CA1074534A - Process for acid gas removal - Google Patents
Process for acid gas removalInfo
- Publication number
- CA1074534A CA1074534A CA263,738A CA263738A CA1074534A CA 1074534 A CA1074534 A CA 1074534A CA 263738 A CA263738 A CA 263738A CA 1074534 A CA1074534 A CA 1074534A
- Authority
- CA
- Canada
- Prior art keywords
- stage
- solution
- gas
- hydrogen sulfide
- alkanolamine
- Prior art date
- Legal status (The legal status is an assumption and is not a legal conclusion. Google has not performed a legal analysis and makes no representation as to the accuracy of the status listed.)
- Expired
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Classifications
-
- B—PERFORMING OPERATIONS; TRANSPORTING
- B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
- B01D—SEPARATION
- B01D53/00—Separation of gases or vapours; Recovering vapours of volatile solvents from gases; Chemical or biological purification of waste gases, e.g. engine exhaust gases, smoke, fumes, flue gases, aerosols
- B01D53/14—Separation of gases or vapours; Recovering vapours of volatile solvents from gases; Chemical or biological purification of waste gases, e.g. engine exhaust gases, smoke, fumes, flue gases, aerosols by absorption
- B01D53/1493—Selection of liquid materials for use as absorbents
-
- B—PERFORMING OPERATIONS; TRANSPORTING
- B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
- B01D—SEPARATION
- B01D53/00—Separation of gases or vapours; Recovering vapours of volatile solvents from gases; Chemical or biological purification of waste gases, e.g. engine exhaust gases, smoke, fumes, flue gases, aerosols
- B01D53/14—Separation of gases or vapours; Recovering vapours of volatile solvents from gases; Chemical or biological purification of waste gases, e.g. engine exhaust gases, smoke, fumes, flue gases, aerosols by absorption
- B01D53/1456—Removing acid components
- B01D53/1468—Removing hydrogen sulfide
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- Chemical & Material Sciences (AREA)
- Engineering & Computer Science (AREA)
- Analytical Chemistry (AREA)
- General Chemical & Material Sciences (AREA)
- Oil, Petroleum & Natural Gas (AREA)
- Chemical Kinetics & Catalysis (AREA)
- Gas Separation By Absorption (AREA)
- Treating Waste Gases (AREA)
- Industrial Gases (AREA)
Abstract
PROCESS FOR ACID GAS REMOVAL
ABSTRACT OF DISCLOSURE
A continuous process for the selective absorption of hydrogen sulfide from a feed gas comprising an acid gas mixture of carbon dioxide and hydrogen sulfide is provided which comprises counter-currently contacting the feed gas with an aqueous alkanolamine solution in an absorption zone having from two to ten separate stages, contacting the feed gas in each stage with lean aqueous alkanolamine solution and withdrawing rich aqueous alkanolamine solution from the bottom of esch stage, in each stage maximizing the equilib-rium approach between the hydrogen sulfide in the gas and liquid phases and minimizing the equilibrium approach between the carbon dioxide in the gas and liquid phases, and maintaining the rich solution loading in each stage within the range from about O.l to about 0.3 mole of acid gas per mole of alkanolamine. The absorption zone is operated in combination with a stripping zone to which rich aqueous alkanolamine solution is passed and in which lean solution is generated. The particular multi-stagP
~bsorption process of the invention allows for the re-covery as stripping zone overhead of gaseous product in which the combined molar ratio of carbon dioxide to hydrogen sulfide is no greater than about 6.
S P E C I F I C A T I O N
ABSTRACT OF DISCLOSURE
A continuous process for the selective absorption of hydrogen sulfide from a feed gas comprising an acid gas mixture of carbon dioxide and hydrogen sulfide is provided which comprises counter-currently contacting the feed gas with an aqueous alkanolamine solution in an absorption zone having from two to ten separate stages, contacting the feed gas in each stage with lean aqueous alkanolamine solution and withdrawing rich aqueous alkanolamine solution from the bottom of esch stage, in each stage maximizing the equilib-rium approach between the hydrogen sulfide in the gas and liquid phases and minimizing the equilibrium approach between the carbon dioxide in the gas and liquid phases, and maintaining the rich solution loading in each stage within the range from about O.l to about 0.3 mole of acid gas per mole of alkanolamine. The absorption zone is operated in combination with a stripping zone to which rich aqueous alkanolamine solution is passed and in which lean solution is generated. The particular multi-stagP
~bsorption process of the invention allows for the re-covery as stripping zone overhead of gaseous product in which the combined molar ratio of carbon dioxide to hydrogen sulfide is no greater than about 6.
S P E C I F I C A T I O N
Description
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FIELD OF THE INVENTION
.
This invention relates to a process for acid gas removal and, more particularly, to the selective removal of hydrogen sulfide from various gas streams.
BACKGROUND OF THE INVENTION
The selective removal of hydrogen sulfide from carbon dioxide containing gas streams by absorption is an important, albeit highly specialized, segment of industrial technology. Hydrogen sulfide is especially useful in the manufacture of elemental sulfur, but in order to use it effectively, the hydrogen sulfide should be made available in a molar ratio of carbon dioxide to hydrogen sulfide approaching and preferably no greater than about 6. In some applications such as, for example, in a Claus plant, the molar ratio should be no grea~er than about 3. San~e the usual molar ratios found in typical feedstocks such as natural gas and synthetic natural gas, are in the neighbor-hood of about 5 to about 140, reduction of these ratios is usually a prerequisite to the use of hydrogen sulfide so obtained. This means that the selectivity of any commer-cially acceptable selective absorption process must be such that a high proportion of carbon dioxide passes through the absorber unabsorbed while a small proportion of the hydrogen sulfide follows that same path. The most advanta-geous process from a competitive point of view is, of course, the one which is most selective and has apparatus and energy requirements equal to or less than other commercial processes.
A principal object of this invention, therefore, is to provide a process for the selective removal of hydrogen sulfide from carbon dioxide containing gas streams whereby
FIELD OF THE INVENTION
.
This invention relates to a process for acid gas removal and, more particularly, to the selective removal of hydrogen sulfide from various gas streams.
BACKGROUND OF THE INVENTION
The selective removal of hydrogen sulfide from carbon dioxide containing gas streams by absorption is an important, albeit highly specialized, segment of industrial technology. Hydrogen sulfide is especially useful in the manufacture of elemental sulfur, but in order to use it effectively, the hydrogen sulfide should be made available in a molar ratio of carbon dioxide to hydrogen sulfide approaching and preferably no greater than about 6. In some applications such as, for example, in a Claus plant, the molar ratio should be no grea~er than about 3. San~e the usual molar ratios found in typical feedstocks such as natural gas and synthetic natural gas, are in the neighbor-hood of about 5 to about 140, reduction of these ratios is usually a prerequisite to the use of hydrogen sulfide so obtained. This means that the selectivity of any commer-cially acceptable selective absorption process must be such that a high proportion of carbon dioxide passes through the absorber unabsorbed while a small proportion of the hydrogen sulfide follows that same path. The most advanta-geous process from a competitive point of view is, of course, the one which is most selective and has apparatus and energy requirements equal to or less than other commercial processes.
A principal object of this invention, therefore, is to provide a process for the selective removal of hydrogen sulfide from carbon dioxide containing gas streams whereby
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the selectivity is such that molar ratios of carbon dioxide to hydrogen sulfide approaching about 6, and preferably about 3, or less are realized for the overall process while maintaining other process requirements at levels equal to or better than competitive processes.
Other objects and advantages of the present invention will become apparent from the following descrip-tion and disclosure.
SUMMARY OF THE INVENTION
The present invention provides improvement upon a continuous process for the selective absorption of hydro-gen sulfide from a feed gas comprising an acid gas mixture of carbon dioxide and hydrogen sulfide, said process comprising the steps of (I) counter-currently contacting the feed gas in an absorption zone with a lean aqueous alkanolamine solution to provide a rich aqueous alkanolamine solution, (II) introducing the rich solution into a stripping zone to provide a mixture of acid gas and water vapor overhead and lean solution as bottoms, and ~III) recycling the lean solution to the absorption zone, the said improvement comprising:
(a) using an alkanolamine having the structural formula, R
R - N
R
where R is an alkanol radical of 2 or 3 carbon atoms and is unsubstituted or methyl-substituted, or an alkyl radical having l to 5 carbon atoms provided that at least one R
is an alkanol radical;
(b) using a lean solution which has a molality of about 1.5 to about 75 and a maximum loading of about 0.1 mole of acid gas per mole of alkanolamine;
(c) using an absorption zone having 2 to lO
separate stages, wherein the equilibrium approach between the hydrogen sulfide in the gas and liquid phases is maxi-mized and the equilibrium approach between the carbon dioxide in the gas and liquid phases is minimized;
(d) adjusting the flow rate of the lean solution through each stage so that ~i) from about 0.1 to about 0.9, by volume, of the hydrogen sulfide passing through each stage is absorbed by the lean solution passing through said stage, and (ii) the rich solution loading in each stage is fro~ about 0.1 to about 0.3 mole of acid gas per mole of alkanolamine; and (e) introducing lean solution from the stripping zone into about the top of each stage, removing rich solution from about the bottom of ea~h stage and introducing rich solution into the stripping zone.
Recoverable as stripping zone overhead is gaseous product in which the molar ratio of carbon dioxide to hydrogen sulfide is no greate.r than about 6.
The combination of conditions employed in accord-ance with the teachings of this invention allows for essentially complete removal from the feed gas of hydrogen sulfide in high selectivity relative to carbon dioxide, in that the conditions are such that displacement of absorbed hydrogen sulfide by carbon dioxide in the feed gas is minimized and the driving force for hydrogen sulfide 4.
~074534 9945-1 absorption is maximized, as explained in greater detail hereinbelow.
BRIEF DESCRIPTION OF THE DRAWING
The sole figure of the drawing is a schematic flow diagram of an illustrative embodiment of the present invention.
DETAILED DESCRIPTION AND PREFERRED EMBODIMENTS
As noted above, in addition to the steps and conditions stated heretofore, each stage of the absorption process functions in a particular manner to accomplish the optimum level of selectivity which can be achieved by the presently described process. The function of the selective removal process of this invention encompasses two primary aspects. The first aspect is to maximize the equilibrium approach between hydrogen sulfide in the liquid and gas phases, and minimize the equilibrium approach between the carbon dioxide in the liquid and gas phases. The ability to control these equilibrium approaches is associated with the mass transfer characteristics of the vapor-liquid system. That is, the equilibrium approach between H2S in the gas and liquid phases is enhanced through any of several techniques designed to increase gas phase mass transfer rate. Furthermore, the equilibrium approach between the C2 in the gas and liquid phases is minimized through techniques capable of reducing the mass transfer rate in the liquid phase. In both cases, these effects may be induced through the proper tray design and selection of operating conditions. There are several means of achieving this particular aspect. One practical means, ~74534 9945-1 aLs an example, is to operate with a high superficial feed gas velocity through each stage sufficient to increase gas phase mass transfer rates without significantly increasing the liquid phase mass transfer r~te by maintaining relatively quiescent liquid phase behavior. Another practical means may be utilized in systems already possessing high gas phase mass transfer characteristics. Here, the velocity of feed gas through each stage and tray, or tray equivalent, is inhibited to such an extent that the flow is sufficient but essentially no greater than that flow which is required to maintain the solution on the tray. In so doing, the liquid mass transfer rates and rich solvent acid gas load-ings are reduced, thus further minimizing the equilibrium approach between the C02 in the gas and liquid phases.
A second and equally important aspect of the process of the present invention is to maintain the rich solution loadings in the range of about 0.1 to 0.3 mole of acid gas per mole of alkanolamine. At greater loadings, as the absorption of CO2 and H2S begins to approach nearer and nearer to equilibrium, the CO2 will begin to displace absorbed H2S in view of the fact that carbon dioxide is a stronger acid than hydrogen sulfide. Such displacement is undesirable because selectivity toward net hydrogen sulfide absorption is adversely affected. Furthermore, maintenance of the net solution loadings within the aforesaid range serves to keep a high rate of H2S transfer from the gas phase into the liquid phase by minimizing the increase in the partial pressure of H2S above the solution. To be con-sidered in this latter respect is the fact that the partial pressure of hydrogen sulfide above the solution is influ-1074S~ 9945-1 .
enced by two factors. First, as H2S loading increases, the partial pressure of H2S above the solution rises. Secondly, and importantly, an increase in CO2 loading will also raise the H2S partial pressure above the solution. Therefore, the net driving force between H2S in the gas phase and H2S
in equilibrium with the liquid phase decreases rapidly as loading increases. Multi-staging in accordance with the teachings of the present invention allows one to maximize the driving force throughout the absorber by replacement of the loaded solution at each stage with the introduction of lean solution, with its inherent low partial pressure of both acid gas components. By removing only a fraction of the acid gases per stage and not allowing the rich solution loadings to rise beyond the stated range, high selectivity toward hydrogen sulfide absorption relative to carbon dioxide absorption is obtained.
As noted, the feed gas which is fed to the absorption zone of the process of this invention comprises an acid gas containing carbon dioxide and hydrogen sulfide.
The feed gas may be (A) a mixture of a process gas and the acid gas or (B) the acid gas. The process gas comprises a hydrocarbon or a mixture of hydrocarbons. Examples of hydrocarbons that can be processed in the system are methane, methane which may be in the form of natural gas or substitute or synthetic natural gas (SNG), ethylene, ethane, propylene, propane, mixtures of such hydrocarbons, and the prepurified effluents from the cracking of naphtha or crude oil or from coal gasification. There are no limits to the throughput of feed gas in the subject process provided that the apparatus is sized correctly, a convention-al engineering problem.
7 ~ r 107453~ 9945-1 The feed gas either contains an acid gas together with the process gas or is the acid gas itself. The acid gas is a mixture of carbon dioxide and hydrogen sulfide generally in a molar ratio of about 5 moles to about 140 moles of carbon dioxide per mole of hydrogen sulfide.
The ratio of process gas to acid gas where both are present can cover a very broad range since the process can handle any ratio. This is simply because the process gas is inert insofar as the absorbent is concerned and, aside from throughput sizing, no provision has to be made for it.
Water can be and is usually present in mixture with all of the feed gas components in the form of water vapor or droplets in amounts running from 0 to saturated and is preferably saturated since saturation minimizes water evaporation in the bottom of the absorption zone. An anhydrous feed gas may ~be used but is very rare. The water referred to here is not considered in the determin-ation of molality unless and until it goes into solution with the alkanolamine.
Impurities as defined herein are represented by (a) any gas not defined above as a process gas, acid gas, or water vapor, and (b) solid particles or liquid droplets (exclusive of water droplets) in the feed gas. They can be present in amounts of up to about 3 weight percent based on the total weight of the feed gas and are preferably present in amounts no greater than about l weight percent and, in many cases, lower than 0.01 percent. Examples of the gaseous impurities are sulfur dioxide, carbonyl sulfide, and carbon disulfide. Examples of the solid or liquid 9945-l !
1~)74534 impurities are iron sulfide, iron oxide, high molecular weight hydrocarbons, and polymers. Any olefins having more than one double bond, triple bond hydrocar~ons, and as a general rule, any material that will polymerize or react in situ is an undesirable impurity.
The absorbent is a solution of an alkanolamine and water, the alkanolamine having the structural formula:
R
R - N
R
wherein each R is the same or different and can be at least one alkanol radical, which has 2 or 3 carbon atoms and is unsubstituted or methyl-substituted, or R can be an alkyl radical having 1 to 5 carbon atoms, provided at least one R
is an alkanol radical.
Examples of the alkanolamines are the preferred methyldiethanolamine (MDEA), triethanolamine (TEA), ethyl-diethanolamine (EDEA), tripropanolamine, and triisopropanol-amine. Although mixtures of alkanolamines can be used, they are nst preferred.
There is enough water in the solution or added to the system to provide a molality in the range of about 1.5 to about 75 and preferably from about 5.5 to about 12.5.
The determinatisn of molality is made on the basis of alkanolamine as solute and water as solvent wherein the molality of the solution is equal ,o the number of moles of solute (alkanolamine) dissolved in 1000 grams of solvent ;
(water). The molality in subject process concerns the lean alkanolamine solution which is used to contact the feed gas in each stage of the absorption zone. Other water in the system is not considered in the determination.
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Where aqueous MDEA solution (lean~ is introduced into a system, the concentration thereof is generally from about 10 percent to about 90 percent by weight MDEA based on the weight of the solution and is preferably from about 45 percent ~o about 55 percent by weight. Again, this solution should either provide the correct molality for the process or additional water must be added to the system to do so.
Examples of typical solutions, in percent by weight, are: MDEA 50 percent - water S0 percent; TEA 60 percent - water 40 percent; and EDEA 55 percent - water 45 percent.
Although generally the amount of water for all alkanolamines lies in the range of about 10 to about 90 percent by weight based on the total weight of the solution and the soIution preferably has the proper viscosity for pumping, the amount of water is determined in the end by molality in the ranges set forth above.
The apparatus used in the process for stripping, heat exchanging, and cooling as well as reboilers, filters, piping, turbines, pumps, flash tanks, etc., are of conven-tional design. A typical stripping or regenerator column used in the system can be described as a sieve tray tower having 15 to 20 actual trays or its equivalent in packing.
The stripper contains in its base, or in an external kettle, a tubular heating element or reboiler and at the top of and external to the stripper are condensers and a water separator.
The absorber is not of conventional design, however, as will be seen below.
10 .
Referrin~ to the Drawin~;
-As can be seen in the drawing, the absorber isnot one column, but is divided into separate stages, which are connected to each other in series. The only direct connection between the stages is line 9 through which passes the feed gas. The liquid from one stage never enters another stage unless and until it passes through a still. There can be two to ten stages.
The number of stages is increased within the two to ten stage range in accordance with increased feed gas purification requirements. It is understood that more than ten stages can be used in special cases to achieve the ulti-mate in purity since only a fraction of hydrogen sulfide in the range of about 0.1 to about 0.9, by volume, is removed in each stage, that is, from about 10 to about 90 percent by volume of the hydrogen sulfide content of the feed gas passing through each stage. Normally, the fraction of hydrogen sulfide removed from the feed gas in each stage is within the range from about 0.3 to about 0.7, by volume.
The removal of relatively high (i.e., up to about 0.9) or relatively low (i.e., down to about 0.1) fractions of hydro-gen sulfide passing through each stage depends on the rich-ness and leanness, respectively, of the hydrogen sulfide in the feed gas stream.
Each stage can have one to ten, and usually has no more than six, actual trays or its conventional equiv-alent in spray towers, co-current contactors, or packed towers, for example. In the discussion to follow, it should be understood that the terms "stage" and "tray"
contemplate the use of equivalent apparatus. For example, 11 .
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a stage having three trays is considered to be the same as a packed tower having the packing equivalent to three trays Least preferred are stages having more than four trays. The number of trays is selected by the operator to achieve the result desired in the work specifications with respect to combined molar ratio of carbon dioxide to hydrogen sulfide required in the stripper overhead, which in this process is set at about six or less. As-suming that the same tray design is used for each tray in the stage, the molar ratio of C02 to H2S in the rich solution in each stage will decrease as the number of the trays in the stage decreases, when operating in accordance with the teachings of the invention. Thus, it is clear that a broad range of purities and molar ratios can be achieved by this process.
From the standpoint of apparatus, the process of the present invention may be carried out in a variety of different types of absorbers. Illustrative of suit-able absorbers are those equipped with sieve trays, bubble cap trays, valve trays and the like, having zero or minimum weir heights. These various types of trays may be operated in one or more specific fashions. For example, in absorbers equipped with sieve trays, the flow of feed gas through the tray may be set to such an extent that the flow is sufficient, but essentially no greater than that flow which is required to maintain the solution in the tray. This may be more easily understood by referring to the Chemical Engineers' Handbook, Perry et al., 4th Edition, McGraw-Hill, 1963, Sectîon 18, pages 18-3 to 18-24, inclusive. At page 18-5, it is noted that weeping takes place when the vapor flow is at a rate ~074S34 insufficient to maintain the liquid on the tray. Note also page 18-14 where Perry points out that "dumping" in bubble-cap trays is analogous to "weeping" in sieve trays. The objective here is to provide a tray in which the rate of flow of the vapor is just sufficient to maintain the liquid on the tray, i.e., just above the weeping point. In other words, the tray is made to operate as inefficiently as p~ssible, which must be considered rather unconventional to say the least. Even though the tray is operated in this manner, the tray design is such that the contact efficiency or area between the gas and liquid phases is maximized while minimizing the liquid hold up time on the tray or, in other words, minimizing the contact time between the gas and liquid phases. Thus, operation of a sieve tray close to the "weep-ing point" meets the aforesaid objective.
At this point it should be noted that, although it is desirable to operate the absorber and/or utilize trays or their equivalent such that the contact time between the gas and the liquid phases is as low as possible, thereby taking advantage of the faster rate of reaction between the alkanolamine solution and H2S relative to the rate at which C02 reacts, low contact time is only one factor to be considered in providing a highly selective H2S removal process. Counteracting the initial and in-herently faster rate of absorption of H2S relative to C02, are factors such as the ability of C02 to displace ab-sorbed H2S, the rise in the partial pressure of H2S above the liquid phase as C2 is absorbed, and mass transfer characteristics. Thus, unless the displacement factor as a function of acid gas loading is minimized and the mass transfer driving forces as a function of loading, are such as to favor H2S absorption and, further, unless the S gas phase transfer is enhanced while minimizing C02 :Liquid phase transfer, a reproducibly highly selective H2S
removal process is not achieved. For example, inasmuch as H2S absorption is gas-film controlled whereas C02 ab-sorption is liquid-film controlled, operation at low con-tact time but with a high liquid phase transfer does not provide absorption of H2S in high selectivity.
The feed gas, which generally contains a molar ratio of carbon dioxide to hydrogen sulfide in the range of about 5 to about 140 moles of carbon dioxide per mole of hydrogen sulfide, is introduced at line 9 into stage 1 below the bottom tray (i.e., tray 4 of the drawing), the inlet temperature of the feed gas entering stage 1 being in the range of about 30C. to about 43C. The feed gas flows upwardly through stage 1 to countercurrently meet the aqueous alkanolamine solution referred to as lean solution, i.e., it contains less than about 0.1 mole of acid gas per mole of alkanolamine, which is introduced into stage 1 at tray 1 through line 11.
The pressure in stage 1 is in the range of about atmospheric pressure to about 2000 psia and is normally in the range of about 25 psia to about 1200 psia.
The lean solution enters stage 1 at a temperature which is normally in the range of about 30C. to about 43C.
The feed gas, which has had roughly about half of its hydrogen sulfide absorbed in passing up the stage 1 column, exits through line 9 and passes into stage 2. The stage 1 outlet temperature of the feed gas is normally in 14.
1~)74534 the range of about 30C to about 43C. It is noted that there is very little heat transfer Cor heat loss) from stage to stage.
The gas exiting the first stage contains a ratio, by volume, of carbon dioxide to hydrogen sulfide in the range of about 10 parts to about 280 parts of carbon dioxide per part of hydrogen sulfide, and is generally, saturated if the initial feed gas was saturated.
After the lean solution absorbs acid gas in stage 1, it is referred to as rich solution, i.e., the rich solution is a mixture of lean solution, absorbed acid gas, additional water picked up from the feed gas, and some impurities. The "rich solution loading", which is expressed as the ratio of moles of acid gas to moles of alkanolamine in the rich solution, is, in the first stage, and each stage thereafter, in the range of about 0.1 to about 0.3, .-as measured at about the bottom of each stage where stage outlets 10, 12, and 14 are located. The rich solution r exits the first stage at or below the bottom tray (i.e., ~0 tray 4 of the drawing) at a first stage outlet temperature which is normally in the range of about 30C. to about 43C~ This rich solution then proceeds to a common line, which is also fed by the rich solutions from the other stages 2 and 3 through lines 12 and 14, respectively.
It will be noted that the feed gas outlet temperature and the rich solution outlet temperature are about the same, and that these temperatures remain about constant from stage to stage. While there are slight appreciations in the temperature of the solution due to the exothermic reaction, which takes place in each stage, these variations 15.
are not meaningful and are accounted for by the use of the term "about".
The absorption process is repeated in each of stages 2 and 3, the process gas, if any, with the remaining acid gas, passing along line 9 to each successive stage. The feed gas inlet temperature for each stage is in the range set forth for stage 1, i.e., each successive inlet temperature is normally from about 30 to about 43C. In sum, the pressure is about the same in each stage; the lean solution enters each stage at about the same temperature; the outlet temperature of the feed gas for each stage is within the range set forth above for the outlet temperature of stage l; the ratio, by volume, of carbon dioxide to hydrogen sulfide in the feed gas increases with each successive stage, the magnitude of the increase depending on the selectivity of the process; and the rich solution outlet temperature for each stage also falls within the range set for the first stage. As previously stated, the aforesaid temperatures of the lean solution and feed gas passed to each stage, as well as the temperatures of the rich solution and treated feed gas as they exit each stage, are normally within the range from about 30C.
to about 43C. It is to be understood, however, that included within the scope of the present invention, is operation of the stages of the absorption zone at lower temperatures (i.e., below 30C.), thereby favoring further increase in selectivity of hydrogen sulfide removal consistent with the kinetics of the system. Therefore, under circumstances wherein added 16.
1()7453~
operating costs due to cooling of the feed gas and lean solution introduced to each stage can be tolerated, the respective temperatures of the lean solution and feed gas fed to each stage can be lowered down to about 10C. in which event the other temperatures referred to above (i.e., the temperature of the treated feed gas and rich solution as they exit each stage), will be correspondingly as low.
The combined rich solution in the common line to which rich solution from each stage is passed, passes through a heat exchanger and then into a conventional stripping zone where it enters, generally, at or near .-the top tray. The rich solution stripper inlet temperature is generally in the range of about 82C. to about 110C.
As noted, the lines, the heat exchanger and the stripper are conventional and not shown in the drawing.
The acid gas and some water are removed from the rich solution in the stripper by distillation. The stripper can be operated by using lower pressure andtor direct heating. Direct heating generates steam internally from the water in the rich solution and can be accomplished by passing lean solution (i.e., bottoms) through reboilers and recycling into the stripper. A
mixture of acid gas and water vapor exit from the top of the stripper. There are approximately 1 to 5 moles of water per mole of acid gas in the overhead. The water can then be removed by condensation and the acid gas recovered by conventional means. All or part of the water may be recycled to the stripper as reflux, ~074534 the preferred mode being to recycle sufficient water to provide the correct molality for the lean solution as noted hereinafter. It should be noted that the water in the stripper has a variety of origins, i.e., feed gas, aqueous alkanolamine solution, and reflux water.
The stripper can be operated at a pressure in the range of about atmospheric to about 65 pounds per square inch absolute (psia) and is generally operated in the range of about 25 psia to about 35 psia and normally at a lower pressure than the absorber. The lean solution leaves the stripper at a stripper bottoms outlet temperature in the range of about 100C. to about 135C. and usually about 118C.
The "lean solution loading" is the ratio of moles of acid gas per mole of alkanolamine in the lean solution and can be about nil to about 0.1, is preferably nil to about 0.05, and is normally about 0.02.
The lean solution then passes from the stripping zone through a common line which proceeds through a heat exchanger and then branches off to lines 11, 13, and 15, which, respectively, supply stages 1, 2, and 3 of the absorption zone with lean solution.
The heat exchanger is known as a lean-rich heat exchanger because the rich solution passing through it is heated up prior to its entry into the stripper and the lean solution is cooled prior to its entry into the stages of the absorber.
It is recommended that the rich solution be filtered after it leaves an absorber stage and that 18.
circulating pumps and/or turbines be used at points along the various lines to maintain the desired circulation rate.
In commercial operations there are losses in the system due to amine entrainment and vaporization, water entrainment, amine degradation, and spillage. These are conventional problems which do not affect the operation of the overall process and will not be treated here.
The lean solution flow rate (or circulation rate) is determined for each stage and is adjusted so that: ;
Ci) from about 0.1 to about 0.9, and normally from about 0.3 to about 0.7, by volume, of hydrogen sulfide passing through each stage is absorbed by the lean solution passing through said stage; and (ii) the rich solution loading at about the bottom of each stage is from about 0.1 to about 0.3 mole of acid gas per mole of alkanolamine, thereby allowing for the recovery of gaseous product, as stripping zone overhead, in which the combined molar ratio of carbon dioxide to hydrogen sulfide is no greater than about 6, and is preferably no greater than about 3.
As a rule of thumb, high rich loadings within the aforesaid range are achieved by reducing the flow rate, and low rich loadings within the aforesaid range are achieved by increasing the flow rate.
Conventional analytical techniques are used to determine amounts of various components in subject process.
It should be understood that the realization of the aforesaid molar ratios of 6 or less which is an objective and advantage of the presently described process, is not necessarily achieved or approached in each stage, but is a combined ratio achieved by the absorption zone 19.
since the rich solution from all stages is combined, as noted, and treated in toto in the stripper. This is important because, when the process deals in low ppm hydro-gen sulfide areas, excellent results are achieved when compared with other processes even though molar ratios may reach even 10 for a particular stage.
The "combined molar ratio" is defined as the molar ratio of carbon dioxide to hydrogen sulfide in the stripper overhead at the end of a complete cycle, i.e., when all the steps of the process have been carried out and all of the conditions are met in all stages, and the rich solution from each stage is being fed to the stripper.
It is not an average ratio for each stage.
The present invention is further illustrated by the following examples:
EXAMPLES 1 to 4 The examples are conducted according to the flow sheet in the drawing and the preferred steps and conditions in the specification.
Each absorber stage is a sieve tray contactor with sieve trays without weirs having a tray spacing of 0.15 meter. The absorber diameter is 3.6 inches.
The still or stripping zone is a packed tower, which is the equivalent of a 17 tray tower. External to the base of the still is one reboiler and external to the top of the still are a condenser, water separator, and pump.
There is a filter and lean-rich heat exchanger in the rich solution common line. There is a pump and the same lean-rich heat exchanger in the lean solution common line.
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107453~
The alkanolamine used as the absorbent is methyl-diethanolamine ~MDEA).
The molality of the lean solution of MDEA enter-ing each stage of the absorber is 8.4, corresponding to an initial MDEA solution which is about fifty percent by .-weight MDEA and fifty percent by weight water.
The apparatus is made of stainless steel (AISI
types 304 and 316~.
In examples l to 3, the feed gas is a mixture of nitrogen, carbon dioxide, hydrogen sulfide and water vapor. The feed gas is saturated and contains essentially no impurities. The nitrogen is used to simulate a process gas, which, in commercial operation, will be a mixture of one or more hydrocarbons and/or other gases such as, for example, a mixture of CH4, C0 and H2, which are inert under the process conditions. In examples 1 to 3, the concentration of carbon dioxide in the feed gas is 15 per-cent by volume, based on the total volume of feed gas exclusive of water vapor and impurities. The molar ratio of C02 to H2S in the feed gas employed in examples 1 to 3 is about 41:1.
In example 4, the feed gas is an acid gas mix-ture of carbon dioxide and hydrogen sulfide and, unlike the feed gas employed in examples 1 to 3, is not diluted with nitrogen. This feed gas is also saturated and contains essentially no impurities. Further, in example 4, the con-centration of carbon dioxide in the feed gas is 94 percent by volume and the concentration of hydrogen sulfide is 6 percent by volume, exclusive of water vapor and impurities.
The molar ratio of C2 to H2S in the feed gas to example 4 is about 16:1.
21.
1()7~534 In each example, the feed gas inlet and outlet temperatures and the lean solution inlet and rich solution outlet temperatures for each stage are about 32C.
The stripper is operated at 30 psia. The rich solution stripper inlet temperature is 104C, the lean solution stripper outlet temperature is 118C, the stripper reboiler temperature is also 118C, and the stripper over- ;
head temperature is 100C. Conventional analytical tech-niques are used to determine amounts of various components.
In examples 1 to 3, the absorber has three stages, the number of trays in each stage being as given in Table I hereinbelow. In example 4, the absorber has six stages and the number of trays per stage is two.
Other test conditions and results are also as set forth in Table I which follows.
ooo~o~
D O ~ O O O Lr~ D O u~ o c~ ~1 ~ c~
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~ ~ U~OO I I 1 000 1 1 1 ~ I
O ~.~ ~, 000 ~0~ 0 000 1 1 1 000 1 1 1 ~0000 1 1 1 O ~ ~, 000 O G co Ul r~l ~ r--l 1~ ~ 0 0 0 1 1 1 0 0 0 1 1 1 ~ ~D ~1 I I I
O O I I I O O 1~ 1 1 1 _I \--1 ~I I I I
O ~1 ~1 000 !~ .
~ O
~
i~i h U~
+
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~ ~ ~ O ~ 0 O O ~ ~ o tn P ~ ~ o ~ a~
o ~ O C)-r~ O O
O O rl ~q O O ~ O J~
n o ~ J~ U ~ J ~ O ~ JJ U ~ J~ ~ ~ ~ J- ~ J J~
X
W~ ~ ~ P~
~074534 ~u~ oooooo oo~ ~ oooooo U~ :
~ ~ o o o ~ , , . . . . . .
,,~ C~l C~ ooo , , o o o .. .. . .
c~l~a~ ~ ~1 ooo _, . . .
`J~O I I I O 000 1 1 1 ., , , . . . . ,_ o ~ C~ Io o o ~4 E~
U~
.,1 a P~ ~
U~ ~
Q) O
o o ~ U~
~ o O ~
O O o ~d~
.,, ~o o ,,~
,~
. ^ O o ~ cq O J~ C~
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a) ~d O
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24 .
~074S34 It is evident from the above data, that the process of the present invention allows for the prefer-ential removal of hydrogen sulfide from carbon dioxide-containing feed gas streams including feed gas containing ~ery dilute concentrations of hydrogen sulfide, and that the hydrogen sulfide-containing gaseous product which is recovered is suitable for further industrial processing such as for the production of elemental sulfur.
1(~74~3~
the selectivity is such that molar ratios of carbon dioxide to hydrogen sulfide approaching about 6, and preferably about 3, or less are realized for the overall process while maintaining other process requirements at levels equal to or better than competitive processes.
Other objects and advantages of the present invention will become apparent from the following descrip-tion and disclosure.
SUMMARY OF THE INVENTION
The present invention provides improvement upon a continuous process for the selective absorption of hydro-gen sulfide from a feed gas comprising an acid gas mixture of carbon dioxide and hydrogen sulfide, said process comprising the steps of (I) counter-currently contacting the feed gas in an absorption zone with a lean aqueous alkanolamine solution to provide a rich aqueous alkanolamine solution, (II) introducing the rich solution into a stripping zone to provide a mixture of acid gas and water vapor overhead and lean solution as bottoms, and ~III) recycling the lean solution to the absorption zone, the said improvement comprising:
(a) using an alkanolamine having the structural formula, R
R - N
R
where R is an alkanol radical of 2 or 3 carbon atoms and is unsubstituted or methyl-substituted, or an alkyl radical having l to 5 carbon atoms provided that at least one R
is an alkanol radical;
(b) using a lean solution which has a molality of about 1.5 to about 75 and a maximum loading of about 0.1 mole of acid gas per mole of alkanolamine;
(c) using an absorption zone having 2 to lO
separate stages, wherein the equilibrium approach between the hydrogen sulfide in the gas and liquid phases is maxi-mized and the equilibrium approach between the carbon dioxide in the gas and liquid phases is minimized;
(d) adjusting the flow rate of the lean solution through each stage so that ~i) from about 0.1 to about 0.9, by volume, of the hydrogen sulfide passing through each stage is absorbed by the lean solution passing through said stage, and (ii) the rich solution loading in each stage is fro~ about 0.1 to about 0.3 mole of acid gas per mole of alkanolamine; and (e) introducing lean solution from the stripping zone into about the top of each stage, removing rich solution from about the bottom of ea~h stage and introducing rich solution into the stripping zone.
Recoverable as stripping zone overhead is gaseous product in which the molar ratio of carbon dioxide to hydrogen sulfide is no greate.r than about 6.
The combination of conditions employed in accord-ance with the teachings of this invention allows for essentially complete removal from the feed gas of hydrogen sulfide in high selectivity relative to carbon dioxide, in that the conditions are such that displacement of absorbed hydrogen sulfide by carbon dioxide in the feed gas is minimized and the driving force for hydrogen sulfide 4.
~074534 9945-1 absorption is maximized, as explained in greater detail hereinbelow.
BRIEF DESCRIPTION OF THE DRAWING
The sole figure of the drawing is a schematic flow diagram of an illustrative embodiment of the present invention.
DETAILED DESCRIPTION AND PREFERRED EMBODIMENTS
As noted above, in addition to the steps and conditions stated heretofore, each stage of the absorption process functions in a particular manner to accomplish the optimum level of selectivity which can be achieved by the presently described process. The function of the selective removal process of this invention encompasses two primary aspects. The first aspect is to maximize the equilibrium approach between hydrogen sulfide in the liquid and gas phases, and minimize the equilibrium approach between the carbon dioxide in the liquid and gas phases. The ability to control these equilibrium approaches is associated with the mass transfer characteristics of the vapor-liquid system. That is, the equilibrium approach between H2S in the gas and liquid phases is enhanced through any of several techniques designed to increase gas phase mass transfer rate. Furthermore, the equilibrium approach between the C2 in the gas and liquid phases is minimized through techniques capable of reducing the mass transfer rate in the liquid phase. In both cases, these effects may be induced through the proper tray design and selection of operating conditions. There are several means of achieving this particular aspect. One practical means, ~74534 9945-1 aLs an example, is to operate with a high superficial feed gas velocity through each stage sufficient to increase gas phase mass transfer rates without significantly increasing the liquid phase mass transfer r~te by maintaining relatively quiescent liquid phase behavior. Another practical means may be utilized in systems already possessing high gas phase mass transfer characteristics. Here, the velocity of feed gas through each stage and tray, or tray equivalent, is inhibited to such an extent that the flow is sufficient but essentially no greater than that flow which is required to maintain the solution on the tray. In so doing, the liquid mass transfer rates and rich solvent acid gas load-ings are reduced, thus further minimizing the equilibrium approach between the C02 in the gas and liquid phases.
A second and equally important aspect of the process of the present invention is to maintain the rich solution loadings in the range of about 0.1 to 0.3 mole of acid gas per mole of alkanolamine. At greater loadings, as the absorption of CO2 and H2S begins to approach nearer and nearer to equilibrium, the CO2 will begin to displace absorbed H2S in view of the fact that carbon dioxide is a stronger acid than hydrogen sulfide. Such displacement is undesirable because selectivity toward net hydrogen sulfide absorption is adversely affected. Furthermore, maintenance of the net solution loadings within the aforesaid range serves to keep a high rate of H2S transfer from the gas phase into the liquid phase by minimizing the increase in the partial pressure of H2S above the solution. To be con-sidered in this latter respect is the fact that the partial pressure of hydrogen sulfide above the solution is influ-1074S~ 9945-1 .
enced by two factors. First, as H2S loading increases, the partial pressure of H2S above the solution rises. Secondly, and importantly, an increase in CO2 loading will also raise the H2S partial pressure above the solution. Therefore, the net driving force between H2S in the gas phase and H2S
in equilibrium with the liquid phase decreases rapidly as loading increases. Multi-staging in accordance with the teachings of the present invention allows one to maximize the driving force throughout the absorber by replacement of the loaded solution at each stage with the introduction of lean solution, with its inherent low partial pressure of both acid gas components. By removing only a fraction of the acid gases per stage and not allowing the rich solution loadings to rise beyond the stated range, high selectivity toward hydrogen sulfide absorption relative to carbon dioxide absorption is obtained.
As noted, the feed gas which is fed to the absorption zone of the process of this invention comprises an acid gas containing carbon dioxide and hydrogen sulfide.
The feed gas may be (A) a mixture of a process gas and the acid gas or (B) the acid gas. The process gas comprises a hydrocarbon or a mixture of hydrocarbons. Examples of hydrocarbons that can be processed in the system are methane, methane which may be in the form of natural gas or substitute or synthetic natural gas (SNG), ethylene, ethane, propylene, propane, mixtures of such hydrocarbons, and the prepurified effluents from the cracking of naphtha or crude oil or from coal gasification. There are no limits to the throughput of feed gas in the subject process provided that the apparatus is sized correctly, a convention-al engineering problem.
7 ~ r 107453~ 9945-1 The feed gas either contains an acid gas together with the process gas or is the acid gas itself. The acid gas is a mixture of carbon dioxide and hydrogen sulfide generally in a molar ratio of about 5 moles to about 140 moles of carbon dioxide per mole of hydrogen sulfide.
The ratio of process gas to acid gas where both are present can cover a very broad range since the process can handle any ratio. This is simply because the process gas is inert insofar as the absorbent is concerned and, aside from throughput sizing, no provision has to be made for it.
Water can be and is usually present in mixture with all of the feed gas components in the form of water vapor or droplets in amounts running from 0 to saturated and is preferably saturated since saturation minimizes water evaporation in the bottom of the absorption zone. An anhydrous feed gas may ~be used but is very rare. The water referred to here is not considered in the determin-ation of molality unless and until it goes into solution with the alkanolamine.
Impurities as defined herein are represented by (a) any gas not defined above as a process gas, acid gas, or water vapor, and (b) solid particles or liquid droplets (exclusive of water droplets) in the feed gas. They can be present in amounts of up to about 3 weight percent based on the total weight of the feed gas and are preferably present in amounts no greater than about l weight percent and, in many cases, lower than 0.01 percent. Examples of the gaseous impurities are sulfur dioxide, carbonyl sulfide, and carbon disulfide. Examples of the solid or liquid 9945-l !
1~)74534 impurities are iron sulfide, iron oxide, high molecular weight hydrocarbons, and polymers. Any olefins having more than one double bond, triple bond hydrocar~ons, and as a general rule, any material that will polymerize or react in situ is an undesirable impurity.
The absorbent is a solution of an alkanolamine and water, the alkanolamine having the structural formula:
R
R - N
R
wherein each R is the same or different and can be at least one alkanol radical, which has 2 or 3 carbon atoms and is unsubstituted or methyl-substituted, or R can be an alkyl radical having 1 to 5 carbon atoms, provided at least one R
is an alkanol radical.
Examples of the alkanolamines are the preferred methyldiethanolamine (MDEA), triethanolamine (TEA), ethyl-diethanolamine (EDEA), tripropanolamine, and triisopropanol-amine. Although mixtures of alkanolamines can be used, they are nst preferred.
There is enough water in the solution or added to the system to provide a molality in the range of about 1.5 to about 75 and preferably from about 5.5 to about 12.5.
The determinatisn of molality is made on the basis of alkanolamine as solute and water as solvent wherein the molality of the solution is equal ,o the number of moles of solute (alkanolamine) dissolved in 1000 grams of solvent ;
(water). The molality in subject process concerns the lean alkanolamine solution which is used to contact the feed gas in each stage of the absorption zone. Other water in the system is not considered in the determination.
9.
1~)7453~
Where aqueous MDEA solution (lean~ is introduced into a system, the concentration thereof is generally from about 10 percent to about 90 percent by weight MDEA based on the weight of the solution and is preferably from about 45 percent ~o about 55 percent by weight. Again, this solution should either provide the correct molality for the process or additional water must be added to the system to do so.
Examples of typical solutions, in percent by weight, are: MDEA 50 percent - water S0 percent; TEA 60 percent - water 40 percent; and EDEA 55 percent - water 45 percent.
Although generally the amount of water for all alkanolamines lies in the range of about 10 to about 90 percent by weight based on the total weight of the solution and the soIution preferably has the proper viscosity for pumping, the amount of water is determined in the end by molality in the ranges set forth above.
The apparatus used in the process for stripping, heat exchanging, and cooling as well as reboilers, filters, piping, turbines, pumps, flash tanks, etc., are of conven-tional design. A typical stripping or regenerator column used in the system can be described as a sieve tray tower having 15 to 20 actual trays or its equivalent in packing.
The stripper contains in its base, or in an external kettle, a tubular heating element or reboiler and at the top of and external to the stripper are condensers and a water separator.
The absorber is not of conventional design, however, as will be seen below.
10 .
Referrin~ to the Drawin~;
-As can be seen in the drawing, the absorber isnot one column, but is divided into separate stages, which are connected to each other in series. The only direct connection between the stages is line 9 through which passes the feed gas. The liquid from one stage never enters another stage unless and until it passes through a still. There can be two to ten stages.
The number of stages is increased within the two to ten stage range in accordance with increased feed gas purification requirements. It is understood that more than ten stages can be used in special cases to achieve the ulti-mate in purity since only a fraction of hydrogen sulfide in the range of about 0.1 to about 0.9, by volume, is removed in each stage, that is, from about 10 to about 90 percent by volume of the hydrogen sulfide content of the feed gas passing through each stage. Normally, the fraction of hydrogen sulfide removed from the feed gas in each stage is within the range from about 0.3 to about 0.7, by volume.
The removal of relatively high (i.e., up to about 0.9) or relatively low (i.e., down to about 0.1) fractions of hydro-gen sulfide passing through each stage depends on the rich-ness and leanness, respectively, of the hydrogen sulfide in the feed gas stream.
Each stage can have one to ten, and usually has no more than six, actual trays or its conventional equiv-alent in spray towers, co-current contactors, or packed towers, for example. In the discussion to follow, it should be understood that the terms "stage" and "tray"
contemplate the use of equivalent apparatus. For example, 11 .
1(~7453~
a stage having three trays is considered to be the same as a packed tower having the packing equivalent to three trays Least preferred are stages having more than four trays. The number of trays is selected by the operator to achieve the result desired in the work specifications with respect to combined molar ratio of carbon dioxide to hydrogen sulfide required in the stripper overhead, which in this process is set at about six or less. As-suming that the same tray design is used for each tray in the stage, the molar ratio of C02 to H2S in the rich solution in each stage will decrease as the number of the trays in the stage decreases, when operating in accordance with the teachings of the invention. Thus, it is clear that a broad range of purities and molar ratios can be achieved by this process.
From the standpoint of apparatus, the process of the present invention may be carried out in a variety of different types of absorbers. Illustrative of suit-able absorbers are those equipped with sieve trays, bubble cap trays, valve trays and the like, having zero or minimum weir heights. These various types of trays may be operated in one or more specific fashions. For example, in absorbers equipped with sieve trays, the flow of feed gas through the tray may be set to such an extent that the flow is sufficient, but essentially no greater than that flow which is required to maintain the solution in the tray. This may be more easily understood by referring to the Chemical Engineers' Handbook, Perry et al., 4th Edition, McGraw-Hill, 1963, Sectîon 18, pages 18-3 to 18-24, inclusive. At page 18-5, it is noted that weeping takes place when the vapor flow is at a rate ~074S34 insufficient to maintain the liquid on the tray. Note also page 18-14 where Perry points out that "dumping" in bubble-cap trays is analogous to "weeping" in sieve trays. The objective here is to provide a tray in which the rate of flow of the vapor is just sufficient to maintain the liquid on the tray, i.e., just above the weeping point. In other words, the tray is made to operate as inefficiently as p~ssible, which must be considered rather unconventional to say the least. Even though the tray is operated in this manner, the tray design is such that the contact efficiency or area between the gas and liquid phases is maximized while minimizing the liquid hold up time on the tray or, in other words, minimizing the contact time between the gas and liquid phases. Thus, operation of a sieve tray close to the "weep-ing point" meets the aforesaid objective.
At this point it should be noted that, although it is desirable to operate the absorber and/or utilize trays or their equivalent such that the contact time between the gas and the liquid phases is as low as possible, thereby taking advantage of the faster rate of reaction between the alkanolamine solution and H2S relative to the rate at which C02 reacts, low contact time is only one factor to be considered in providing a highly selective H2S removal process. Counteracting the initial and in-herently faster rate of absorption of H2S relative to C02, are factors such as the ability of C02 to displace ab-sorbed H2S, the rise in the partial pressure of H2S above the liquid phase as C2 is absorbed, and mass transfer characteristics. Thus, unless the displacement factor as a function of acid gas loading is minimized and the mass transfer driving forces as a function of loading, are such as to favor H2S absorption and, further, unless the S gas phase transfer is enhanced while minimizing C02 :Liquid phase transfer, a reproducibly highly selective H2S
removal process is not achieved. For example, inasmuch as H2S absorption is gas-film controlled whereas C02 ab-sorption is liquid-film controlled, operation at low con-tact time but with a high liquid phase transfer does not provide absorption of H2S in high selectivity.
The feed gas, which generally contains a molar ratio of carbon dioxide to hydrogen sulfide in the range of about 5 to about 140 moles of carbon dioxide per mole of hydrogen sulfide, is introduced at line 9 into stage 1 below the bottom tray (i.e., tray 4 of the drawing), the inlet temperature of the feed gas entering stage 1 being in the range of about 30C. to about 43C. The feed gas flows upwardly through stage 1 to countercurrently meet the aqueous alkanolamine solution referred to as lean solution, i.e., it contains less than about 0.1 mole of acid gas per mole of alkanolamine, which is introduced into stage 1 at tray 1 through line 11.
The pressure in stage 1 is in the range of about atmospheric pressure to about 2000 psia and is normally in the range of about 25 psia to about 1200 psia.
The lean solution enters stage 1 at a temperature which is normally in the range of about 30C. to about 43C.
The feed gas, which has had roughly about half of its hydrogen sulfide absorbed in passing up the stage 1 column, exits through line 9 and passes into stage 2. The stage 1 outlet temperature of the feed gas is normally in 14.
1~)74534 the range of about 30C to about 43C. It is noted that there is very little heat transfer Cor heat loss) from stage to stage.
The gas exiting the first stage contains a ratio, by volume, of carbon dioxide to hydrogen sulfide in the range of about 10 parts to about 280 parts of carbon dioxide per part of hydrogen sulfide, and is generally, saturated if the initial feed gas was saturated.
After the lean solution absorbs acid gas in stage 1, it is referred to as rich solution, i.e., the rich solution is a mixture of lean solution, absorbed acid gas, additional water picked up from the feed gas, and some impurities. The "rich solution loading", which is expressed as the ratio of moles of acid gas to moles of alkanolamine in the rich solution, is, in the first stage, and each stage thereafter, in the range of about 0.1 to about 0.3, .-as measured at about the bottom of each stage where stage outlets 10, 12, and 14 are located. The rich solution r exits the first stage at or below the bottom tray (i.e., ~0 tray 4 of the drawing) at a first stage outlet temperature which is normally in the range of about 30C. to about 43C~ This rich solution then proceeds to a common line, which is also fed by the rich solutions from the other stages 2 and 3 through lines 12 and 14, respectively.
It will be noted that the feed gas outlet temperature and the rich solution outlet temperature are about the same, and that these temperatures remain about constant from stage to stage. While there are slight appreciations in the temperature of the solution due to the exothermic reaction, which takes place in each stage, these variations 15.
are not meaningful and are accounted for by the use of the term "about".
The absorption process is repeated in each of stages 2 and 3, the process gas, if any, with the remaining acid gas, passing along line 9 to each successive stage. The feed gas inlet temperature for each stage is in the range set forth for stage 1, i.e., each successive inlet temperature is normally from about 30 to about 43C. In sum, the pressure is about the same in each stage; the lean solution enters each stage at about the same temperature; the outlet temperature of the feed gas for each stage is within the range set forth above for the outlet temperature of stage l; the ratio, by volume, of carbon dioxide to hydrogen sulfide in the feed gas increases with each successive stage, the magnitude of the increase depending on the selectivity of the process; and the rich solution outlet temperature for each stage also falls within the range set for the first stage. As previously stated, the aforesaid temperatures of the lean solution and feed gas passed to each stage, as well as the temperatures of the rich solution and treated feed gas as they exit each stage, are normally within the range from about 30C.
to about 43C. It is to be understood, however, that included within the scope of the present invention, is operation of the stages of the absorption zone at lower temperatures (i.e., below 30C.), thereby favoring further increase in selectivity of hydrogen sulfide removal consistent with the kinetics of the system. Therefore, under circumstances wherein added 16.
1()7453~
operating costs due to cooling of the feed gas and lean solution introduced to each stage can be tolerated, the respective temperatures of the lean solution and feed gas fed to each stage can be lowered down to about 10C. in which event the other temperatures referred to above (i.e., the temperature of the treated feed gas and rich solution as they exit each stage), will be correspondingly as low.
The combined rich solution in the common line to which rich solution from each stage is passed, passes through a heat exchanger and then into a conventional stripping zone where it enters, generally, at or near .-the top tray. The rich solution stripper inlet temperature is generally in the range of about 82C. to about 110C.
As noted, the lines, the heat exchanger and the stripper are conventional and not shown in the drawing.
The acid gas and some water are removed from the rich solution in the stripper by distillation. The stripper can be operated by using lower pressure andtor direct heating. Direct heating generates steam internally from the water in the rich solution and can be accomplished by passing lean solution (i.e., bottoms) through reboilers and recycling into the stripper. A
mixture of acid gas and water vapor exit from the top of the stripper. There are approximately 1 to 5 moles of water per mole of acid gas in the overhead. The water can then be removed by condensation and the acid gas recovered by conventional means. All or part of the water may be recycled to the stripper as reflux, ~074534 the preferred mode being to recycle sufficient water to provide the correct molality for the lean solution as noted hereinafter. It should be noted that the water in the stripper has a variety of origins, i.e., feed gas, aqueous alkanolamine solution, and reflux water.
The stripper can be operated at a pressure in the range of about atmospheric to about 65 pounds per square inch absolute (psia) and is generally operated in the range of about 25 psia to about 35 psia and normally at a lower pressure than the absorber. The lean solution leaves the stripper at a stripper bottoms outlet temperature in the range of about 100C. to about 135C. and usually about 118C.
The "lean solution loading" is the ratio of moles of acid gas per mole of alkanolamine in the lean solution and can be about nil to about 0.1, is preferably nil to about 0.05, and is normally about 0.02.
The lean solution then passes from the stripping zone through a common line which proceeds through a heat exchanger and then branches off to lines 11, 13, and 15, which, respectively, supply stages 1, 2, and 3 of the absorption zone with lean solution.
The heat exchanger is known as a lean-rich heat exchanger because the rich solution passing through it is heated up prior to its entry into the stripper and the lean solution is cooled prior to its entry into the stages of the absorber.
It is recommended that the rich solution be filtered after it leaves an absorber stage and that 18.
circulating pumps and/or turbines be used at points along the various lines to maintain the desired circulation rate.
In commercial operations there are losses in the system due to amine entrainment and vaporization, water entrainment, amine degradation, and spillage. These are conventional problems which do not affect the operation of the overall process and will not be treated here.
The lean solution flow rate (or circulation rate) is determined for each stage and is adjusted so that: ;
Ci) from about 0.1 to about 0.9, and normally from about 0.3 to about 0.7, by volume, of hydrogen sulfide passing through each stage is absorbed by the lean solution passing through said stage; and (ii) the rich solution loading at about the bottom of each stage is from about 0.1 to about 0.3 mole of acid gas per mole of alkanolamine, thereby allowing for the recovery of gaseous product, as stripping zone overhead, in which the combined molar ratio of carbon dioxide to hydrogen sulfide is no greater than about 6, and is preferably no greater than about 3.
As a rule of thumb, high rich loadings within the aforesaid range are achieved by reducing the flow rate, and low rich loadings within the aforesaid range are achieved by increasing the flow rate.
Conventional analytical techniques are used to determine amounts of various components in subject process.
It should be understood that the realization of the aforesaid molar ratios of 6 or less which is an objective and advantage of the presently described process, is not necessarily achieved or approached in each stage, but is a combined ratio achieved by the absorption zone 19.
since the rich solution from all stages is combined, as noted, and treated in toto in the stripper. This is important because, when the process deals in low ppm hydro-gen sulfide areas, excellent results are achieved when compared with other processes even though molar ratios may reach even 10 for a particular stage.
The "combined molar ratio" is defined as the molar ratio of carbon dioxide to hydrogen sulfide in the stripper overhead at the end of a complete cycle, i.e., when all the steps of the process have been carried out and all of the conditions are met in all stages, and the rich solution from each stage is being fed to the stripper.
It is not an average ratio for each stage.
The present invention is further illustrated by the following examples:
EXAMPLES 1 to 4 The examples are conducted according to the flow sheet in the drawing and the preferred steps and conditions in the specification.
Each absorber stage is a sieve tray contactor with sieve trays without weirs having a tray spacing of 0.15 meter. The absorber diameter is 3.6 inches.
The still or stripping zone is a packed tower, which is the equivalent of a 17 tray tower. External to the base of the still is one reboiler and external to the top of the still are a condenser, water separator, and pump.
There is a filter and lean-rich heat exchanger in the rich solution common line. There is a pump and the same lean-rich heat exchanger in the lean solution common line.
20.
107453~
The alkanolamine used as the absorbent is methyl-diethanolamine ~MDEA).
The molality of the lean solution of MDEA enter-ing each stage of the absorber is 8.4, corresponding to an initial MDEA solution which is about fifty percent by .-weight MDEA and fifty percent by weight water.
The apparatus is made of stainless steel (AISI
types 304 and 316~.
In examples l to 3, the feed gas is a mixture of nitrogen, carbon dioxide, hydrogen sulfide and water vapor. The feed gas is saturated and contains essentially no impurities. The nitrogen is used to simulate a process gas, which, in commercial operation, will be a mixture of one or more hydrocarbons and/or other gases such as, for example, a mixture of CH4, C0 and H2, which are inert under the process conditions. In examples 1 to 3, the concentration of carbon dioxide in the feed gas is 15 per-cent by volume, based on the total volume of feed gas exclusive of water vapor and impurities. The molar ratio of C02 to H2S in the feed gas employed in examples 1 to 3 is about 41:1.
In example 4, the feed gas is an acid gas mix-ture of carbon dioxide and hydrogen sulfide and, unlike the feed gas employed in examples 1 to 3, is not diluted with nitrogen. This feed gas is also saturated and contains essentially no impurities. Further, in example 4, the con-centration of carbon dioxide in the feed gas is 94 percent by volume and the concentration of hydrogen sulfide is 6 percent by volume, exclusive of water vapor and impurities.
The molar ratio of C2 to H2S in the feed gas to example 4 is about 16:1.
21.
1()7~534 In each example, the feed gas inlet and outlet temperatures and the lean solution inlet and rich solution outlet temperatures for each stage are about 32C.
The stripper is operated at 30 psia. The rich solution stripper inlet temperature is 104C, the lean solution stripper outlet temperature is 118C, the stripper reboiler temperature is also 118C, and the stripper over- ;
head temperature is 100C. Conventional analytical tech-niques are used to determine amounts of various components.
In examples 1 to 3, the absorber has three stages, the number of trays in each stage being as given in Table I hereinbelow. In example 4, the absorber has six stages and the number of trays per stage is two.
Other test conditions and results are also as set forth in Table I which follows.
ooo~o~
D O ~ O O O Lr~ D O u~ o c~ ~1 ~ c~
CO ~C~ . . ~C~I~ . . . oooooo oo ooo 0 O O O ~ a I I 1 0 0 0 1 1 1 O ~ C~ I I I
~ ~ U~OO I I 1 000 1 1 1 ~ I
O ~.~ ~, 000 ~0~ 0 000 1 1 1 000 1 1 1 ~0000 1 1 1 O ~ ~, 000 O G co Ul r~l ~ r--l 1~ ~ 0 0 0 1 1 1 0 0 0 1 1 1 ~ ~D ~1 I I I
O O I I I O O 1~ 1 1 1 _I \--1 ~I I I I
O ~1 ~1 000 !~ .
~ O
~
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24 .
~074S34 It is evident from the above data, that the process of the present invention allows for the prefer-ential removal of hydrogen sulfide from carbon dioxide-containing feed gas streams including feed gas containing ~ery dilute concentrations of hydrogen sulfide, and that the hydrogen sulfide-containing gaseous product which is recovered is suitable for further industrial processing such as for the production of elemental sulfur.
Claims (28)
1. In a continuous process for the selective absorption of hydrogen sulfide from a feed gas comprising an acid gas mixture of carbon dioxide and hydrogen sulfide, said process comprising the steps of (I) counter-currently contacting the feed gas in an absorption zone with a lean aqueous alkanolamine solution to provide a rich aqueous alkanolamine solution, (II) introducing the rich solution into a stripping zone to provide a mixture of acid gas and water vapor overhead and lean solution as bottoms, and (III) recycling the lean solution to the absorption zone, the improvement which comprises:
(a) using an alkanolamine having the formula, (R)3N, wherein R is an alkanol radical, which has 2 or 3 carbon atoms and is unsubstituted or methyl-substituted, or an alkyl radical having 1 to 5 carbon atoms, provided at least one of the R groups is an alkanol radical;
(b) using a lean solution, which has a molality of about 1.5 to about 75 and a maximum loading of about 0.1 mole of acid gas per mole of alkanolamine;
(c) using an absorption zone having 2 to 10 separate stages, wherein the equilibrium approach between the hydrogen sulfide in the gas and liquid phases is maxi-mized and the equilibrium approach between the carbon dioxide in the gas and liquid phases is minimized;
(d) adjusting the flow rate of the lean solution through each stage so that (i) from about 0.1 to about 0.9, by volume, of the hydrogen sulfide passing through each stage is absorbed by the lean solution passing through 26.
said stage, and (ii) the rich solution loading in each stage is from about 0.1 to about 0.3 mole of acid gas per mole of alkanolamine; and (e) introducing lean solution from the stripping zone into about the top of each stage, removing rich solu-tion from about the bottom of each stage and introducing said rich solution into the stripping zone.
(a) using an alkanolamine having the formula, (R)3N, wherein R is an alkanol radical, which has 2 or 3 carbon atoms and is unsubstituted or methyl-substituted, or an alkyl radical having 1 to 5 carbon atoms, provided at least one of the R groups is an alkanol radical;
(b) using a lean solution, which has a molality of about 1.5 to about 75 and a maximum loading of about 0.1 mole of acid gas per mole of alkanolamine;
(c) using an absorption zone having 2 to 10 separate stages, wherein the equilibrium approach between the hydrogen sulfide in the gas and liquid phases is maxi-mized and the equilibrium approach between the carbon dioxide in the gas and liquid phases is minimized;
(d) adjusting the flow rate of the lean solution through each stage so that (i) from about 0.1 to about 0.9, by volume, of the hydrogen sulfide passing through each stage is absorbed by the lean solution passing through 26.
said stage, and (ii) the rich solution loading in each stage is from about 0.1 to about 0.3 mole of acid gas per mole of alkanolamine; and (e) introducing lean solution from the stripping zone into about the top of each stage, removing rich solu-tion from about the bottom of each stage and introducing said rich solution into the stripping zone.
2. A process as defined in claim 1 in which the feed gas contains, in addition to the acid gas, a process gas comprising a hydrocarbon or mixture of hydro-carbons.
3. A process as defined in claim 1 in which the flow rate of the lean solution through each stage is adjusted such that no more than about 0.7, by volume, of the hydrogen sulfide passing through each stage is absor-bed by the lean solution passing through each stage.
4. A process as defined in claim 1 in which the alkanolamine is methyldiethanolamine.
5. A process as defined in claim 2 in which the alkanolamine is methyldiethanolamine.
6. A process as defined in claim 3 in which the alkanolamine is methyldiethanolamine.
7. A process as defined in claim 1 in which each stage of the absorption zone has from one to six actual trays or tray equivalents.
27.
27.
8. A process as defined in claim 1 in which each stage of the absorption zone has from one to four actual trays.
9. A process as defined in claim 1 wherein stripping zone overhead is recovered in which the molar ratio of carbon dioxide to hydrogen sulfide is no more than about 6:1.
10. A process as defined in claim 1 wherein stripping zone overhead is recovered in which the molar ratio of carbon dioxide to hydrogen sulfide is no more than about 3:1.
11. In a continuous process for the selective absorption of hydrogen sulfide from a feed gas selected from the group consisting of (A) a mixture of a process gas and an acid gas and (B) an acid gas, where the process gas comprises a hydrocarbon or mixture of hydrocarbons and the acid gas is a mixture of carbon dioxide and hydro-gen sulfide, and the said process comprises the steps of (I) counter-currently contacting the feed gas in an absorption zone with a lean aqueous alkanolamine solution to provide a rich aqueous alkanolamine solution, (II) in-troducing the rich solution into a stripping zone to pro-vide a mixture of acid gas and water vapor overhead and lean solution as bottoms, and (III) recycling the lean solution to the absorption zone, the improvement which comprises:
(a) using an alkanolamine having the formula, (R)3N, wherein each R is the same or different and can be an alkanol radical, which has 2 or 3 carbon atoms and is 28.
unsubstituted or methyl-substituted, or an alkyl radical having 1 to 5 carbon atoms, provided at least one R is an alkanol radical;
(b) using a lean solution having a molality of about 1.5 to about 75 and a maximum loading of about 0.1 mole of acid gas per mole of alkanolamine;
(c) using an absorption zone having 2 to 10 separate stages, wherein the equilibrium approach between the hydrogen sulfide in the gas and liquid phases is maxi-mized while the equilibrium approach between the carbon dioxide in the gas and liquid phases is minimized;
(d) adjusting the flow rate of the lean solution through each stage so that (i) from about 0.3 to about 0.7, by volume, of the hydrogen sulfide passing through said stage is absorbed by the lean solution passing through said stage, and (ii) the rich solution loading at about the bottom of each stage is from about 0.1 to about 0.3 mole of acid gas per mole of alkanolamine;
(e) introducing lean solution from the stripping zone into about the top of each stage and removing rich solution from about the bottom of each stage and intro-ducing said rich solution into the stripping zone, the molar ratio of carbon dioxide to hydrogen sulfide in the stripping zone overhead being no more than about 6:1.
(a) using an alkanolamine having the formula, (R)3N, wherein each R is the same or different and can be an alkanol radical, which has 2 or 3 carbon atoms and is 28.
unsubstituted or methyl-substituted, or an alkyl radical having 1 to 5 carbon atoms, provided at least one R is an alkanol radical;
(b) using a lean solution having a molality of about 1.5 to about 75 and a maximum loading of about 0.1 mole of acid gas per mole of alkanolamine;
(c) using an absorption zone having 2 to 10 separate stages, wherein the equilibrium approach between the hydrogen sulfide in the gas and liquid phases is maxi-mized while the equilibrium approach between the carbon dioxide in the gas and liquid phases is minimized;
(d) adjusting the flow rate of the lean solution through each stage so that (i) from about 0.3 to about 0.7, by volume, of the hydrogen sulfide passing through said stage is absorbed by the lean solution passing through said stage, and (ii) the rich solution loading at about the bottom of each stage is from about 0.1 to about 0.3 mole of acid gas per mole of alkanolamine;
(e) introducing lean solution from the stripping zone into about the top of each stage and removing rich solution from about the bottom of each stage and intro-ducing said rich solution into the stripping zone, the molar ratio of carbon dioxide to hydrogen sulfide in the stripping zone overhead being no more than about 6:1.
12. A process as defined in claim 11 wherein each stage has from 1 to 6 actual trays or tray equivalent.
13. A process as defined in claim 11 wherein each stage has 1 to 4 actual trays.
29.
29.
14. A process as defined in claim 11 wherein the alkanolamine is methyldiethanolamine.
15. A process as defined in claim 12 wherein the alkanolamine is methyldiethanolamine.
16. A process as defined in claim 13 wherein the alkanolamine is methyldiethanolamine.
17. In a continuous process for the selective absorption of hydrogen sulfide from a feed gas selected from the group consisting of (A) a mixture of a process gas and an acid gas and (B) an acid gas, wherein the process gas is a hydrocarbon or a mixture of hydrocarbons and the acid gas is a mixture of carbon dioxide and hydrogen sulfide, and wherein said process comprises the steps of (I) counter-currently contacting the feed gas in an absorption zone with a lean aqueous alkanolamine solution to provide a rich aqueous alkanolamine solution, (II) introducing the rich solution into a stripping zone to provide a mixture of acid gas and water vapor overhead and lean solution as bottoms, and (III) recycling the lean solution to the absorption zone, the improvement which comprises:
(a) using an alkanolamine having the formula, (R3)N, wherein R is an alkanol radical having 2 or 3 carbon atoms, or an alkyl radical having 1 to 5 carbon atoms, pro-vided at least one R is an alkanol radical;
(b) using a lean solution, which has a molality of about 5.5 to about 12.5 and a maximum loading of nil to about 0.05 mole of acid gas per mole of alkanolamine;
30.
(c) using an absorption zone having 2 to 10 separate stages, wherein the equilibrium approach between the hydrogen sulfide in the gas and liquid phases is max-imized while the equilibrium approach between the carbon dioxide in the gas and liquid phases is minimized;
(d) adjusting the flow rate of the lean solution through each stage so that (i) from about 0.3 to about 0.7, by volume, of the hydrogen sulfide passing through each stage is absorbed by the lean solution passing through said stage, and (ii) the rich solution loading at about the bottom of each stage is about 0.1 to about 0.3 mole of acid gas per mole of alkanolamine; and (e) introducing lean solution from the stripping zone into about the top of each stage, removing rich solu-tion from about the bottom of each stage, introducing said rich solution into the stripping zone, and recovering as stripping zone overhead, gaseous product in which the com-bined molar ratio of carbon dioxide to hydrogen sulfide is no greater than about 6:1.
(a) using an alkanolamine having the formula, (R3)N, wherein R is an alkanol radical having 2 or 3 carbon atoms, or an alkyl radical having 1 to 5 carbon atoms, pro-vided at least one R is an alkanol radical;
(b) using a lean solution, which has a molality of about 5.5 to about 12.5 and a maximum loading of nil to about 0.05 mole of acid gas per mole of alkanolamine;
30.
(c) using an absorption zone having 2 to 10 separate stages, wherein the equilibrium approach between the hydrogen sulfide in the gas and liquid phases is max-imized while the equilibrium approach between the carbon dioxide in the gas and liquid phases is minimized;
(d) adjusting the flow rate of the lean solution through each stage so that (i) from about 0.3 to about 0.7, by volume, of the hydrogen sulfide passing through each stage is absorbed by the lean solution passing through said stage, and (ii) the rich solution loading at about the bottom of each stage is about 0.1 to about 0.3 mole of acid gas per mole of alkanolamine; and (e) introducing lean solution from the stripping zone into about the top of each stage, removing rich solu-tion from about the bottom of each stage, introducing said rich solution into the stripping zone, and recovering as stripping zone overhead, gaseous product in which the com-bined molar ratio of carbon dioxide to hydrogen sulfide is no greater than about 6:1.
18. A process as defined in claim 17 wherein the alkanolamine is methyldiethanolamine.
19. A process as defined in claim 18 wherein each stage has no more than six trays or tray equivalents.
20. A process as defined in claim 18 in which each stage has no more than four actual trays.
21. A process as defined in claim 18 in which the combined molar ratio of carbon dioxide to hydrogen sul-fide in the gaseous product recovered as stripping zone overhead is no greater than about 3:1.
31.
31.
22. In a continuous process for the selective absorption of hydrogen sulfide from a feed gas comprising an acid gas mixture of carbon dioxide and hydrogen sulfide, said process comprising the steps of (I) counter-currently contacting the feed gas in an absorption 7-one with a lean aqueous alkanolamine solution to provide a rich aqueous alkanolamine solution, (II) introducing the rich solution into a stripping zone to provide a mixture of acid gas and water vapor overhead and lean solution as bottoms, and (III) recycling the lean solution to the absorption zone, the improvement which comprises:
(a) using an alkanolamine having the formula, (R)3N, wherein R is an alkanol radical having two or three carbon atoms, or an alkyl radical having one to five carbon atoms, provided at least one of the R groups is said alkanol radical;
(b) using a lean solution, which has a molality of about 1.5 to about 75 and a maximum loading of about 0.1 mole of acid gas per mole of alkanolamine;
(c) using an absorption zone having two to ten individual stages wherein, in each stage, the equilibrium approach between the hydrogen sulfide in the gas and liquid phases is maximized and the equilibrium approach between the carbon dioxide in the gas and liquid phases is minimized;
(d) adjusting the flow rate of the lean solution through each stage so that (i) from about 0.1 to about 0.9, by volume, of the hydrogen sulfide passing through each stage is absorbed by the lean solution passing there-through, and (ii) the rich solution loading at about the 32.
bottom of each stage is from about 0.1 to about 0.3 mole of acid gas per mole of alkanolamine;
(e) maintaining the temperature conditions from stage to stage at a substantially constant level;
(f) introducing lean solution from the stripping zone into about the top of each stage, removing rich solu-tion from about the bottom of each stage, and introducing said rich solution into the stripping zone; and (g) withdrawing from the stripping zone, gaseous overhead in which the combined molar ratio of carbon dioxide to hydrogen sulfide is no greater than about 6:1.
(a) using an alkanolamine having the formula, (R)3N, wherein R is an alkanol radical having two or three carbon atoms, or an alkyl radical having one to five carbon atoms, provided at least one of the R groups is said alkanol radical;
(b) using a lean solution, which has a molality of about 1.5 to about 75 and a maximum loading of about 0.1 mole of acid gas per mole of alkanolamine;
(c) using an absorption zone having two to ten individual stages wherein, in each stage, the equilibrium approach between the hydrogen sulfide in the gas and liquid phases is maximized and the equilibrium approach between the carbon dioxide in the gas and liquid phases is minimized;
(d) adjusting the flow rate of the lean solution through each stage so that (i) from about 0.1 to about 0.9, by volume, of the hydrogen sulfide passing through each stage is absorbed by the lean solution passing there-through, and (ii) the rich solution loading at about the 32.
bottom of each stage is from about 0.1 to about 0.3 mole of acid gas per mole of alkanolamine;
(e) maintaining the temperature conditions from stage to stage at a substantially constant level;
(f) introducing lean solution from the stripping zone into about the top of each stage, removing rich solu-tion from about the bottom of each stage, and introducing said rich solution into the stripping zone; and (g) withdrawing from the stripping zone, gaseous overhead in which the combined molar ratio of carbon dioxide to hydrogen sulfide is no greater than about 6:1.
23. A process as defined in claim 22 in which the lean solution is introduced to each stage of the absorption zone at about the same temperature.
24. A process as defined in claim 22 in which the flow rate of the lean solution through each stage of the absorption zone is adjusted such that no more than about 0.7, by volume, of the hydrogen sulfide passing through each stage is absorbed by the lean solution passing through each stage.
25. A process as defined in claim 22 in which each stage of the absorption zone has from one to six actual trays or tray equivalents.
26. A process as defined in claim 22 in which the contact area of the gas and liquid phases is maximized and the contact time between the gas and liquid phases is minimized.
33.
33.
27. A process as defined in claim 22 in which the feed gas contains, in addition to the acid gas, a process gas comprising a hydrocarbon or mixture of hydro-carbons.
28. A process as defined in claim 22 in which the combined molar ratio of carbon dioxide to hydrogen sulfide in gaseous product withdrawn from the stripping zone is no greater than about 3:1.
34.
34.
Applications Claiming Priority (2)
| Application Number | Priority Date | Filing Date | Title |
|---|---|---|---|
| US62721175A | 1975-10-30 | 1975-10-30 | |
| US05/723,161 US4093701A (en) | 1975-10-30 | 1976-09-17 | Process for acid gas removal |
Publications (1)
| Publication Number | Publication Date |
|---|---|
| CA1074534A true CA1074534A (en) | 1980-04-01 |
Family
ID=27090369
Family Applications (1)
| Application Number | Title | Priority Date | Filing Date |
|---|---|---|---|
| CA263,738A Expired CA1074534A (en) | 1975-10-30 | 1976-10-20 | Process for acid gas removal |
Country Status (25)
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| JP (1) | JPS5254680A (en) |
| AR (1) | AR208638A1 (en) |
| AT (1) | AT366929B (en) |
| BR (1) | BR7607246A (en) |
| CA (1) | CA1074534A (en) |
| CS (1) | CS199650B2 (en) |
| DD (1) | DD130570A5 (en) |
| DE (1) | DE2649719C2 (en) |
| DK (1) | DK490776A (en) |
| ES (1) | ES452862A1 (en) |
| FI (1) | FI763095A (en) |
| FR (1) | FR2329328A1 (en) |
| GB (1) | GB1561492A (en) |
| GR (1) | GR63150B (en) |
| IN (1) | IN146105B (en) |
| IT (1) | IT1070933B (en) |
| MX (1) | MX145154A (en) |
| MY (1) | MY8100134A (en) |
| NL (1) | NL186435C (en) |
| NO (1) | NO763710L (en) |
| PH (1) | PH14111A (en) |
| PT (1) | PT65778B (en) |
| RO (1) | RO72255A (en) |
| SE (1) | SE424815B (en) |
| TR (1) | TR19515A (en) |
Families Citing this family (6)
| Publication number | Priority date | Publication date | Assignee | Title |
|---|---|---|---|---|
| JPS53108880A (en) * | 1977-03-05 | 1978-09-22 | Idemitsu Kosan Co Ltd | Washing method for gas with amine process |
| NL190316C (en) * | 1978-03-22 | 1994-01-17 | Shell Int Research | METHOD FOR REMOVING ACID GASES FROM A GAS MIX |
| FR2439613A1 (en) * | 1978-10-27 | 1980-05-23 | Elf Aquitaine | PROCESS FOR THE SELECTIVE EXTRACTION OF H2S FROM GAS MIXTURES CONTAINING CO2 |
| FI107171B (en) * | 1997-10-22 | 2001-06-15 | Metso Paper Inc | Paper machine / board machine suction roll suction box sealing structure |
| JP6117027B2 (en) | 2013-07-04 | 2017-04-19 | 株式会社神戸製鋼所 | Absorption method and apparatus using fine flow path |
| US9962644B2 (en) * | 2015-12-28 | 2018-05-08 | Exxonmobil Research And Engineering Company | Process for increased selectivity and capacity for hydrogen sulfide capture from acid gases |
Family Cites Families (4)
| Publication number | Priority date | Publication date | Assignee | Title |
|---|---|---|---|---|
| US3563695A (en) * | 1968-03-22 | 1971-02-16 | Field And Epes | Separation of co2 and h2s from gas mixtures |
| DE1903065A1 (en) * | 1969-01-22 | 1970-08-27 | Basf Ag | Process for removing carbon dioxide from gas mixtures |
| NL168548C (en) * | 1970-07-15 | 1982-04-16 | Shell Int Research | PROCESS FOR THE SELECTIVE REMOVAL OF HYDROGEN SULFUR FROM GASES CONTAINING SULFUR HYDROGEN AND CARBON DIOXIDE. |
| US4085192A (en) * | 1973-05-11 | 1978-04-18 | Shell Oil Company | Selective removal of hydrogen sulfide from gaseous mixtures |
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1976
- 1976-10-11 SE SE7611274A patent/SE424815B/en not_active IP Right Cessation
- 1976-10-20 CA CA263,738A patent/CA1074534A/en not_active Expired
- 1976-10-27 TR TR19515A patent/TR19515A/en unknown
- 1976-10-29 IN IN1972/CAL/76A patent/IN146105B/en unknown
- 1976-10-29 PH PH19066A patent/PH14111A/en unknown
- 1976-10-29 GB GB44983/76A patent/GB1561492A/en not_active Expired
- 1976-10-29 ES ES452862A patent/ES452862A1/en not_active Expired
- 1976-10-29 JP JP51129595A patent/JPS5254680A/en active Granted
- 1976-10-29 AT AT0805476A patent/AT366929B/en not_active IP Right Cessation
- 1976-10-29 GR GR52033A patent/GR63150B/en unknown
- 1976-10-29 NO NO763710A patent/NO763710L/no unknown
- 1976-10-29 FI FI763095A patent/FI763095A/fi not_active Application Discontinuation
- 1976-10-29 DK DK490776A patent/DK490776A/en unknown
- 1976-10-29 BR BR7607246A patent/BR7607246A/en unknown
- 1976-10-29 PT PT65778A patent/PT65778B/en unknown
- 1976-10-29 NL NLAANVRAGE7612016,A patent/NL186435C/en not_active IP Right Cessation
- 1976-10-29 MX MX166853A patent/MX145154A/en unknown
- 1976-10-29 CS CS766993A patent/CS199650B2/en unknown
- 1976-10-29 IT IT28906/76A patent/IT1070933B/en active
- 1976-10-29 FR FR7632797A patent/FR2329328A1/en active Granted
- 1976-10-29 DE DE2649719A patent/DE2649719C2/en not_active Expired
- 1976-10-29 RO RO7688261A patent/RO72255A/en unknown
- 1976-10-29 AR AR265281A patent/AR208638A1/en active
- 1976-11-01 DD DD7600195545A patent/DD130570A5/en unknown
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1981
- 1981-12-30 MY MY134/81A patent/MY8100134A/en unknown
Also Published As
| Publication number | Publication date |
|---|---|
| FR2329328B1 (en) | 1980-06-13 |
| FR2329328A1 (en) | 1977-05-27 |
| SE424815B (en) | 1982-08-16 |
| DE2649719C2 (en) | 1985-01-10 |
| CS199650B2 (en) | 1980-07-31 |
| MY8100134A (en) | 1981-12-31 |
| TR19515A (en) | 1979-06-27 |
| PT65778B (en) | 1978-04-27 |
| AT366929B (en) | 1982-05-25 |
| SE7611274L (en) | 1977-05-01 |
| ES452862A1 (en) | 1978-03-16 |
| RO72255A (en) | 1982-05-10 |
| PT65778A (en) | 1976-11-01 |
| IT1070933B (en) | 1985-04-02 |
| DK490776A (en) | 1977-05-01 |
| JPS5254680A (en) | 1977-05-04 |
| GB1561492A (en) | 1980-02-20 |
| ATA805476A (en) | 1981-10-15 |
| DD130570A5 (en) | 1978-04-12 |
| GR63150B (en) | 1979-09-25 |
| IN146105B (en) | 1979-02-24 |
| NO763710L (en) | 1977-05-03 |
| DE2649719A1 (en) | 1977-05-12 |
| MX145154A (en) | 1982-01-12 |
| FI763095A (en) | 1977-05-01 |
| BR7607246A (en) | 1977-09-13 |
| PH14111A (en) | 1981-02-26 |
| AR208638A1 (en) | 1977-02-15 |
| NL7612016A (en) | 1977-05-03 |
| NL186435B (en) | 1990-07-02 |
| JPS6111658B2 (en) | 1986-04-04 |
| NL186435C (en) | 1990-12-03 |
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