those polar substituents that confer enhanced water solubility when using aqueous systems. With alkanolamines which contain more than one hydroxyl group such as DEA, TEA or MDEA, the possibility of CO2 reaction at one or more of the available hydroxyl sites obviously arises so that reaction at these sites can be inhibited to the extent that the hydroxyls are capped by conversion to alkoxy groups. Thus, with DEA, one or both hydroxyls may be converted to alkoxy, preferably methoxy, groups and with TEA, from one to three of the hydroxyls may be converted in this way. Of course, the extent to which the carbonation reaction is inhibited depends upon the proportion of the hydroxyl groups which are effectively deactivated.
Among the capped alkanolamines that may be used in the present process are the following:
Other alternative capped secondary alkanolamines include the methoxy, ethoxy-, propoxy- and butoxy-capped ethers derived from the secondary aminothers described in U.S. Pat. No. 4,471,138, such capped ethers including the t-butylaminoethoxyethyl ethers, the 2-(2-t-butylamino)propoxyethyl ethers, the 2-(2-isopropylamino)propoxyethyl ethers, and the (1-methyl-1-ethylpropylamino)ethoxyethyl ethers.
As shown above, the amine functionality may be provided by a primary or a secondary or a tertiary amine group. Secondary amine groups provide additional steric hindrance from the two adjacent carbons than a hindered primary amine group and are generally preferred. This steric hindrance inhibits the reaction with the CO2 at conditions approaching the hydrosulfide/CO2 equilibrium when the kinetically faster reaction with the H2S has taken place.
Molecular weight is a consideration in the selection of a commercially useful absorbent since sorption operates on a molecular basis but absorbents are sold on a weight basis. Low molecular weight is therefore desirable if consistent with other factors especially selectivity. This factor therefore favors the use of ethanolamine and propanolamine ethers but their molecular weight and therefore absorption capacity per unit weight will need to be balanced against their selectivity. One example of this balancing is with the tertiary amine, dimethylamino ethyl methyl ether (DMAE-OMe), which is attractive from the viewpoint of low molecular weight (103 amu); this amine forms a bicarbonate in aqueous solution, but in non-aqueous systems the tertiary amine cannot form a carbamate or a bicarbonate and is thereby free to react exclusively with H2S. The secondary amine, 2-N-methylamino-2-methyl-prop-1-yl methyl ether (MAP-OMe) has a comparable molecular weight (115 amu) but generally has low inherent selectivity for H2S and is therefore not favored in this application although it is effective for CO2 separation. Thus, although the present capping procedure is effective for improving the inherent selectivity of an alkanolamine, it does not achieve high selectivity values with all alkanolamines. If high selectivity is the primary process objective to the exclusion of other considerations, the ethers of tertiary amines such as MDEA would be preferred with operation in a non-aqueous solvent: tertiary amines have no protons available for carbamate formation and in non-aqueous media cannot form bicarbonate; very good selectivities are therefore to be expected in such systems. Alternatively, more strongly basic secondary or tertiary amines of the guanidine/amidine/biguanide-type cannot react with CO2 in non-aqueous systems to form bicarbonates and because of the larger delta in pKa, between the amine and the acid gas, there exists a driving force for faster kinetics and higher selectivities for H2S absorption.
In the listing of exemplary capped alkanolamines above, one example of a partially capped alkanolamine is the 2-methoxyethyl-N-methyl-ethanolamine (conceptually a derivative of MDEA) which retains one hydroxyl function available for reaction with CO2. The completely capped alkanolamine is the succeeding one, bis-(2-methoxyethyl)-N-methylamine where both hydroxyls originating from the MDEA have been capped off by methoxy functionality and thus are unable to participate in the carbonation reaction with CO2. A similar progressive reduction in available hydroxyl functionality can be conceptualized with TEA where the hydroxyl groups might be successively converted to effect a stepwise progressive reduction in the hydroxyl functionality of the original molecule, passing from TEA to bis-(2-hydroxyethyl)-2-methoxyethyl-N-methylamine through the intermediate bis-(2-methoxyethyl)-2-hydroxyethyl-N-methylamine to the final tris-(2-hydroxyethyl)-N-methylamine.
Capped tertiary alkanolamines are also useful in the high pressure separation process; while tertiary amino alkanolamines are susceptible to reaction by carbonation on the hydroxyl groups with CO2 under higher pressure, the capped counterparts are largely immune and so offer a path to improved H2S selectivity. Thus, for example, etherifying the hydroxyl groups in MDEA to form bis-(methoxyethyl)-aminomethane inhibits the absorption of carbon dioxide and increases H2S/CO2 selectivity:
Other tertiary alkanolamines may be capped by etherification in a similar manner to improve their H2S selectivity.
The cyclic absorption process is normally operated with a solvent for the absorbent in order to permit ready circulation through the unit, especially to prevent undue viscosity increases with the H2S/capped amine reaction products in the rich solution leaving the bottom of the absorption tower. Aqueous and non-aqueous solutions may be used but while aqueous solutions may be preferred for reasons of economy, the optimal degree of H2S selectivity will be achieved with non-aqueous solutions since certain reaction products formed with CO2 are less stable in water and so apt to be more readily desorbed/hydrolyzed in the regeneration tower with a consequent decrease in H2S selectivity. As indicated by the comparative testing reported below, high H2S selectivity will be achieved by operating in non-aqueous systems and for this reason, non-aqueous solvents are normally preferred for optimum H2S selectivity although judicious selection of the solvent on an empirical basis may become necessary especially when operating with higher molecular weight absorbents as the hydrosulfide salts formed by reaction of the H2S at the amino nitrogen may be less soluble in non-aqueous media. Non-aqueous solvents would also be expected to be less corrosive, enabling the use of cheaper metallurgies, e.g., carbon steel, with reduced concern about corrosion at higher loadings; more polar non-aqueous solvents also minimize hydrocarbon solubility when they are evolved from natural gas wells at elevated levels.
Polar non-aqueous solvents such as toluene with a relatively low dipole moment may be found to be effective although in general, higher values for the dipole moment (Debye) of at least 2 and preferably at least 3 are to be preferred. Polar solvents such as DMSO (dimethyl sulfoxide), DMF (N,N-dimethylformamide), NMP (N-methyl-2-pyrrolidone), HMPA (hexamethylphosphoramide), THF (tetrahydrofuran) and the like are preferred from the viewpoint of potential reaction product solubility.
The preferred solvents preferably have a boiling point of at least 65° C. and preferably 70° C. or higher in order to reduce solvent losses in the process and higher boiling points are desirable depending on the regeneration conditions which are to be used. Use of higher boiling point solvents will conserve valuable energy which would otherwise be consumed in vaporization of the solvent.
Solvents potentially effective include toluene, sulfolane (tetramethylene sulfone) and dimethylsulfoxide (DMSO). Other solvents of suitable boiling point and dipole moment would include acetonitrile, N,N-dimethylformamide (DMF), tetrahydrofuran (THF), N-methyl-2-pyrrolidone (NMP), propylene carbonate, dimethyl ethers of ethylene and propylene glycols, ketones such as methyl ethyl ketone (MEK), esters such as ethyl acetate and amyl acetate, and halocarbons such as 1,2-dichlororobenzene (ODCB). Dipole moments (D) and boiling points for selected solvents are:
The positive effect resulting from the capping of free hydroxyl groups should be enhanced in non-aqueous systems, especially at higher pressures. Also, with regeneration at temperatures below the boiling point of water, desorption by PSA and/or TSA techniques becomes more readily attainable. The trends for selectivity, capacity and energy requirements are also favorable. If necessary, the incoming gas stream may be dried to reduce water accumulation in non-aqueous absorbent systems; for example, the incoming gas stream may be dried using conventional drying agents such as a glycol, usually diethylene glycol (DEG), triethylene glycol (TEG), propylene carbonate, or a solid dessicant such as activated alumina, granular silica gel, a small pore zeolite such as Zeolite-4A or a salt drying agent such as calcium chloride, potassium chloride, lithium chloride, sodium sulfate, or magnesium sulfate.
Both types of solvent—aqueous and non-aqueous—will, of course, tend to take up CO2 by direct physisorption under the high pressure conditions employed; desorption under the conditions in the regeneration can be expected with some decrease in selectivity.
The concentration of the capped alkanolamines absorbent in the solvent is determined empirically in the light of the particular operational mode, the concentration of acidic gases in the incoming gas stream, the selected absorbent and the solubility of the reaction products in the selected solvent with attention also to the viscosity of the rich solution. While a high concentration of the absorbent will favor lower circulation rates and possibly smaller unit size, viscosity and solubility issues may favor less concentrated solutions. In general terms, aqueous solutions (if used) may comprise from about 30 to 70 w/w percent of the absorbent while non-aqueous solutions may require a lower concentration as a result of the trend towards lower solubility with these systems.
The concentration of the capped alkanolamine in the selected solvent can vary over a wide range. Alkanolamine concentrations may typically range from 5 or 10 weight percent to about 70 weight percent, more usually in the range of 20 to 60 weight percent. Mixtures of capped alkanolamines can be used in comparable total concentrations. The concentration of the capped alkanolamine may be optimized for specific alkanolamine/solvent mixtures in order to achieve the maximum total absorbed H2S concentration, which typically is achieved at the highest alkanolamine concentration although a number of counter-balancing factors force the optimum to lower concentrations. Among these are limitations imposed by solution viscosity, solubilities of the alkanolamine and/or of the H2S reaction product, and solution corrosivity. In addition, as the concentration of the capped alkanolamine affects the nature of the H2S reaction product formed, the alkanolamine concentration also directly affects the required regeneration energy for a specific mixture. Therefore, the optimal alkanolamine concentration is selected to balance the maximum total absorbed H2S concentration and the lowest required regeneration energy, contingent upon the viscosity, solubility and corrosivity limitations described above; this concentration is likely to vary for individual combinations and is therefore to be selected on an empirical basis which also factors in the gas feed rate relative to the rate of sorbent circulation in the unit. The temperature and pKa of the capped alkanolamine compound also play into this equation.
The formation of precipitates is regarded as generally undesirable since, if precipitates are formed, the concentration of the active amine sorbent decreases and the amount of amine available for H2S capture, decreases accordingly. The formation of sulfide precipitates may, be exploited by separation of the solid or slurry of the solid, e.g., by hydrocyclone or centrifuge, followed by desorption of the H2S from the solid by heating. This enables the absorbent amine to be regenerated with lower energy requirements since much less solvent needs to be stripped, heated or vaporized.
Examples 1 to 4 below illustrate the synthesis of capped alkanolamines useful as absorbents in the present process.
The secondary amine 2-N-methylaminoethanol (3.76 g, 0.05 mol), N,N-diisopropylethylamine (DIPEA) (6.46 g, 0.075 mol), 2-methoxyethyl bromide (7.30 g, 0.0525 mol) and 30 mL acetonitrile were placed in a round bottom flask and stirred at room temperature under nitrogen. After completion of the reaction, (˜6 h, monitored by HPLC) the reaction mixture was evaporated under reduced pressure in a rotary evaporator. The residue was dissolved in 50 mL of dichloromethane and washed with 50 mL of 50% sodium hydroxide solution in water. The aqueous layer was washed with 3×15 mL portions of dichloromethane. The collected organic fractions were dried over sodium sulfate and the solvent was then removed under reduced pressure in a rotary evaporator at low 0-5° C. to yield the crude product finally purified by fractional vacuum distillation under sodium hydroxide to yield the product (1.6 g, 0.013 mol, b.p. ˜115° C., pressure is not available) as a colorless oil in 25% yield.
2-methoxyethyl-N-methyl-ethanolamine (MDEA-OMe), a colorless oil, was collected in a yield of 25%. 1H NMR (300 MHz, CDCl3) δ 3.63-3.56 (m, 2H), 3.48 (t, J=5.6 Hz, 2H), 3.36 (s, 3H), 2.93 (s, 1H), 2.64 (t, J=5.6 Hz, 2H), 2.61-2.55 (m, 2H), 2.33 (s, 3H). 13C NMR (75 MHz, CDCl3) δ 70.8, 59.0, 58.9, 58.9, 56.7, 42.8.
Bis(2-methoxyethyl)amine (35.45 g, 0.26 mol) was cooled to 0° C. in a 2-L round-bottom flask containing a stir bar. Following dropwise addition of 88% aqueous formic acid (47 mL, 0.91 mol), 37% aqueous formaldehyde (56 mL, 0.69 mol) was added. Controlled heating to 60° C. initiated rapid gas evolution. The reaction was allowed to proceed without further heating until gas evolution decreased (˜6 h) and was then heated to 80° C. for 24 h. The reaction mixture was cooled, acidified with 20% aqueous HCl, and extracted three times with 100 mL portions of diethyl ether. The aqueous layer was stirred in a salt/ice bath and brought to pH 12 by dropwise addition of 40% aqueous NaOH without allowing the internal temperature to exceed 25° C. Following separation of the resulting amine/aqueous layers, the aqueous layer was further extracted three times with 100 mL portions of diethyl ether. The combined organic layers were dried over sodium sulfate and solvent rotary evaporated under reduced pressure at low temperature. The resulting crude product was subjected to fractional vacuum distillation under sodium hydroxide to yield the product (17.42 g, 0.13 mol, b.p. 120-122° C., 35 Torr) as a colorless oil in 50% yield.
Bis-(2-methoxyethyl)-N-methylamine (MDEA-(OMe)2), a colorless oil, was collected in a yield of 50%. 1H NMR (300 MHz, CDCl3) δ3.44 (t, J=5.8 Hz, 4H), 3.29 (s, 6H), 2.57 (t, J=5.8 Hz, 4H), 2.27 (s, 3H). 13C NMR (75 MHz, CDCl3) δ 70.7, 58.8, 57.2, 43.2.
This Example demonstrates the synthesis of two alkoxy propylamine derivatives in a three stage synthesis in which the amino group on an initial propanolamine compound is first protected by p-methoxyphenyl protection (PMP-protection) to form a protected aminoalcohol which is then methylated on the hydroxyl group after which the protecting PMP group is removed to form the final methoxy substituted amine.
Step 1 p-Methoxyphenyl Protection (PMP-Protection)
A mixture of the selected amino alcohol (1 eq) and p-anisaldehyde (1.1 eq) was heated under reflux in benzene with azeotropic removal of water during 24 hours. The reaction was concentrated under reduced pressure. The desired products were recrystallized from hexane.
PMP-AP-OH, was collected as white microcrystals in a yield of 96%. 1H NMR (300 MHz, CDCl3) δ 8.09 (s, 1H), 7.53 (d, J=8.7 Hz, 2H), 6.80 (d, J=8.7 Hz, 2H), 3.74 (s, 3H), 3.69-3.49 (m, 1H), 3.48-3.27 (m, 2H), 1.10 (d, J=6.5 Hz, 3H). 13C NMR (75 MHz, CDCl3) δ 161.57, 160.86, 129.90, 128.75, 113.83, 67.61, 67.18, 55.32, 18.48.
PMP-AMP-OH was collected as white microcrystals in a yield of 92%, mp 52-53° C. (hexane). 1H NMR (300 MHz, CDCl3) δ 8.24 (s, 1H), 7.67 (d, J=8.7 Hz, 2H), 7.39 (d, J=8.4 Hz, 2H), 6.90 (t, J=8.6 Hz, 4H), 5.49 (s, 1H), 3.82 (s, 3H), 3.79 (s, 3H), 3.70 (d, J=7.5 Hz, 1H), 3.56 (d, J=7.4 Hz, 1H), 3.50 (s, 2H), 1.30 (s, 6H), 1.23 (s, 6H). 13C NMR (75 MHz, CDCl3) δ 161.6, 159.7, 156.9, 132.0, 129.6, 128.4, 127.2, 114.0, 113.8, 91.9, 77.9, 71.9, 60.4, 59.9, 55.3, 26.9, 26.3, 24.1.
The PMP-protected amino alcohol (1 eq) was reacted with sodium hydride (60% in mineral oil, 1.1 eq) in dry THF at 0° C. After 4 h stirring at room temperature, methyl iodide (1.1 eq) was added by dropwise to the reaction mixture. The resulting mixture was stirred at room temperature for 12 h. The reaction was quenched in water and extracted with dichloromethane. The organic layer was dried over sodium sulfate, filtered and concentrated under reduced pressure in rotary evaporator to give the desired methoxy ether of the initial PMP-protected amino alcohol.
PMP-AP-OMe, yield 95%, yellow oil. 1H NMR (300 MHz, CDCl3) δ 8.23 (s, 1H), 7.67 (d, J=8.7 Hz, 2H), 6.89 (d, J=8.6 Hz, 2H), 3.80 (s, 3H), 3.60-3.06 (m, 6H), 1.22 (d, J=6.0 Hz, 3H). 13C NMR (75 MHz, CDCl3) δ 161.49, 159.81, 129.77, 129.26, 113.71, 77.58, 65.70, 59.01, 55.31, 19.08.
PMP-AMP-OMe, yield 95%, yellow oil. 1H NMR (300 MHz, CDCl3) δ 8.24 (s, 1H), 7.69 (d, J=8.9 Hz, 2H), 6.90 (d, J=8.9 Hz, 2H), 3.82 (s, 3H), 3.36 (m, 5H), 1.26 (s, 6H). 13C NMR (75 MHz, CDCl3) δ 161.4, 156.4, 130.1, 129.6, 113.9, 81.6, 60.2, 59.6, 55.4, 24.7.
The methoxy ether of the PMP-protected amino alcohol was stirred 24 h in 250 mL of 5 N hydrochloric acid solution in water at room temperature. The reaction mixture was then washed with 3×75 mL portions of diethyl ether to extract p-anisaldehyde. The aqueous layer was stirred in a salt/ice bath and brought to pH 12 by dropwise addition of 40% aqueous NaOH without allowing the internal temperature to exceed 25° C., then was further extracted three times with 100 mL portions of diethyl ether. The combined organic layers were dried over sodium sulfate and the solvent was evaporated under reduced pressure at low temperature. The resulting crude product was subjected to fractional distillation under sodium hydroxide (bp of AMP-OMe ˜98-101° C.; AP-OMe ˜95-98° C.) at atmospheric pressure.
AP-OMe (308), yield 30%, colorless oil. 1H NMR (300 MHz, CDCl3) δ 3.35 (s, 3H), 3.27 (m, 1H), 3.19-3.03 (m, 2H), 1.03 (d, J=5.9 Hz, 3H). 13C NMR (75 MHz, CDCl3) δ 79.27, 58.31, 45.91, 19.34.
AMP-OMe, yield 70%, colorless oil. 1H NMR (300 MHz, CDCl3) δ 3.37 (s, 3H), 3.12 (s, 2H), 1.09 (s, 6H). 13C NMR (75 MHz, CDCl3) δ 82.9, 58.9, 49.7, 27.1.
The methoxy ether of PMP-protected amino alcohol (1 eq) was reacted with methyl triflate (1.1 eq) in dichloromethane under reflux to form the iminium salt in (monitored by 1H NMR) which was hydrolyzed with 100 mL of sodium hydroxide as 30% solution in water during 1 h at 20° C. The desired product was extracted with dichloromethane (2×100 ml) then dried under sodium sulfate followed by fractional distillation under sodium hydroxide at atmospheric pressure (bp of MAP-OMe ˜102-106° C.; MAMP-OMe ˜105-110° C.).
MAP-OMe, yield 39%, colorless oil. 1H NMR (300 MHz, CDCl3) δ 3.44-3.14 (m, 5H), 2.90-2.61 (m, 1H), 2.42 (s, 3H), 1.01 (d, J=6.4 Hz, 3H). 13C NMR (75 MHz, CDCl3) δ 76.94, 58.72, 54.20, 33.71, 16.34.
MAMP-OMe, yield 25%, colorless oil. 1H NMR (300 MHz, CDCl3) δ 1.03 (s, 6H), 2.29 (s, 3H), 3.18 (s, 2H), 3.36 (s, 3H). 13C NMR (75 MHz, CDCl3) δ 23.3, 28.4, 52.8, 59.1, 9.1.
Examples 5 to 13 below illustrate the extent to which capped and uncapped alkanolamines differ in their ability to react with CO2 in aqueous and non-aqueous solvents. The experiments were run as single component uptake experiments with CO2 only (which reacts with amine and —OH) in order to confirm CO2 uptake via O-carbonation of alkanolamines and absence of O-carbonation of methoxylated amines. In the presence of H2S, methoxylated amines will react preferentially with the H2S rather than with CO2 under conditions short of equilibrium between the two absorbing species (i.e. with short contact times) because the amino group tends to react faster with H2S and the methoxy group is no longer reactive towards the CO2.
The experimental setup for monitoring of amine acid gas uptake by was built inside a wide bore 400 MHz Bruker Avance™ nuclear magnetic resonance (NMR) spectrometer equipped with variable temperature capabilities. A 10 mm NMR tube placed inside the instrument and containing a solution of the desired amine, typically in H2O or d6-dimethylsulfoxide (DMSO-d6), was contacted with an acid gas, e.g., CO2 at desired pressure inside the instrument while recording quantitative 1H and 13C{1H} NMR spectra. Desorption/regeneration experiments were performed by decreasing the CO2 pressure and increasing the solution temperature if needed.
13C and 1H spectra taken before, during, and after the absorption/desorption sequence(s) gave quantitative information about the starting solution, reaction kinetics, and intermediate/final sorption products. The reaction products seen in 13C NMR spectra were identified and quantified by integration of the 13C NMR carbonyl resonance(s) at 165-164 ppm (representing CO2 as an ammonium carbamate), 161-160 ppm (representing CO2 as an ammonium bicarbonate), 159-158 ppm (representing CO2 in O-carbonate) versus resonances representing the amine —OCH2CH2N— and (if present) —NCH3 groups. Carbon dioxide dissolved in solution was detected at 125-124 ppm and accounted for as additional gas uptake. When desired, samples were transferred into a 5 mm NMR tube for more accurate ex-situ 10 and 2D NMR analysis on a Bruker Avance III™ narrow bore 400 MHz spectrometer.
The severely hindered secondary amine 2-N-methylamino-2-methylprop-1-yl methyl ether which has a methyl capped hydroxyl group was studied as an example of a compound, which does not react with CO2 but reacts with H2S in non-aqueous solution. The presence of the methoxy group also prevents an additional CO2 reaction with the hydroxyloxygen of the alkanolamine via an O-carbonation reaction. Such severely hindered amines such as MeO-MAMP in anhydrous solution can be used for very efficient separation of CO2/H2S based on fast reaction rates of an amine with H2S and of the slower CO2 reaction with the amine because of its steric hindrance.
The severely hindered secondary amine with a methyl capped hydroxyl group, 2-N-methylamino-2-methylprop-1-yl methyl ether (MeO-MAMP), was studied as an example of compound with slow CO2 reaction rates with the amine in aqueous solution. The methoxy-group of MeO-MAMP also prevents an additional CO2 reaction with the hydroxyloxygen of the amine. Severely hindered amines with capped hydroxyl groups such as MeO-MAMP can be used for kinetic separation of CO2/H2S based on fast reaction rates of an amine with H2S and slow reaction rates with CO2.
Unlike regular nucleophilic primary and secondary amines such as monoethanolamine (MEA) and N-methylaminoethanol (MAE), the hindered secondary amine MeO-MAMP does not form a carbamate reaction product with CO2 and directly forms bicarbonate/carbonate species. This reaction mechanism is characterized by a very long reaction constant characteristic of tertiary amines such as dimethylaminoethanol (DMAE) or triethanolamine (TEA) where the rate constant for direct bicarbonate formation with CO2 is 10-100 times lower.
The severely hindered secondary alkanolamine, 2-methylamino-2-methylpropan-1-ol (MAMP), was studied as a comparative example of a compound with slow CO2 reaction rates with an amine in aqueous solution. In contrast to MeO-MAMP with a methoxy-group, the hydroxyl group of MAMP is responsible for additional CO2 reaction with the hydroxyloxygen of an alkanolamine, which increase CO2 loading and decreases CO2/H2S separation efficiency.
The severely hindered secondary amine with methyl capped hydroxyl groups, bis-(2-methoxyethyl)-N-methylamine (2-MDEA-(OMe)2) was studied as an example of a compound, which does not react with CO2 but reacts with H2S in non-aqueous solution. The methoxy-groups of MDEA-(OMe)2 also prevent an additional CO2 reaction with the hydroxyloxygen of the amine via an O-carbonation reaction. Tertiary amines such as MDEA-(OMe)2 in anhydrous solution can be used for very efficient separation of CO2/H2S based on fast reaction rates of the amine with H2S and the slower CO2 reaction with the amine.
The tertiary amine MDEA-(OMe)2 with capped hydroxyl groups was studied as an example of a compound with slow CO2 reaction rates with an amine in aqueous solution. The methoxy-groups of MDEA-(OMe)2 prevent an additional CO2 reaction with the hydroxyloxygen of an amine. Tertiary amines with capped hydroxyl groups such as MDEA-(OMe)2 can be used for kinetic separation of CO2/H2S based on fast reaction rates of an amine with H2S and slow reaction rates with CO2.
Unlike regular nucleophilic primary and secondary amines such as monoethanolamine (MEA) and N-methylaminoethanol (MAE), tertiary amine MDEA-(OMe)2 does not form a carbamate reaction product with CO2 and directly forms bicarbonate/carbonate species. This reaction mechanism is characterized by a very long reaction constant. The rate constant for direct bicarbonate formation of tertiary or severely hindered amines with CO2 is 10-100 times lower.
The tertiary alkanolamine, methyldiethanolamine (MDEA), was studied as a comparative example because it is used commercially for H2S/CO2 separation. In contrast to MDEA-(OMe)2, the hydroxyl groups of MDEA are available for additional CO2 reaction which increases CO2 loading and decreases CO2/H2S separation efficiency.
The severely hindered primary amine 2-amino-2-methylprop-1-yl methyl ether which has a methyl capped hydroxyl group, was studied as an example of a compound, which reacts with CO2 slowly in non-aqueous solution while reaction with H2S is expected significantly faster. The presence of the methoxy group also prevents an additional CO2 reaction with the hydroxyloxygen of the alkanolamine via an O-carbonation reaction. Such severely hindered amines such as MeO-AMP in anhydrous solution can be used for very efficient separation of CO2/H2S based on fast reaction rates of an amine with H2S and of the slower CO2 reaction with the amine because of its steric hindrance.
O-carbonation reaction products in the region 159-158 ppm were not observed as well (see
The severely hindered primary amine with a methyl capped hydroxyl group, 2-amino-2-methylprop-1-yl methyl ether (MeO-AMP), was studied as an example of compound with slow CO2 reaction rates with the amine in aqueous solution. The methoxy-group of MeO-AMP also prevents an additional CO2 reaction with the hydroxyloxygen of the amine. Severely hindered amines with capped hydroxyl groups such as MeO-AMP can be used for kinetic separation of CO2/H2S based on fast reaction rates of an amine with H25 and slow reaction rates with CO2.
The severely hindered primary alkanolamine, 2-amino-2-methylpropan-1-ol (AMP), was studied as a comparative example of a compound with slow CO2 reaction rates with an amine in aqueous solution. In contrast to MeO-AMP with a methoxy-group, the hydroxyl group of AMP is responsible for additional CO2 reaction with the hydroxyloxygen of an alkanolamine, which increase CO2 loading and decreases CO2/H2S separation efficiency.
The moderately hindered secondary amine with a methyl capped hydroxyl group, 2-N-methylamino-prop-1-yl methyl ether (MeO-MAP), was studied as an example of compound with fast CO2 reaction rates with the amine in aqueous solution. The methoxy-group of MeO-MAP prevents an additional CO2 reaction with the hydroxyloxygen of the amine but helps to maintain solution viscosity. Moderately hindered amines with capped hydroxyl groups such as MeO-MAP cannot be used for kinetic separation of CO2/H2S based because reaction rates of H2S and CO2 with an amine are similar. However, moderately hindered secondary amines with capped hydroxyl groups such as MeO-MAP can be utilized for effective CO2 capture from various gases such as flue gas and natural gas
At 10.0 bar of CO2 and 45° C., carbamate completely hydrolyzed into bicarbonate (not shown here). The equilibrium CO2 loading at given conditions is 1.00 CO2 per amine with all CO2 molecules present in bicarbonate. O-carbonation and dissolved CO2 was not detected by 13C NMR.
The moderately hindered secondary alkanolamine, 2-N-methylamino-propan-1-ol (MAP), was studied as an example of compound with fast CO2 reaction rates with the amine in aqueous solution. In contrast to MeO-MAP with a methoxy-group, the hydroxyl group of MAP is responsible for additional CO2 reaction with the hydroxyloxygen of an alkanolamine, which increase CO2 loading. Moderately hindered alkanolamines such as MAP cannot be used for kinetic separation of CO2/H2S based because reaction rates of H2S and CO2 with an amine are similar. However, moderately hindered secondary alkanolamines such as MAP can be utilized for effective CO2 capture from various gases such as flue gas and natural gas
At 10.0 bar of CO2 and 30° C., carbamate completely hydrolyzed into bicarbonate. The equilibrium CO2 loading at given conditions is 1.00 CO2 per amine with all CO2 molecules present in bicarbonate. The equilibrium CO2 loading and the contribution of reaction products for the alkanolamines and amino-ethers tested in Examples 5 to 13 are summarized in Table 1 together with the results of testing the uncapped alkanolamines in aqueous solution with product speciation as CO2/amine molecule) at 10.0 bar of CO2 and 45° C. in aqueous and non-aqueous solution. Use of non-aqueous solvent and capping of the hydroxyl group(s) of an alkanolamine leaves the amine nitrogen free to preferentially react with hydrogen sulfide.
Selectivity studies on removal of H2S from CO2-containing gas feeds were performed by purging an acid gas mixture through a reactor vessel containing an amine absorbent solution and analyzing the gas composition exiting the reaction vessel.
The experimental setup consisted of six main elements: (i) an N2 purge gas supply, (ii) an acid gas supply containing a mixture of H2S/CO2/N2, (iii) a 4-way valve to facilitate switching gas feeds between inlets for the N2 gas and the acid gas mixture, (iv) a bubbler type reactor vessel containing an amine solution, (v) a mass spectrometer and (vi) acid gas scrubber. The 4-way valve selects the feed gas (N2 or the acid gas mixture) and directs it to either the reactor vessel or to the scrubber. The outlet from the reactor vessel was connected to the scrubber with a connection to the mass spectrometer to permit real time analysis of the effluent gas composition. Approximately 15 cc of amine solution was placed into the reaction vessel (40 cc) containing an inlet tube that reaches near to the bottom of the vessel and an outlet tube connected to the gas scrubber and the mass spectrometer.
The reaction vessel containing the amine solution was first flushed with inert gas (e.g., N2) to remove air from the head space. Flow of H2S and CO2 of a given concentration in N2 was initiated and directed into the scrubber vessel in order to flush lines and stabilize the flow. After the reaction vessel was flushed with N2 and the mass spectrometer detected low concentrations of O2, H2O and CO2, the 4-way valve was turned to expose the amine solution to the H2S/CO2 mixture, at which point the run was considered to be at zero time. The mass spectrometer quantitatively detects the real-time off-gas composition, namely the concentration of CO2 and H2S as a function of time in the gas after treatment by the amine solution. Rigorous unit calibration was performed to calibrate the mass spectrometer signals taking into account for the delayed gas breakthrough due to filling the finite system volume. Each experimental sequence was composed of two runs: (1) gas flowing through an empty reactor vessel without amine and (2) the same gas composition flowing through the reactor vessel containing the amine solution.
Representative experimental data is presented below for 1M solutions of three amines Di-MeO-MDEA, MeO-MAMP, and 1,1,3,3-tetramethylguanidine (TMG) dissolved in NMP and two gas mixtures containing 0.1% H2S/9.0% CO2/90.9% N2 and 0.5% H2S/5.0% CO2/94.5% N2 purged through the amine solution at a flow rate of 100 sccm. The data includes the gas composition after treatment with an amine solution, the derived rates of H2S and CO2 capture as a function of reaction time and H2S/CO2 selectivities calculated as a ratio of relative concentrations of H2S and CO2 in the liquid and gas phases.
The results of the testing show that the equilibrium absorption factors and kinetics of the separation process should be factored into the operation of the unit: while the H2S is initially absorbed selectively relative to the CO2, continued passage of the acid gas mixture through the absorbent solution eventually leads to displacement of the H2S by reaction of the amine groups with the CO2. For this reason, the relative flow rates of the incoming gas mixture and of the absorbent solution should be controlled in combination with the compositions of the solution and the gas mixture so as to maintain the separation within the regime affording selectivity for H2S absorption.
A gas mixture containing 0.1% H2S, 9.0% CO2 and 90.9% N2 was purged through 15.1 g of a 1M solution of MDEA-(MeO)2 in NMP at 22.5° C. and 0.4 psig (3 kPag).
A gas mixture containing 0.5% H2S, 5.0% CO2 and 94.5% N2 was purged through 15.1 g of neat MDEA-(MeO)2 at 22.5° C. and 0.4 psig (3 kPaG).
A gas mixture containing 0.5% H2S, 5.0% CO2 and 94.5% N2 was purged through 15.0 g of a 1M solution of MeO-MAMP in NMP at 22.5° C. and 0.4 psig (3 kPag).
A gas mixture containing 0.1% H2S, 9.0% CO2 and 90.9% N2 was purged through 15.0 g of a 1M solution of 1,1,3,3-tetramethylguanidine (TMG) in NMP at 22.5° C. and 0.4 psig (3 kPag).
A gas mixture containing 0.5% H2S, 5.0% CO2 and 94.5% N2 was purged through 15.0 g of a 1M solution of 1,1,3,3-tetramethylguanidine (TMG) (pKa 15.2) in DMSO at 22.5° C. and 0.4 psig (3 kPag).
This application claims priority under 35 USC 120 from U.S. Application Ser. No. 61/859,325, filed 29 Jul. 2013.
| Number | Date | Country | |
|---|---|---|---|
| 61859325 | Jul 2013 | US |
