CS199650B2 - Continuous method for the selective absorption of hydrogen sulphide from dosed gas - Google Patents

Abstract

PROCESS FOR ACID GAS REMOVAL 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 equilibrium 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 recovery 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

(54) Continuous method of selective absorption of hydrogen sulphide from the feed gas

The invention relates to a continuous process for the selective absorption of hydrogen sulfide from various gas mixtures.

The selective removal of hydrogen sulfide from gases containing carbon dioxide by absorption is an important, albeit highly specialized, field of industrial technology. Hydrogen sulfide is particularly suitable for the production of elemental sulfur, but to be used efficiently, the hydrogen sulfide content must be such that the molar ratio of carbon dioxide to hydrogen sulfide is close to 6 and preferably not greater than 6. For some applications, for example in the Claus unit, Since the conventional molar ratio in typical materials such as natural gas and synthetic natural gas is about 5 to 140, it is generally necessary to reduce this ratio in order to utilize the hydrogen sulfide thus obtained. That is, the selectivity of any operationally acceptable absorption process must be such that a high proportion of carbon dioxide passes through the absorber while a small proportion of hydrogen sulfide behaves the same. From a competitive point of view, the most advantageous process is that which is most selective with the same or smaller apparatus and energy requirements than other industrial processes.

It is therefore an object of the present invention to provide a method for selectively removing hydrogen sulfide from gases containing carbon dioxide, the selectivity of which is such that the molar ratio of carbon dioxide to hydrogen sulfide is close to 6 and preferably 3 or even less throughout the duration of the process parameters at the same or better level than known methods.

The invention relates to a continuous process for the selective absorption of hydrogen sulfide from a feed gas containing an acidic gas which is a mixture of carbon dioxide and hydrogen sulphide and optionally a process gas which is a hydrocarbon or a hydrocarbon mixture, characterized in that the feed gas is introduced into an absorption zone consisting from 2 to 10 separate stages, an aqueous alkanolamine solution of formula I is fed to each stage upstream of the stream

R

R — N

R (I), where is

R is a C 2 -C 3 alkanol group optionally substituted by a methyl or C 1 -C 5 alkyl group, provided that at least one R is an alkanol group with a molarity of 1.5 to 75 and a maximum content of 0.1 mole of acid gas per mole of alkanolamine at a temperature of up to 45 ° C and at a pressure in the range of atmospheric pressure to 14 MPa, wherein the flow rate of the aqueous alkanolamine solution is adjusted so that it is absorbed in each step from 0.1 to 10 bar. 0.9 parts by volume of hydrogen sulfide and 0.1 to 0.3 moles of acid gas per mole of alkanolamine, the aqueous alkanolamine solution from the bottom of each stage is introduced into the stripping zone to separate the oxygen gas.

A gaseous product is obtained from the head of the stripping zone in which the molar ratio of carbon dioxide to hydrogen sulfide is at most 6.

The combination of the conditions in the process according to the invention results in substantially complete removal of hydrogen sulfide from the feed gas under high selectivity with respect to carbon dioxide. Conditions are created under which the transfer of absorbed hydrogen sulfide through the carbon dioxide in the feed gas is minimized and the driving force for hydrogen sulfide absorption is maximized.

As mentioned above, in addition to the above operations and conditions, each absorption stage operates separately in order to achieve the optimum selectivity obtainable by the method of the invention. The function of the selective absorption process involves two aspects. First of all, it is necessary to maximize the equilibrium between hydrogen sulfide in liquid. and vapor phase and minimize the approach to equilibrium between the liquid and vapor phase carbon dioxide. The ability to control the approach to equilibrium is associated with ^cha: rakterlstikami mass transfer of vapor - liquid. This means that approaching the equilibrium state between the hydrogen sulphide in the gas and liquid phases is promoted in any way. to increase the degree of mass transfer in the gas phase. In addition, the approximation to the equilibrium between carbon dioxide in the gas and liquid phases is minimized by a technique suitable to reduce the degree of mass transfer into the liquid phase. In both cases, these phenomena can be induced by a suitable tray design and a suitable choice of working conditions. There are various options to achieve this goal. For example, it is practical to work with a higher surface velocity of the feed gas by each stage sufficient to increase the rate of mass transfer in the gas phase without substantially increasing the rate of mass transfer of the liquid phase while maintaining the liquid phase at relative rest. Other practical options can be utilized in a system with already high gas phase mass transfer characteristics. In this case, the rate of feed gas at each stage and tray or tray equivalent is inhibited to such an extent that the flow rate is sufficient but not substantially greater than the flow rate required to maintain the solution on the tray. In such a case, the rate of mass transfer in the liquid state and the saturation of the rich gas with hydrogen sulfide decreases, thus further minimizing the approach to equilibrium of carbon dioxide in the gas and liquid phases.

A second and equally important aspect of the process of the invention is to maintain a saturated solution in the art of about 0.1 to 0.3 moles of acid gas per mole of alkanolamine. With greater saturation. the absorption of carbon dioxide and hydrogen sulfide is increasingly close to equilibrium, and carbon dioxide begins to displace the absorbed hydrogen sulfide, since carbon dioxide is a stronger acid than hydrogen sulfide. This extrusion is undesirable because this affects selectivity with respect to the absorption of pure hydrogen sulfide. In addition, maintaining the saturation of the solution within this range assists a high degree of hydrogen sulfide transfer from the gas phase to the liquid phase by minimizing the increase in hydrogen sulfide partial pressure above the solution. In this respect, it is an important fact that the partial pressure of hydrogen sulfide above the solution is influenced by two factors. In particular, if the hydrogen sulfide saturation increases, the hydrogen sulfide partial pressure increases above the solution. Furthermore, it is important that increasing carbonation saturation also increases the hydrogen sulfide partial pressure above the solution. Thus, the net propulsion power between the gaseous hydrogen sulphide and the hydrogen phase equilibrium with the liquid phase rapidly decreases with increasing saturation. The several stages of the invention maximize the propulsion force by the absorber by replacing the saturated solution at each stage with a lean solution with lower partial pressure of both. By removing only the fraction of the acid gas in the step and avoiding saturation of the rich solution above said limit, a high selectivity with respect to the absorption of hydrogen sulfide as compared to the absorption of carbon dioxide is achieved.

As already mentioned, the feed gas to be introduced into the absorption zone of the process of the invention is an acid gas containing carbon dioxide and hydrogen sulfide. The feed gas may be (A) a process gas / acid gas mixture and (B) an acid gas. The process gas comprises a hydrocarbon or a mixture of hydrocarbons. Hydrocarbons which may be included in the system are methane, methane which is in the form of natural gas or a substitute or synthetic natural gas (SNG), ethylene, ethane, propylene, propane, mixtures of such hydrocarbons and pretreated products from petroleum cracking or crude oil from coal gasification. There are no limits to the throughput of the feed gas in the process of the invention, provided that:. the unit is large enough, which is a common technological problem. _

The feed gas either contains the acid gas together with the process gas or contains the acid gas itself. The acid gas is a mixture of carbon dioxide and hydrogen sulfide generally in a molar ratio of 5 mol to 140 moles of carbon dioxide per mole of hydrogen sulfide.

The ratio of process gas to acid gas can be very wide since the process is suitable for any ratio. This is because the process gas is. inert with respect to the absorbent and need not be considered in terms of enforcement.

The water may be and is generally present in a mixture with the other components of the feed gas in the form of water vapor or water droplets in an amount of 0 until saturation, and the mixture is preferably saturated because saturation minimizes water evaporation at the bottom of the absorption zone. Anhydrous feed gas may also be used, but this is rare. The water referred to in this context is not included in the determination of the molarity when it comes into solution with an alkanolamine.

The impurities present are (a) any gas not yet defined as process gas, acid gas or water vapor, and (b) solid particles or liquid droplets (except water droplets) in the feed gas. The impurities may be present in an amount of up to 3% by weight, based on the total weight of the feed gas, and preferably present in no more than 1% by weight, in many cases present in an amount of less than 0.01%. As examples of gaseous. impurities include sulfur dioxide, carbonyl sulfide and carbon disulfide. Examples of solid and liquid impurities are iron sulfide, iron oxide, high molecular weight hydrocarbons and polymers. The undesirable impurity is any olefin having more than one double bond, triple bonded hydrocarbons, and generally any polymerizing or in situ reactive material.

The absorbent is a solution of alkanolamine and water, the alkanolamine having the general formula

R i

R — N,

Wherein R is each the same or different is at least one C 2 -C 3 alkanol unsubstituted or substituted by methyl, or R is C 1 -C 5 alkyl, provided that at least one R is alkanol.

Of the alkanolamines, methyldiethanolamine (MDEA), triethanolamine (TEA), ethyldiethanolamine (EDEA), tripropanolamine, and triisopropanolamine are particularly preferred. Although mixtures of alkanolamines may be used, this is not preferred.

There is either enough water in the system or water is added to the system to achieve a molarity of 1.5 to 75 and preferably 5.5 to 12.5. The molarity is determined with respect to alkanolamine as the solute and water as the solvent, the molarity of the solution being equal to the number of moles of solute (alkanolamine) dissolved in 1000 g of solvent (water). The molarity of the process of the invention is directed to a lean alkanolamine solution that is used to contact the feed gas at each stage of the absorption zone. Other water is not considered in the system.

When an aqueous (lean) methyldiethanolamine solution is introduced into the system, its concentration is generally 10 to 90% by weight of methyldiethanolamine calculated on the weight of the solution, and is preferably 45 to 55% by weight. Thus, the solution must either have the correct molarity for carrying out the process of the invention or additional water must be added to the system.

Examples of typical solution compositions in weight% are:

N / a methyldiethanolamine and 50% water; 60% triethanolamine and 40% water;

% of ethyldiethanolamine and 45% water.

However, in general, the amount of water for all alkanolamines. the idea is 10 to 90% by weight, based on the total weight of the solution, and the solution preferably has a suitable viscosity with respect to it. for pumping, the amount of molybdenum water in the above-mentioned field is determined. ,

The apparatus used in the method of the invention for stripping, heat exchange and cooling as well as digesters, filters, pipes, turbines, pumps, expansion tanks, etc. are common designs. A typical stripping or regeneration column used in the process of the invention may be characterized as a sieve tray tower having 15 to 20 actual trays or an equivalent packing. The striper has a tubular heating element or a bottom reboiler at its bottom or in the outside pin and a condenser and water separator in or outside the head.

The absorber is not of conventional construction, as will be mentioned.

As is apparent from the drawing, the absorber is not a column but a column. is divided into separate stages which are connected in series. The only direct connection between the stages is the conduit 9 through which the feed gas passes. . Liquid from one stage is never introduced into the other stage, except when passing through a distillation apparatus. ₍ The unit can be two to ten degrees. '

The number of stages is two to ten in accordance with the increasing requirements for cleaning the feed gas. Of course, more than ten stages can be used in special cases to achieve extraordinary purity, since in each stage only the hydrogen sulfide fraction in the range of 0.1 to 0.9 parts by volume is removed, representing 10 to 90 volume percent hydrogen sulfide contained in the feed gas. passing through each stage.

Normal amount of hydrogen sulphide removed from. of the feed gas in each step is in the range of 0.3 to 0.7 parts by volume. The removal of a relatively high fraction (e.g. up to about 0.9 parts) or a relatively low fraction (e.g. less than 0.1 part) of hydrogen sulfide as it passes through the steps depends on the high or low hydrogen sulfide content of the feed gas to be introduced.

Each stage has one to ten and, as a rule, no more than six actual trays or their common equivalent in spray towers, downstream contactors or filled towers. In the following, "grade" and "floor" are considered equivalent terms. For example, a stage having three levels is considered equivalent to a tower having a filling equivalent to three levels. Steps having more than four levels are least preferred. The number of trays is selected by the operator to achieve the desired operating parameters with respect to the combination of the molar ratio of carbon dioxide to hydrogen sulfide required in the stripper head, which is set to six or less in the process of the invention. Assuming that the same tray design is used for all of the trays in the step, the molar ratio of carbon dioxide to hydrogen sulfide in the rich solution will decrease in each step with the decreasing number of trays in the step of the process of the invention. Thus, it is clear that a wide variation in purity and molar ratio can be achieved in the process of the invention.

With regard to the device, the process according to the invention can be carried out in a wide variety of absorbers. Examples of suitable absorbers are those equipped with sieve trays, cap trays and flaps trays and similar trays with zero or minimum overflow height. These different types of floors can work in one or more specific ways. For example, in absorbers equipped with sieve trays, the tray flow can be adjusted to be. sufficient but not more than the flow needed to maintain the solution on the tray. This requirement is further clarified in Chemical Engineers' Handbook, Perry et al.,. 4th edition, McGraw-Hill, 1963, section 18, pp. 18-3 to 18-24. On page 18-5, water leakage occurs when steam flows at a rate insufficient to hold liquid on the palate. It is also recalled on pages 18-14, where Perry states that emptying on the hat trays is analogous to the release on sieve trays. There is a need to provide a tray on which the flow rate of the steam is just sufficient to hold the liquid on the tray, i.e. just above the point of discharge. In other words, the tray must work as inefficiently as possible, which seems unusual. In order for the tray to work in this way, the tray construction must be such that the contact efficiency or the surface area is limited! the liquid and gaseous phases were maximal in the minimum liquid residence time per tray or, in other words, in. minimum contact times. between the liquid and gaseous phases. Under this condition, the sieve tray close to the 'leakage point' meets the above condition.

It should be recalled that, although it is desirable for the absorber to operate and / or utilize the trays or their equivalent, the contact time between the gaseous and liquid phases is kept as low as possible, thereby utilizing a quicker reaction between the alkanolamine solution and hydrogen sulfide. with respect to the reaction of alkanolamine with carbon dioxide, it is low. contact time with only one factor decisive for a highly selective method of hydrogen sulphide removal. Factors such as the ability of oxide counteract the initial and substantially faster absorption of hydrogen sulfide compared to carbon dioxide. carbon dioxide to displace the absorbed hydrogen sulphide, increase the hydrogen sulphide partial pressure above the liquid phase during the absorption of carbon dioxide, and mass transfer characteristics. Thus, if displacement as a function of acid gas saturation is kept to a minimum and the driving force of mass transfer as a function of saturation is such that. furthermore, if the gaseous transfer of hydrogen sulfide is promoted while keeping the carbon dioxide transport in the liquid phase to a minimum, reproducibly high selective removal of hydrogen sulfide is not achieved. For example, since the absorption of hydrogen sulfide is controlled in the gas film, while the absorption of carbon dioxide is controlled in the liquid film, operation at low contact time, but at high liquid phase transfer, results in highly selective absorption of hydrogen sulfide.

The feed gas, in which the mole ratio of carbon dioxide to hydrogen sulfide in the range of 5 to 140 moles of carbon dioxide per mole of hydrogen sulfide, is generally introduced through the duct 9 to stage 1 below the bottom of the tray (e.g. tray 4 of the drawing). the feed gas entering step 1 is 30 to 43 ° C. The feed gas flows upwardly through stage 1 countercurrently with respect to an aqueous alkanolamine solution referred to as a lean solution containing, for example, less than 0.1 mol of acid gas per mole of alkanolamine introduced into stage 1 on the first floor via line 11.

The overpressure in step 1 is in the range of atmospheric pressure to 14.0 MPa and normally the overpressure is 0.17 to 8.4 MPa.

The poor solution exits to step 1 at a temperature of generally 30 to 43 ° C.

The feed gas, whose roughly half the hydrogen sulfide content is absorbed as it passes through stage 1 of the column, exits through line 9 and is introduced. The outlet temperature of the feed gas from Step 1 is normally 30 to 43 ° C. It is recalled that there is a slight heat transfer (and therefore a slight heat loss) from one stage to another.

The gas leaving the first stage has a fine carbon dioxide to hydrogen sulfide ratio in the range of 10 to 200 parts carbon dioxide per part hydrogen sulfide and is generally saturated if the initial feed gas has been saturated.

After the acid gas is absorbed by the lean solution in step 1, the lean solution is referred to as a rich solution, i.e., the rich solution is a mixture of lean solution, absorbed acid gas, additional water entrained from the feed gas and any impurities. The "rich solution saturation", expressed as the ratio of moles of acid gas to moles of alkanolamine in the rich solution, is 0.1 to 0.3 in the first stage and in each subsequent stage when measured at the bottom of each stage where they are The rich solution exits from the first stage at or below the bottom tray (tray 4 in the figure) at the outlet temperature of the first stage, which is normally 30 to 43 ° C. This rich solution then goes through a common line to which also rich solutions from the other stages 2 and 3 are supplied via lines 12 and 14. It is recalled that the outlet temperature of the feed gas and the outlet temperature of the rich solution are almost the same and that these temperatures remain constant throughout degrees. Although there is a slight increase in solution temperature due to the exothermic reaction occurring at each stage, these changes are not substantial and it can be said that the temperature is still substantially constant.

The absorption process is repeated in each of stages 2 and 3, the process gas, when used, with the remaining acid gas flowing through line 9 from one stage to the second stage. The inlet temperature of the feed gas for each stage is in the range of temperatures indicated for stage 1, i.e. each additional inlet temperature is about 30 to

43 ° C. Likewise, the pressure in each stage is quite the same; the lean solution enters each stage at approximately the same temperature; the outlet temperature of the feed gas for each stage is within the above-mentioned range for the outlet temperature of stage 1; the volume ratio of carbon dioxide to hydrogen sulfide in the feed gas increases in each subsequent stage, the magnitude of which depends on the selectivity of the process; also the outlet temperature for each stage is in the field indicated for stage 1. As already mentioned, said lean solution and feed gas temperatures to each stage, as well as the rich solution and treated feed gas temperatures at their exit from each stage are normally within the range 30. to 43 ° C. However, the process of the invention also includes operation in the absorption zone stages at lower temperatures (i.e. below 30 ° C), which promotes a further increase in selectivity of hydrogen sulfide removal in relation to the system kinetics. Thus, in circumstances where the additional operating costs of cooling the feed gas and the lean solution introduced into each stage can be tolerated, the temperatures of the lean solution and the feed gas to each stage may be lowered to below about 10 ° C, already discussed (for example, the temperature of the treated feed gas and the temperature of the rich solution leaving each stage) are reduced accordingly.

The combined rich solution in a common conduit into which the rich solution from each stage is passed is passed through a heat exchanger and introduced into a conventional stripping zone which generally enters at or near the uppermost floor. The inlet temperature of the rich solution into the stripper is generally 82 to 110 ° C.

As already mentioned, the pipe, heat exchanger and stripper are common and are therefore not shown in FIG.

In the stripper, the acid gas and some water are separated from the rich solution by distillation. The striper may operate using lower pressure and / or direct heating. Direct heating generates steam internally from the water contained in the rich solution and can be carried out by passing a lean solution (for example, the bottom through the bottom digester and recycling to the stripper. The acid gas / water vapor exits the stripper head. All or some or part of the water can be recycled to the stripper as reflux, it is preferable to recycle enough water to maintain the correct molarity for the lean solution, as mentioned above. The water in the stripper is of different origin, such as water from the feed gas, aqueous alkanolamine solution and reflux water.

The striper may operate at atmospheric pressure to 0.46 kg / cm ² and typically operates at 0.17 to 0.24 kg / cm ² and normally at a lower pressure than that of the absorber. The poor solution leaves the stripper at the bottom of the stripper at an outlet temperature of about 100-135 ° C and generally at a temperature of about 118 ° C.

"Lean solution saturation" means the ratio of moles of acid gas per mole of alkanolamine in lean solution and may be zero to 0.1 in the art, preferably zero to 0.05 and normally 0.02.

The poor solution then proceeds from the stripping zone through a common conduit that passes through the heat exchanger and then branches to the conduits 11, 13 and 15, which supply stages 1, 2 and 3 of the absorption zone with the lean solution. The heat exchanger is known as a lean-rich heat exchanger because the passing rich solution passing through the exchanger is heated before entering the stripper and the lean solution is cooled before entering the absorber stage.

It is recommended that the rich solution be filtered prior to entering the absorber stage and that circulation pumps or turbines be used at points along the various pipelines to maintain the desired circulation rate.

In industrial operations. there are system losses due to amine entrainment and evaporation, water entrainment, amine degradation and overflow. These are common problems that do not affect the operation and are therefore not discussed in detail here.

The flow rate or circulation rate of the lean solution is determined for each stage and is adjusted as follows: (i) about 0.1 to 0.9 and normally about 0.3 to 0.7 parts by volume of hydrogen sulfide passing through each stage is absorbed by the lean solution passing through the respective stage degree; and (ii) saturating the rich solution at the bottom. of each step is about 0.1 to 0.3 moles of acid gas per mole of alkanolamine, allowing the recovery of the gaseous product as. product from the stripping head. a zone in which the combined molar ratio of carbon dioxide to hydrogen sulphide is not more than 6 and preferably not more than 3.

As a rule, high saturation of the rich solution within the above limits is achieved by reducing the flow rate, and low saturation of the rich solution is achieved within this range by increasing the flow rate.

It is used . conventional analytical methods for determining the amounts of the various components in the method of the invention.

It is pointed out that achieving said molar ratio of 6 or even lower is an object and advantage of the process of the invention and this value does not need to be reached or approached at each stage, but a combined ratio of the absorption zone is achieved as the rich solution is combined. from all stages and processed at once in the stripper. This is important because performing the process at low ppm hydrogen sulfide contents achieves excellent results compared to other processes even though the molar ratio is achieved at 10 separate steps.

'Combined molar ratio' means the molar ratio of carbon dioxide to hydrogen sulphide in the stripper head at the end of the complete cycle, that is to say, after carrying out all the stages of the process and fulfilling all conditions at all stages and at. introducing a rich solution from each stage. into the stripper. It is not a mean ratio for each grade.

Way . according to the invention the attached examples are illustrated.

Example 1 to. 4

The examples are carried out in the manner shown schematically in FIG. and in view of the preferred embodiment and conditions of the process according to the invention.

Each absorption stage is a contactor with sieve trays without overflows with a plate distance of 0.15 m. The absorber diameter is 9 cm.

The distillation or stripping zone is a tower with a charge corresponding to the tower. se 17. patry. Outside the bottom of the distillation. Apparatus is the bottom digester and outside the head of the distillation apparatus is a condenser, water separator and pump.

The device includes a lean filter and heat exchanger. - Rich in piping for rich solution. The equipment includes a pump and the same poor - rich heat exchanger in a lean solution pipe.

The alkanolamine used as the absorbent is methyldiethanolamine (MDEA).

The molarity of the lean methyldiethanolamine solution entering each absorber stage is 8.4, which means that the initial methyldiethanolamine solution contains 50 wt% methyldiethanolamine and 50 wt% water.

The apparatus is made of. stainless steel (AISI types 304 and 316).

In the case of samples 1 to 3, the feed gas is a mixture of nitrogen, oxide. carbon dioxide, hydrogen sulphide and water vapor. The feed gas is saturated and essentially free of impurities. Nitrogen is used to simulate a process gas, which in industrial plants is a mixture of one or more hydrocarbons and / or other gases, such as a mixture of methane, carbon monoxide and hydrogen, which is inert under operating conditions. In the examples 1 to 3, the concentration of carbon dioxide in the feed gas was 15% by volume, based on the total feed gas volume excluding water vapor and impurities. The molar ratio of carbon dioxide to hydrogen sulfide in the feed gas of Examples 1 to 3 was 41: 1.

In the case of Example 4, the feed gas is an acidic gaseous mixture of carbon dioxide and hydrogen sulfide and, unlike the feed gas used in Examples 1 to 3, is not diluted with nitrogen. The metering gas is also saturated and substantially free of impurities. In the case of Example 4, the concentration of carbon dioxide in. 94% by volume of the feed gas and 6% by volume of hydrogen sulfide. percent except water vapor and impurities. The molar ratio of carbon dioxide to hydrogen sulfide in the feed gas of Example 4 is 16: 1.

In each example, the inlet and outlet temperatures of the feed gas and the inlet temperature of the lean solution and the outlet temperature of the rich solution for each stage are 32 ° C.

The strlper operates at a pressure of 0.21 MPa, at an inlet temperature of the rich solution of 104 ° C, at an outlet temperature of the lean solution of 118 ° C, at the bottom boiler temperature also of 118 ° C and. at a head temperature of 100 ° C. To determine the quantity. The various components are used by conventional techniques.

In the case of Examples 1 to 3, the absorber has three stages, the number of plates in each stage being shown in the table. In the case of Example 4, the absorber has six stages and there are two levels in each stage.

The other test conditions and results are given in Table 1.

199850

Ta bulka I Example 1 2 3 4 entering carbon dioxide into. absorber 15 Dec 15 Dec 15 Dec 94 % By volume entering hydrogen sulphide into the absorber, 0.3650 0.3650 0.3650 6 % vol overpressure in each stage, MPa 0.98 1.97 3.93 0.07 dosing gas rate, liter / minute (1) 708 944 1420 651 rate of lean solution entry of each absorber grade cm ³ / minute 355 700 550 800 Hydrogen sulfide input (volume): degree 1 3650 ppm 3650 ppm 3650 ppm 6.0% degree 2 1500 ppm 1700. ppm 1600 ppm 4,0 ''% Grade 3 500 ppm 650 ppm 500 ppm 2.5% degree 4 - - - 1.5 o / o Grade 5 - - - 0.75 «/ o Grade 6 - - 0.36% Hydrogen sulfide output (volume) degree 1 1500 ppm 1700 ppm 1600 ppm 4.0% degree 2 500 ppm 650 ppm 500 ppm 2.5 o / ₀ Grade 3 170 ppm 260 ppm 200 ppm 1,5 0/0 degree 4 - . - 0,75 '% Grade 5 - - -. 0.36% Grade 6 - - ' 0,10% Output of rich solution, mole of oxide carbon dioxide + hydrogen sulfide. per mol methyldi- ethanolamine degree 1 0.14 0,23 ' ⁴ 0.20 0.20 degree 2 0.16 0.18 0.16 0.18 Grade 3 0.11 0.18 0.12 0.21 Grade 4 , - - - 0.20 Grade 5 - - -. 0.13 Grade 6 - - - 0.12 Stripper head, mole of carbon dioxide per mole of hydrogen sulphide: degree 1 1.4 2.5 1.5 0.28 degree 2 4.1 3.3 3.5 0.46 Grade 3 10.0 9.4 6.5 1,8 degree 4 - - - - 2.2 Grade 5 - - - 2.5 Grade 6 ^z - - - 4,5 combined 3.0 3.5 2.6 1’1 Number of floors: degree 1 2 1 2 2 degree 2 4 1 2 2 Grade 3 2 1 1 2 (2) Saturation of lean solution on entry, mole of carbon dioxide hydrogen sulfide per mole of methyldiethanolamine - degree 1 0,035 . 0.03 0.015 0.02 degree. 2 0.015 0,04 / 0.02 0.02 Grade 3 0.015 0.03 0.02 0.02 degree 4 - - - 0.02 Grade 5 . - -. - 0.02 Grade 6 . - - - 0.02 (11 Under standard temperature and pressure (STP) conditions (2) In each of the remaining steps 4, 5 and 6 of Example 4, the number of trays is also two

and that the gaseous hydrogen sulphide-containing product obtained is suitable for further industrial processing such as the production of elemental sulfur. . '

From the above figures, it is clear that the process of the invention is advantageous for the removal of hydrogen sulfide from a feed gas based on oxide. carbon dioxide, which contains very

Claims (8)

Subject CLAIMS 1. A continuous process for the selective absorption of hydrogen sulfide from a feed gas containing an acidic gas which is a mixture of carbon dioxide and hydrogen sulphide and optionally a process gas which is a hydrocarbon or a hydrocarbon mixture, characterized in that the feed gas is introduced into 10 separate stages, each with a countercurrent aqueous alkanolamine solution being fed to each stage. of formula I R R — N R (I), where is R is a C 2 -C 3 alkanol group optionally substituted by a methyl or C 1 -C 5 alkyl group, provided that at least one R is an alkanol group with a molarity of 1.5 to 75 and a maximum content of 0 1 mol of acid gas per mole of alkanolamine at a temperature of up to 45 ° C and at a pressure in the range of atmospheric pressure to 14 MPa, the flow rate of the aqueous alkanolamine solution. is adjusted so that 0.1 to 0.9 parts by volume of hydrogen sulphide and 0.1 to 0.3 moles of acid gas per mole of alkanolamine are absorbed in each stage, the aqueous alkanolamine solution from the bottom of each stage being introduced into the stripping zone. separation of oxygen gas. VYNALEZU 2. The process according to claim 1, wherein the aqueous alkanolamine solution is recirculated to the individual stages from the stripping zone. 3. A process according to claim 1, wherein the feed gas is contacted countercurrently with an aqueous solution of methyldiethanolamine. 4. A method according to claim 1 or 2, characterized in that the recycled cold solution is introduced into the absorption zone, each stage having 1 to 6 actual trays or tray equivalents. 5. A method according to claim 1 or 2, characterized in that the recycled solution is introduced into an absorption zone, each stage having 1 to 4 actual trays or tray equivalents. 6. A process according to claim 2, wherein the flow rate of the aqueous alkanolamine solution is adjusted so that 0.7 parts by volume of hydrogen sulfide is absorbed at each stage. 7. The process of claim 2, wherein the feed gas is contacted with an aqueous solution of an alkanolamine of formula I wherein R is as hereinbefore defined with a molarity of 5.5 to 18.5 and a maximum saturation of 0.005 mol of acid gas. per mole of alkanolamine. 8. The process of claims 2 to 7, wherein the flow rate of the aqueous alkanolamine solution is adjusted to absorb 0.3 to 0.7 parts by volume of hydrogen sulfide by each stage.