a technical report on gas sweetening

A Technical Report on

Gas Sweetening by Amines

Subhasish Mitra, Sr. Process Engineer Petrofac Engineering (I) Ltd, Mumbai, India

Gas sweetening by amine

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Content

1.0 Introduction 3

2.0 Gas sweetening basics 7

3.0 Alkanolamine gas treatment basics 8

4.0 Alkanolamine gas treating chemistry 13

5.0 Alkanolamine processes-strengths and weakness/solvent selection 20

6.0 Amine system description 24

7.0 Operational issues of amine sweetening system 33

8.0 Troubleshooting guide 41

9.0 Prevention of BTEX emission 46

10.0 Bulk CO2 removal technology by membrane unit 47

11.0 New developments 49

Appendix - 1: Typical process specification for gas sweetening package 50

Appendix - 2: Typical process flow sheet for amine absorption unit prepared in 58 Hysys simulator package


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List of abbreviation

AGR Acid Gas Removal

BTX Benzene Toluene Xylene

DEA Di Ethyl Amine

DGA Di Glycol Amine Agent

DIPA Di Iso-Propanol Amine

HSS Heat Stable Salts

LNG Liquefied Natural Gas

LPG Liquefied Petroleum Gas

MDEA Methyl Di Ethyl Amine

MEA Mono Ethyl Amine

SRU Sulphur Recovery Unit

TEA Tri Ethyl Amine

VLE Vapour Liquid Equilibrium

VOC Volatile Organic Compound


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1.0 Introduction:

The use of natural gas as an industrial and domestic fuel has become a prime source of energy generation. There are a number of processes utilized between the wellhead and the consumer to render the natural gas fit for consumption. These processes are vital for removal of .contaminants. within the gas stream which, if left in the gas, would cause problems with freezing, corrosion, erosion, plugging, environmental, health and safety hazards. Contaminants can be generalized as mentioned in Table 1,

Table 1. Principal gas phase impurities

Hydrogen sulfide (H2S) Carbon di-oxide (CO2) Water vapor (H2O) Sulfur di-oxide (SO2) Nitrogen Oxides (NOX) VOC Volatile Chlorine Compounds (HCl,Cl2 etc) Volatile fluorine compounds (HF, SiF4 etc.) Basic Nitrogen Compounds Carbon Mono-oxide Carbonyl Sulfide Carbon di-sulfide Organic sulfur compounds Hydrogen cyanide

As consumption of natural gas as an inevitable fuel is increasing worldwide, gas treating is getting more complex due to emissions requirements established by environmental regulatory agencies. Upstream gas preconditioning, or final steps for gas conditioning downstream of the gas-treating unit, are emerging as the best options to comply with the most stringent regulations emerging in the industry. The final steps of gas conditioning are a combination of different processes to remove impurities such as elemental sulphur, solids, heavy hydrocarbons and mercaptans.

Table 2: Typical product specifications

In general, gas purification involves the removal of vapor-phase impurities from gas streams. The processes which have been developed to accomplish gas purification vary from simple once-through wash operations to complex multiple-step recycle systems. In many cases, the


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process complexities arise from the need for recovery of the impurity or reuse of the material employed to remove it. The primary operation of gas purification processes generally falls into one of the following five categories:

1. Absorption into a liquid 2. Adsorption on a solid 3. Permeation through a membrane 4. Chemical conversion to another compound 5. Condensation

Absorption:

It refers to the transfer of a component of a gas phase to a liquid phase in which it is soluble. Stripping is exactly the reverse-the transfer of a component from a liquid phase in which it is dissolved to a gas phase. Absorption is undoubtedly the single most important operation of gas purification processes and is used widely..

Adsorption:

It is the selective concentration of one or more components of a gas at the surface of a micro-porous solid. The mixture of adsorbed components is called the adsorbate, and the micro-porous solid is the adsorbent. The attractive forces holding the adsorbate on the adsorbent are weaker than those of chemical bonds, and the adsorbate can generally be released (desorbed) by raising the temperature or reducing the partial pressure of the component in the gas phase in a manner analogous to the stripping of an absorbed component from solution. When an adsorbed component reacts chemically with the solid, the operation is called chemisorption and desorption is generally not possible.

Membrane permeation:

It is a relatively new technology in the field of gas purification. In this process, polymeric membranes separate gases by selective permeation of one or more gaseous components from one side of a membrane barrier to the other side. The components dissolve in the polymer at one surface and are transported across the membrane as the result of a concentration gradient. The concentration gradient is maintained by a high partial pressure of the key components in the gas on one side of the membrane barrier and a low partial pressure on the other side. Although membrane permeation is still a minor factor in the field of gas purification, it is rapidly finding new applications.

Chemical conversion:

It is the principal operation in a wide variety of processes, including catalytic and non-catalytic gas phase reactions and the reaction of gas phase components with solids. The reaction of gaseous Species with liquids and with solid particles suspended in liquids is considered to be a special case of absorption and is discussed under that subject.

Condensation:


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This process is of interest primarily for the removal of VOCs from exhaust gases. The process consists of simply cooling the gas stream to a temperature at which the Organic compound has a suitably low vapor pressure and collecting the condensate. 2.0 Gas sweetening basics:

Gas sweetening is one of the important purification processes which is employed to remove acidic contaminants from natural gases prior to sale. This includes removal of H2S and CO2 from gas streams by using absorption technology and chemical solvents. Sour gas contains H2S, CO2, H2O, hydrocarbons, COS/CS2, solids, mercaptans, NH3, BTEX, and all other unusual impurities that require additional steps for their removal.

There are many treating processes available however no single process is ideal for all applications. The initial selection of a particular process may be based on feed parameters such as composition, pressure, temperature, and the nature of the impurities, as well as product specifications. The second selection of a particular process may be based on acid/sour gas percent in the feed, whether all CO2, all H2S, or mixed and in what proportion, if CO2 is significant, whether selective process is preferred for the SRU/TGU feed, and reduction of amine unit regeneration duty. The final selection could be based on content of C3 + in the feed gas and the size of the unit (small unit reduces advantage of special solvent and may favor conventional amine). Final selection is ultimately based on process economics, reliability, versatility, and environmental constraints.

Clearly, the selection procedure is not a trivial matter and any tool that provides a reliable mechanism for process design is highly desirable.

Hydrogen sulfide and carbon dioxide removal processes can be grouped into the seven types indicated in Table 3, which also suggests the preferred areas of application for each process type.

Table 3: Selection of treatment process

Both absorption in alkalime solution (e.g., aqueous diethanolamine) and absorption in a physical solvent (e.g., polyethylene glycol dimethyl ether) are suitable process techniques for treating high-volume gas streams containing hydrogen sulfide andor carbon dioxide. However, physical absorption processes are not economically competitive when the acid gas partial pressure is low because the capacity of physical solvents is a strong function of partial


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pressure. Physical absorption is generally favored at acid gas partial pressures above 200 psia, while alkaline solution absorption is favored at lower partial pressures. A lower pressure limit (60 - 100 psia) has also been mentioned in literature above which physical solvents are favored.

Membrane permeation is particularly applicable to the removal of carbon dioxide from high-pressure gas. The process is based on the use of relatively small modules, and an increase in plant capacity is accomplished by simply using proportionately more modules. As a result, the process does not realize the economies of scale and becomes less competitive with absorption processes as the plant size is increased.

At very high acid-gas concentrations (over about 15% carbon dioxide), a hybrid process (amine + membrane) proved to be more economical than either type alone. The hybrid process uses the membrane process for bulk removal of carbon dioxide and the amine process for final cleanup.

When hydrogen sulfide and carbon dioxide are absorbed in alkaline solutions or physical solvents, they are normally evolved during regeneration without undergoing a chemical change. If the regenerator off-gas contains more than about 10 tons per day of sulfur (as hydrogen sulfide), it is usually economical to convert the hydrogen sulfide to elemental sulfur in a conventional Claus-type sulfur plant. For cases that involve smaller quantities of sulfur, because of either a very low concentration in the feed gas or a small quantity of feed gas, direct oxidation may be the preferred route.

Direct oxidation can be accomplished by absorption in a liquid with subsequent oxidation to form slurry of solid sulfur particles or sorption on a solid with or without oxidation. The solid sorption processes are particularly applicable to very small quantities of feed gas where operational simplicity is important, and to the removal of traces of sulfur compounds for final cleanup of synthesis gas streams. Solid sorption processes are also under development for treating high temperature gas streams, which cannot be handled by conventional liquid absorption processes.

Adsorption is a viable option for hydrogen sulfide removal when the amount of sulfur is very small and the gas contains heavier sulfur compounds (such as mercaptans and carbon disulfide) that must also be removed. For adsorption to be the preferred process for carbon dioxide removal there must be a high CO2 partial pressure in the feed, the need for a very low concentration of carbon dioxide in the product, and the presence of other gaseous impurities that can also be removed by the adsorbent.

3.0 Alkanolamine gas treatment basics

The removal of sour or acid gas components such as hydrogen sulfide (H2S), carbon dioxide (CO2), carbonyl sulfide (COS) and mercaptans (RSH) from gas and liquid hydrocarbon streams is a process requirement in many parts of the hydrocarbon processing industry. This is especially true with the increasingly stringent environmental considerations coupled with the need to process natural gas and crude oil with increasingly higher sulfur levels. The chemical solvent process, using the various alkanolamines, is the most widely employed gas treating process.


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These processes utilize a solvent, either an alkanolamine or an alkali-salt (hot carbonate processes) in an aqueous solution, which reacts with the acid gas constituents (H2S and CO2) to form a chemical complex or bond. This complex is subsequently reversed in the regenerator at elevated temperatures and reduced acid gas partial pressures releasing the acid gas and regenerating the solvent for reuse. They are well suited for low operating pressure applications where the acid gas partial pressures are low and low levels of acid gas are desired in the residue gas since their acid gas removal capacity is relatively high and insensitive to acid gas partial pressure as compared to physical solvents. The chemical solvent processes are generally characterized by a relatively high heat of acid gas absorption and require a substantial amount of heat for regeneration. The alkanolamines are widely used in both the natural gas and the refinery gas processing industries treating a wide variety of applications. Figure 1 illustrates the process flow for a typical gas treating plant employing an alkanolamine.

Gas to be purified is passed through an inlet separator and/or a gas-liquid coalescer to remove any entrained liquids or solids, the sour gas is introduced at the bottom of the absorber or contactor. Normally packed or trayed tower is used and the gas is contacted counter-currently with the aqueous amine solution absorbing the acid gas in the amine upward through the absorber, countercurrent to a stream of the solution. The rich solution from the bottom of the absorber is heated by heat exchange with lean solution from the bottom of the stripping column and is then fed to the stripping column at some point near the top. In units treating sour hydrocarbon gases at high pressure, it is customary to flash the rich solution in a flash drum maintained at an intermediate pressure to remove dissolved and entrained hydrocarbons before acid gas stripping. When heavy hydrocarbons condense from the gas stream in the flash drum may be used to skim off liquid hydrocarbons as well as to remove dissolved gases. The flashed gas is often used locally as fuel.

A water wash is used primarily in MEA systems, especially at low absorber operating pressures, as the relatively high vapor pressure of MEA may cause appreciable vaporization losses. The other amines usually have sufficiently low vapor pressures to make water washing unnecessary, except in rare cases when the purified gas is used in a catalytic process and the catalyst is sensitive even to traces of amine vapors. If acid gas condensate from the regenerator reflux drum (contains water) is used for this purpose, no draw-off tray is required because it is necessary to readmit this water to the system at some point. It should be noted however, that this condensate is saturated with acid gas at regenerator condenser operating conditions and that this dissolved acid gas will be reintroduced into the gas stream if the water is used as it is for washing. If the gas volume is very large, compared to the amount of wash water, this may be of no consequence. However, if calculations indicate that the quantity of acid gas so introduced is excessive, a water stripper can be included in the process. Alternatively, a recirculating water wash with a dedicated water wash pump can be utilized. This design uses a comparatively small wash water make-up and wash water purge. The number of trays used for water wash varies from two to five in commercial installations. An efficiency of 40 or 50% per tray has been reported in literature under typical absorber operating conditions. From this, it would appear that four trays would be ample to remove over 80% of the vaporized amine from the purified gas and, incidentally, a major portion of the amine carried as entrained droplets in the gas stream. It is probable that even greater tray efficiency is obtained in the water wash section of the stripping column. However, because of the higher temperature involved, the amine content of the vapors entering this section may be quite high. Four to six trays are commonly used for this service.


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A small packed tower with a lean amine wash may be installed on top of the flash drum to remove H2S from the flashed gas if sweet fuel gas is required. Lean solution from the stripper, after partial cooling in the lean-to-rich solution heat exchanger, is further cooled by heat exchange with water or air, and fed into the top of the absorber to complete the cycle. Acid gas that is removed from the solution in the stripping column is cooled to condense a major portion of the water vapor. This condensate is continually fed back to the system to prevent the amine solution from becoming progressively more concentrated. Generally, all of this water, or a major portion of it, is fed back to the top of the stripping column at a point above the rich-solution feed and serves to absorb and return amine vapors carried by the acid gas stream.

Many modifications to the basic flow scheme have been proposed to reduce energy consumption or equipment costs. For example, power recovery turbines are sometimes used on large, high-pressure plants to capture some of the energy available when the pressure is reduced on the rich solution. A minor modification aimed at reducing absorber column cost is the use of several lean amine feed points. Most of the lean solution is fed near the midpoint of the absorber to remove the bulk of the acid gas in the lower portion of the unit. Only a small stream of lean solution is needed for final clean-up of the gas in the top portion of the absorber, which can therefore be smaller in diameter. A modification that has been used successfully to increase the acid gas loading of the rich amine (and thereby decrease the required solution flow rate) is the installation of a side cooler (or intercooler) to reduce the temperature inside the absorber. The optimum location for a side cooler is reported to be the point where half the absorption occurs above and half below the cooler, which results in a location near the bottom of the column.

Figure 1. Typical gas sweetening plant PFD

The alkanolamine gas treating basic process flow scheme as presented in Figure 1 has remained relatively unaltered over the years. The principal technological development has been the introduction of additional alkanolamines for use as gas treating solvents. TEA was


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utilized in early applications but was quickly displaced by MEA and DEA as the alkanolamines of principal commercial interest. Other amines of significant commercial importance include DIPA, DIGLYCOLAMINE Agent, 2-(2-aminoethoxy) ethanol, (DGA) and MDEA. Of late, a great deal of interest in formulated MDEA specialty solvents has developed in order to take advantage of MDEAs unique features as a gas treating solvent.

3.1 Amine concentration:

The choice of amine concentration may be quite arbitrary and is usually made on the basis of operating experience. Typical concentrations of MEA range from 12 wt% to a maximum of 32 wt% however it should be noted that higher amine concentrations, up to 32 wt% MEA, may be used when corrosion inhibitors are added to the solution and when CO2 is the only acid gas component. DEA solutions that are used for treatment of refinery gases typically range in concentration from 20 to 25 wt% while concentrations of 25 to 30 wt% are commonly used for natural gas purification. DGA solutions typically contain 40 to 60 wt% amine in water and MDEA solution concentrations may range from 35 to 55 wt%. It is obvious that increasing the amine concentration will generally reduce the required solution circulation rate and therefore the plant cost. However, the effect is not as great as might be expected, the principal reason being that the acid-gas vapor pressure is higher over more concentrated solutions at equivalent acid-gas/amine mole ratios. In addition, when an attempt is made to absorb the same quantity of acid gas in a smaller volume of solution, the heat of reaction results in a greater increase in temperature and a consequently increased acid-gas vapor pressure over the solution.

The effect of increasing the amine concentration in a specific operating plant using DGA solution for the removal of about 15% acid gas from associated gas is shown in Figure 2. The graph indicates that the optimum DGA strength for this case is about 50 wt%. The effect of the increasing amount of DGA at higher concentrations is almost nullified by the decreasing net acid gas absorption per mole of DGA.

Figure2. Effect of DGA conc. on maximum plant capacity and net solution loading


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3.2 Thermal effects:

Considerable heat is released by the absorption and subsequent reaction of the acid gases in the amine solution. A small amount of heat may also be released (or absorbed) by the condensation (or evaporation) of water vapor. To avoid hydrocarbon condensation the lean solution is usually fed into the top of the absorber at a slightly higher temperature than that of the sour gas, which is fed into the bottom. As a result, heat would be transferred from the liquid to the gas even in the absence of acid gas absorption. The heat of reaction is generated in the liquid phase, which raises the liquid temperature and causes further heat transfer to the gas. However, the bulk of the absorption (and therefore heat generation) normally occurs near the bottom of the column, so the gas is first heated by the liquid near the bottom of the column, and then cooled by the incoming lean solution near the top of the column.

When gas streams containing relatively large proportions of acid gases (over about 5%) are purified, the quantity of solution required is normally so large that the purified gas at the top of the column is cooled to within a few degrees of the temperature of the lean solution. In such cases essentially all of the heat of reaction is taken up by the rich solution, which leaves the column at an elevated temperature. This temperature can be calculated by a simple heat balance around the absorber since the temperatures of the lean solution, feed gas, and product gas are known, and the amount of heat released can be estimated from available heat of solution data.

A typical temperature profiles for an absorber (Glycol-amine system, similar profile observed for MEA & DGA plants also) of this type is shown in Figure 3. The temperature bulge is a result of the cool inlet gas absorbing heat from the rich solution at the bottom of the column, and then later losing this heat to the cooler solution near the upper part of the column. The size, shape, and location of the temperature bulge depend upon where in the column the bulk of the acid gas is absorbed, the heat of reaction, and the relative amounts of liquid and gas flowing through the column. In general, for CO2 absorption, the bulge is sharper and lower in the column for primary amines, broader for secondary amines, and very broad for tertiary amines, which absorb CO2 quite slowly and also have a low heat of solution.


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Figure3. Temperature bulge in acid gas absorber

Since heat is transferred from the hot liquid to the cooler gas at the bottom of the column and in the opposite direction near the top, the temperature profiles for gas and liquid cross each other near the temperature bulge. This effect is shown in Figure 4 for an absorber treating 840 psig natural gas containing 7.56% CO2 and a trace of H2S with a 27 wt% DEA solution.

Figure4. Composition & temperature profile in acid gas absorber

System design requirements:

The design of amine plants centers around the absorber, which performs the gas purification step, and the stripping system which must provide adequately regenerated solvent to the absorber. After selecting the amine type and concentration, key items i.e. solution flow rate; absorber and stripper types (tray or packed), absorber and stripper heights and diameters: and the thermal duties (heating and cooling) of all heat transfer equipment are to be appropriately chosen to meet the required product specification.

4.0 Alkanolamine gas treating chemistry

Hydrogen sulfide (H2S) and carbon dioxide (CO2) are called acid gases because in water or an aqueous solution they dissociate to form weak acids. The alkanolamines are weak organic bases. When the sour gas stream containing H2S and/or CO2 is contacted counter-currently with the aqueous alkanolamine solution, the acid gas and the amine base react to form an acid-base complex, a salt. This acid-base complex is reversed in the stripper when the acid gas rich amine is stripped by steam, releasing the acid gas for disposal or further processing


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and regenerating the amine solution for reuse, thus removing the acid gas from the inlet gas stream.

The alkanolamines are classified by the degree of substitution on the central nitrogen; a single substitution denoting a primary amine, a double substitution, a secondary amine, and a triple substitution, a tertiary amine. Each of the alkanolamines has at least one hydroxyl group and one amino group. In general, the hydroxyl group serves to reduce vapor pressure and increase water solubility, while the amine group provides the necessary alkalinity in water solutions to promote the reaction with acid gases. It is readily apparent looking at the molecular structures that the non-fully substituted alkanolamines have hydrogen atoms at the non-substituted valent sites on the central nitrogen, whereas the tertiary amines are fully substituted on the central nitrogen. This structural characteristic plays an important role in the acid gas removal capabilities of the various treating solvents.

Amines which have two hydrogen atoms directly attached to a nitrogen atom, such as MEA and DGA, are called primary amines and are generally the most alkaline. DEA and DPA have one hydrogen atom directly attached to the nitrogen atom and are called secondary amines. TEA and MDEA represent completely substituted ammonia molecules with no hydrogen atoms attached to the nitrogen, and are called tertiary amines.

Primary amines:

Monoethanolamine (MEA) DIGLYCOLAMINE Agent (DGA)

C2H4OH - NH2 HOC2H4OC2H4 - NH2

Secondary amines

Diethanolamine (DEA) Diisopropanolamine (DIPA)

C2H4OH - NH - C2H4OH C3H5OH - NH- C3H5OH

Tertiary amines

Triethanolamine (TEA) Methyldiethanolamine (MDEA)

2H4OH - NH - C2H4OH C2H4OH - NH - C2H4OH


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Figure 5: Structural formulae of Alkanolamines used in gas treating

In an aqueous solution, H2S and CO2 dissociate to form a weakly acidic solution.

Ionization of water:

H2O = H+ + OH-

Ionization of dissolved H2S:

H2S = H+ + HS-

Hydrolysis and ionization of dissolved CO2:

CO2 + H2O = HCO3- + H+

When a gas stream containing H2S and/or CO2 is contacted by an aqueous amine solution, the acid gases react with the amine to form a soluble acid-base complex, a salt, in the treating solution. The reaction between both H2S and CO2 is exothermic and a considerable amount of heat is liberated. Regardless of the structure of the amine, H2S reacts instantaneously with the primary, secondary or tertiary amine via a direct proton transfer reaction as shown in Equation 1 below to form the amine hydrosulfide:

R1R2R3N + H2S R1R2R3NH+ HS - Equation 1

The reaction between the amine and CO2 is a bit more complex because CO2 absorption can occur via two different reaction mechanisms. When dissolved in water, CO2 hydrolyses to form carbonic acid, which in turn, slowly dissociates to bicarbonate. The bicarbonate then undertakes an acid-base reaction with the amine to yield the overall reaction shown by Equation 2 below:

CO2 + H2O H2CO3 (Carbonic Acid) - Equation 2

H2CO3 H+ + HCO3 - (Bicarbonate) - Equation 3


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H+ + R1R2R3N R1R2R3NH+ -Equation 4

CO2 + H2O + R1R2R3N R1R2R3NH+ HCO3 - Equation 5

This acid-base reaction may occur with any of the alkanolamines regardless of the amine structure but it is slow kinetically because the carbonic acid dissociation step to the bicarbonate is relatively slow. A second CO2 reaction mechanism as shown by Equation 3 below requiring the presence of labile hydrogen in the molecular structure of the amine may also occur.

CO2 + R1R2NH R1R2N+ HCOO - - Equation 6

R1R2N+ HCOO- + R1R2NH R1R2NCOO- + R1R2NH2 - Equation 7

CO2 + 2R1R2NH R1R2NH2 + R1R2NCOO- - Equation 8

This second reaction mechanism for CO2, which results in the formation of the amine salt of a substituted carbamic acid, is called the carbamate formation reaction and may only occur with primary and secondary amines. The CO2 reacts with one primary or secondary amine molecule to form the carbamate intermediate which in turn reacts with a second amine molecule to form the amine salt. The rate of CO2 absorption via the carbamate reaction is rapid, much faster than the CO2 hydrolysis reaction, but somewhat slower than the H2S absorption reaction. The stoichiometry of the carbamate reaction indicates that the capacity of the amine solution for CO2 is limited to 0.5 mole of CO2 per mole of amine if the only reaction product is the amine carbamate. But, the carbamate can undergo partial hydrolysis to form bicarbonate, regenerating free amine. Hence CO2 loadings greater than 0.5, as experienced in some plants employing DEA, are possible through the hydrolysis of the carbamate intermediate to bicarbonate. The fact that CO2 absorption may occur by two reaction mechanisms with significantly different kinetic characteristics has a great impact upon the relative absorption rates of H2S and CO2 among the different alkanolamines.

For primary and secondary amines, very little difference exists between the H2S and CO2 reaction rates. This rate equivalence is due to the availability of the rapid carbamate formation reaction for CO2 absorption. Therefore, the primary and secondary amines achieve essentially complete removal of H2S and CO2. However, because the tertiary amines are fully substituted, they can not form the carbamate. Tertiary amines must react with CO2 via the slow CO2 hydrolysis mechanism discussed earlier. For MDEA, since the CO2 reaction with water to form bicarbonate is slow and the H2S reaction is fast, it is generally felt that the H2S reaction is gas phase limited while the CO2 reaction is liquid phase limited. With only the slow acid-base reaction available for CO2 absorption, MDEA and several of the formulated MDEA products yield significant selectivity toward H2S relative to CO2.

A little insight to the solubility phenomenon of acid gases (H2S, CO2) exhibits a physical solubility relationship in aqueous medium. Figure 3 displays a graphical representation of the acid gas reactions with aqueous phase. Here (g) designates the molecule in the vapor phase while (aq) designates the molecule physically dissolved in water. Under these premises, Henrys law can be applied to relate the vapor and physically dissolved liquid concentrations:

iyiP = imiHi (i = H2S, CO2) - Equation 9


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where i = fugacity coefficient of component i yi = mole fraction of component i in vapor phase P = total pressure of the system i = activity coefficient of component i mi = concentration of component i in liquid phase Hi = Henrys constant of component i. -

Figure 6. Acid gas VLE representation

Further acid gas solubility is present in the form of chemically dissolved ions. Since H2S and CO2 are only considered weak acids, very little ionization occurs unless a basic compound (such as an amine) is also present. Taking H2S as an example, the total equivalent H2S in the aqueous phase will be the sum of free physically dissolved H2S, bisulfide ion (HS-), and sulfide ion (S2-).

Water and ammonia/alkanolamines (designated generically as R3N) obey a vapor pressure relationship across the liquid vapor phase boundary. For water the relationship is:

-Equation 9

Within the aqueous phase, a number of acid-base chemical reactions are present as depicted in Figure 1. Table 1 indicates all the primary reactions necessary to model the system along with equilibrium relationships obeyed (equations 3-9). Every equilibrium relationship mentioned in Table 1 can be tried to Hydrogen ion concentration (H+) by the below mentioned thermodynamic relationship,

- Equation 10


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Considering an infinite dilution in essentially aqueous phase at standard conditions followed by substitution of molarity unit the following well known expression is obtained,

- Equation 11

Since hydrogen ion is present everywhere, solution pH plays an important role for modeling the chemistry of this system.

Table 4. Aqueous phase chemical reactions & equilibrium relationships

To understand how pH can alter the ion distribution in a polybasic acid such as H2S in the presence of a weak base such as MDEA, a dilute solution is assumed where activity coefficients () are unity. The total solution H2S and MDEA concentrations are defined to set the material balances:

- Equation 12

The fractional sulphide and amine species concentrations are defined as,

Following relationships are derived based on above data,


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A model derived from the above figures shows that when pH of the aqueous solution is raised i.e. solution is made more basic, the fraction of total H2S present the solution shifts from free physically dissolved H2S to bisulphide (HS-) ions and ultimately to sulphide (S2-) ions. This drives the equilibrium towards dissolving more total H2S. Addition of alkanolamines (basic in nature) as solvent accomplishes this shift (Refer Figure 4). An alternate way to achieve proper absorption of acid gas in scrubbing solvent is to increase partial pressure of acid gas (Vide equation 4) which in turn increases solubility of physically dissolved gas.

Figure 7. Distribution of H2S & MDEA ions v/s pH


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5.0 Alkanolamine processes-Strengths & Weakness/Solvent selection:

5.1 Monoethanolamine (MEA):

The use of MEA in gas treating applications is well established and the subject of a tremendous amount of literature. However, MEA is no longer the predominant gas treating alkanolamine; its use has declined in recent years.

The advantages of MEA include:

Low solvent cost, Good thermal stability, Partial removal of COS and CS2, which requires a reclaimer, and High reactivity due to its primary amine character, a grain H2S specification can usually be achieved and CO2 removal to 100 ppmv for applications at low to moderate operating pressures.

Some of the disadvantages of MEA are:

High solvent vapor pressure which results in higher solvent losses than the other alkanolamines, Higher corrosion potential than other alkanolamines, High energy requirements due to the high heat of reaction with H2S and CO2, Nonselective removal in a mixed acid gas system, and Formation of irreversible degradation products with CO2, COS and CS2, which requires continuous reclaiming.

The MEA-CO2 degradation reaction produces oxazolidone-2, 1-(2-hydroxyethyl) imidazolidone-2, N-(2-hydroxyethyl) ethylenediamine (HEED), and higher polyamines which accelerate corrosion in addition to representing a loss of MEA. In applications where the gas to be treated is at low pressures, and maximum removal of H2S and CO2 is required or no minor contaminants such as COS and CS2 are present, MEA may still have a window of application and should not be overlooked. However, more efficient solvents, particularly for the treatment of high-pressure natural gas are rapidly replacing MEA.

5.2 Diethanolamine (DEA):

Probably the most widely employed gas treating solvent, DEA being a secondary amine is generally less reactive than MEA. Applications with appreciable amounts of COS and CS2, besides H2S and CO2, such as refinery gas streams, can generally be treated successfully.

The advantages of DEA are:

Resistance to degradation from COS and CS2, Low solvent vapor pressure which results in potentially lower solvent losses, Reduced corrosive nature when compared to MEA, and Low solvent cost.


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Some of the disadvantages of DEA include:

Lower reactivity compared to MEA and DGA Agent, Essentially nonselective removal in mixed acid gas systems due to inability to slip an appreciable amount of CO2, Higher circulation requirements, and Non-reclaimable by conventional reclaiming techniques.

Degradation products resulting from the reaction of DEA and CO2 at elevated temperatures include hydroxyethyloxazolidone-1,dihydroxyethylpiperazine,3-(2-ydroxyethyl)oxazolidone-2(HEOD), N,N.bis(2-hydroxyethyl) piperazine (BHEP) and N,N,N-tris(2-hydroxyethyl) ethylenediamine (THEED).

An explanation for DEAs wide utilization within the gas treating industry is due to DEAs ability to balance three key gas treating process considerations,

1) Reactivity, i.e. ability to make specification product. 2) Corrosiveness, generally less than that of MEA. 3) Energy utilization allowing a wider array of gas treating applications than other solvents. di-glycolamine agent (DGA).

5.3 Diglycolamine (DGA):

Being a primary amine, DGA Agent is similar in many respects to MEA except that its lower vapor pressure permits higher solvent concentrations, typically 50 to 60 weight percent, to be utilized, resulting in significantly lower circulation rates and energy utilization. DGA Agent treating units are processing natural gas and refinery gas streams containing from 1.5 to 35.0% total acid gas. Most units are treating gases with both CO2 and H2S with CO2/H2S ratios varying from 300/1 to 0.1/1. Treating pressure covers the entire spectrum from 75 psig to 1,000 psig [517 to 6,985 kPA].

The advantages of DGA Agent include:

Capital and operating cost savings due to lower circulation requirements, Removal of COS and CS2, High reactivity, grain H2S specification can generally be obtained for applications with low operating pressures and high operating temperatures, Enhanced mercaptan removal in comparison to other alkanolamines, Low freeze point for 50 weight percent solution of -30 F [-34.4 C], whereas 15 wt. % MEA and 25 wt. % DEA solutions freeze at 25 and 21 F [-3.9 and -6.1 C], respectively, and Excellent thermal stability. Atmospheric reclaiming to reverse the BHEEU formed by the reaction of DGA with CO2 and COS.

Some of the disadvantages of DGA Agent are:

Nonselective removal in mixed acid gas systems, Absorbs aromatic compounds from inlet gas which potentially complicates the sulfur recovery unit design, Higher solvent cost relative to MEA and DEA.


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DGA Agent reacts with CO2 and COS to form BHEEU, N,N,bis-(hydroxyethoxyethyl) urea, via Equation 1 and with COS and CS2 to form BHEETU, N,N,bis(hydroxyethoxyethyl) thiourea, via Equation 2 as shown below: 2R-NH2 + (CO2 or COS) (R-NH)2CO + (H2O or H2S) 2R-NH2 + (COS or CS2) (R-NH)2CS + (H2O or H2S)

The major chemical by-product in a DGA solution is BHEEU. It is formed by the reaction of two moles of DGA Agent with 1 mole of either CO2 or COS. A second by-product can also form by the reaction of 1 mole of either CS2 or COS with two moles of DGA Agent yielding a thiouera (BHEETU). Experience indicates the dominant reaction with COS will be to form BHEEU. The reactions between CO2, COS, or CS2 and DGA are reversible at temperatures of 340 to 360 F [171.1 to 182.2 C].

5.4 Methyldiethanolamine (MDEA):

In recent years, the specialty formulated MDEA solvents offered by several solvent vendors have gained a significant share of the market. The introduction of the formulated MDEA solvents has been the major innovation within the gas treating industry over the past decade. This commercial success is due principally to the ability of MDEA to selectively remove H2S when treating a gas stream containing both H2S and CO2 while slipping a significant portion of the CO2. This slippage of CO2 can be useful in applications requiring the upgrading of H2S content for sulfur plant feed gas or adjusting the CO2 content of the treated gas while at the same time removing H2S to less than 1/4 grain per 100 scf (4 ppmv). Originally, the most significant application of MDEA and the various formulated MDEA solvents were in tail gas treating units but increasingly the formulated solvents have displaced primary and secondary amines in refinery primary treating systems and in high pressure natural gas applications.

The advantages of MDEA and the formulated MDEA solvents are:

Selectivity of H2S over CO2 in mixed acid gas applications, Essentially complete H2S removal while only a portion of CO2 is removed enriching the acid gas feed to the sulfur recovery unit (SRU), Low vapor pressure which results in potentially lower solvent losses, Less corrosive, High resistance to degradation, and Efficient energy utilization (capital and operating cost savings).

The disadvantages of MDEA and the formulated MDEA solvents include:

Highest solvent cost relative to MEA, DEA and DGA Agent, Lower comparative reactivity, Non-reclaimable by conventional reclaiming techniques, and Minimal COS, CS2 removal.

Although degradation is not normally a problem with MDEA, certain circumstances have shown that MDEA is degradable. TGTU systems are especially vulnerable to degradation from SO2 breakthrough. Not only is a noticeable build-up of Heat-Stable-Salts seen, but MDEA degradation into primary and secondary amines is also likely. Reactions are possible which will lead to the formation of bicine, a strong metal chelate. Corrosion is a major


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concern when degradation products are formed and bicine is present. As with all alkanolamines, the presence of oxygen increases the likelihood of product degradation and corrosion concerns.

Table- 5: Comparative Study of Solvents:

Solvent Name MEA (Mono Ethanol Amine )

DEA (Di- Ethanol Amine)

DGA (Di-Glycol Amine Agent)

MDEA (Methy Di Ethanol Amine)

Solvent Cost Low Solvent Cost Low Solvent Cost Relatively high solvent cost Highest Solvent Cost

Solvent Loss

High solvent vapor pressure results in higher solvent loss.

Low solvent vapor pressure results potentially lower solvent loss.

Low vapor pressure which results in potentially low solvent loss.

Selectivity

Non-selective removal in a mixed acid gas system. Partial removal of COS and CS2.

Non-selective removal in a mixed acid gas system.

Non-selective removal in a mixed acid gas system. Removal of COS and CS2.

Selectivity of H2S over CO2 in mixed acid gas applications. Essentially complete H2S removal while only a portion of CO2 is removed enriching the acid gas feed to the sulfur recovery unit. Minimal COS and CS2 removal.

Thermal Stability

Good Thermal Stability

Excellent Thermal Stability

Reactivity

High reactivity due to its primary amine characteristics.

Low reactivity compared to MEA and DGA Agent.

High reactivity, 1/4 grain H2S specification can generally be obtained for applications with low operating pressures & high operating temperatures.

Lower comparative reactivity

Corrosion Higher Corrosion potential

Reduces corrisive nature compared to MEA. Less corrosive

Recovery (Reclaimation )

Requires continuous reclaiming.

Non-reclaimable by conventional reclaiming techniques.

Non-reclaimable by conventional reclaiming techniques.


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Table 6: Comparative features of various gas sweetening substances:

6.0 Amine System Description:

6.1 Inlet separation / Pre-treatment:

The design and type of inlet separation should be carefully considered. Inlet separation equipment can vary from slug catchers, which are generally designed to catch large slugs of liquids from gas gathering systems where condensing hydrocarbons are prevalent, to cutting edge technology reverse flow filter-coalescers. Experience indicates that inlet feed gas filtration is very important and critical in the trouble-free operation of the amine treating system. The cleaner an amine system is, the better the system operates. Many of the contaminants that cause poor performance can enter the amine system via the inlet feed gas. In most cases, the inlet separator of the amine system is sized based on the feed gas being a relatively dry stream, removing only condensed water and hydrocarbons. The separator is typically a vertical vessel with a side inlet and top outlet for the feed gas to the absorber with a wire-mesh mist pad in the top of the separator. Standard mist elimination pads common in inlet separation vessels have 99% efficiency down to about 10 microns. But, the efficiency


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drops rapidly for droplets below 10 microns. Wire-mesh pads have been reported to have 97 per cent removal efficiency at 8 microns; falling off to 50 per cent efficiency at the 2-micron level. In applications where it is anticipated that the inlet gas may contain particulate such as FeS, a filter-separator may be required. This equipment typically consists of a horizontal vessel with filters in the inlet end of the vessel to remove the FeS followed by mist pads or impingement baffles with a separator chamber to collect any separated liquids. Aerosols, which may be as small as micron, are not removed effectively by standard mist elimination pads. If aerosols are determined to be present, high technology coalescing filtration systems are available which can remove aerosols in the sub-micron range. A water wash system on the inlet feed gas consisting of a small trayed (4-5 trays) or packed column is also effective in removing aerosols formed by upstream equipment. Consideration of a reverse flow coalescer may also be dictated by the necessity to remove iron sulfide from the inlet feed gas that can be as small as sub-micron in size.

6.2 Flash vessel:

The rich amine flash vessel is designed to remove soluble and entrained hydrocarbons from the amine solution and should be operated at as low a pressure as possible in order to maximize hydrocarbon recovery. The removal of hydrocarbons reduces the amine solution foaming potential. Normal operating pressure of the flash vessel ranges from 5 psig to 75 psig, depending upon the disposition of the flash vessel vent stream. A rich amine pump is usually required to pump the rich amine through the lean/rich cross exchanger to the regenerator if the flash vessel operating pressure is lower than 50 psig. A flash vessel should be considered a process requirement in refinery gas treating applications and should be strongly considered in gas plant applications treating wet natural gas (> 8 % C2+) or where a considerable amount of hydrocarbon may be present due to condensation or pipeline slugging. If significant quantities of hydrocarbon gases are flashed from the amine solution in the flash vessel, an absorber with 4-6 trays or an equivalent amount of packing is installed on the top of the flash vessel. A slipstream of lean amine is fed to this absorber to remove H2S and CO2 from the hydrocarbon flash gas prior to going into the fuel gas system. The flash vessel should have adequate instrumentation and level gauges to enable operational personnel to check periodically for the presence of a hydrocarbon layer on top of the amine solution. The flash vessel design should incorporate an internal baffle system as shown in Figure 2 above that allows the hydrocarbon collected in the vessel to be routinely skimmed off. A minimum design residence time for a three phase flash vessel of 20 minutes based on the flash vessel operating half full is recommended. Amine systems treating very dry natural gas (


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Figure 8. Schematic representation of a flash tank

6.3 Absorber:

The absorber diameter is determined primarily by the flow rate of the inlet feed gas. The circulation rate of the amine solution is best determined by rigorous equilibrium loading calculations based on the acid gas content of the inlet sour gas, the strength of the amine solution, the volume of inlet sour gas and the type of amine. For a given absorber application and amine type, a set of curves can be developed if one of the three variables is relatively constant. For example, if inlet feed gas flow rate is relatively constant; a series of curves can be developed utilizing the acid gas content and the amine solution strength as independent variables. Rigorous calculations and simulations should be performed to confirm the quick estimates, especially for applications utilizing MDEA and the formulated MDEA solvents. The amine solution temperature entering the absorber should be 10 to 15 F higher than the inlet feed gas temperature to prevent condensation of hydrocarbon in the contactor, which can cause foaming. The inlet feed gas typically enters the absorber at 100 - 120 F. Therefore, the typical range of lean amine solvent temperature is 115 - 135 F. As a practical maximum, though dependent upon the particular amine and absorber application, the lean amine solvent temperature should generally not exceed 135 F. High lean solvent temperatures can lead to poor solvent performance due to H2S equilibrium problems on the top tray of the absorber or increased solution losses due to excessive vaporization losses.

A differential pressure instrument should be installed on the absorber and stripper tower to monitor the differential pressure across the trays or packing. The differential pressure should be measured from just below the first tray or section of packing to just above the last tray or section. A sharp increase in the absorber/stripper differential pressure is an excellent indication that a foaming problem exists in the system. The typical absorber design does not usually include a provision for several water wash trays (2-4 trays) above the last amine-contacting tray to reduce amine entrainment/carryover into the sweet gas residue. However,


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with the increasing use of specialty solvents in gas treating, amine loss control is becoming an important issue; therefore, an absorber water wash system on the absorber overhead may be justifiable in newer amine system designs. Following similar logic, many existing amine systems are being retrofitted with an absorber overhead carryover scrubber to recover amine carryover from the absorber.

6.4 Tray & packed type absorber:

In general both packed and tray type absorbers are used however when selective removal of H2S is preferred to CO2, then packed tower becomes the obvious choice. H2S reacts much faster with the solvent than CO2 and this aspect of the reaction kinetics is employed in packed tower which owing to low hold up provides shorter contact time between the phases to achieve preferential absorption.

Table 7, gives a comparison between performances of both type of towers for similar operating conditions.

Table 7: Tray vs Packing in selective removal application

Although bubble-cap trays and raschig ring packings were once commonly used in amine plant absorbers and strippers, modem plants are generally designed to use more effective trays (e.g.. sieve or valve types) and improved packing shapes (e.g., Pall rings or high-performance proprietary designs). Very high-performance structured packing is seldom used for large commercial gas treating plants because of its high cost and sensitivity to plugging by small particles suspended in the solution. The choice between trays and packing is somewhat arbitrary because either can usually be designed to do an adequate job, and the overall economics are seldom decisively in favor of one or the other. At this time, sieve tray columns are probably the most popular for both absorbers and strippers in conventional, huge commercial amine plants; while packed columns are often used for revamps to increase capacity or efficiency and for special applications. Tray columns are particularly applicable for high pressure columns, where pressure drop is not an important consideration and gas purity specifications can readily be attained with about 20 trays. Packing is often specified for CO2 removal columns, where a high degree of CO2 removal is desired and the low efficiency of trays may result in objectionably tall columns. Packing is also preferred for columns where


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pressure drop and possible foam formation are important considerations. Packing should not be used in absorbers treating unsaturated gases that can readily polymerize (propadiene, butadiene, butylene, etc.) as gum formation can lead to plugging of the packing. Also, packing should not be used in treating gases containing H2S which are contaminated with oxygen because of the potential for plugging with elemental sulfur. Table 1-5 represents a simplified design guide for both tray and packed type amine absorption column.

Table 8: Trays vs. packing in selective treating with 50% MDEA

After establishing the liquid and gas flow rates, the column operating conditions and the physical properties of the two streams, the required diameters of both the absorber and stripping column can be calculated by conventional techniques. Various correlations have been proposed and available in literature. Pressure drop and flooding data for proprietary packing designs are available from the manufacturers. It is usually necessary to use a conservative safety factor in conjunction with published packing correlations because of the possibility of foaming and solids deposition in gas treating applications.


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Figure 9: Estimation of diameter for tray type amine absorption column

Column heights for amine plant absorbers and strippers are usually established on the basis of experience with similar plants. Almost all installations that utilize primary or secondary amines for essentially complete acid gas removal are designed with about 20 trays (or a packed height equivalent to 20 trays) in the absorber. In bulk acid gas removal applications, experience has shown that if a 20-tray column is supplied with sufficient amine so that the rich solvent leaving the absorber has an acid gas loading that is 75 to 80% of the equilibrium value, then the amine on the upper 5 to 10 absorber trays is very close to equilibrium with the H2S in the treated gas leaving these trays. Therefore, in these circumstances, the H2S content of the treated gas is independent of the absorber design and depends only on the lean amine temperature and the amine regenerator performance.

Absorbers with 20 trays can usually meet all common treated gas CO2 specifications; however, more than 20 trays may be required if CO2 in the treated gas is to be close to equilibrium with the lean amine. Therefore, in applications such as synthesis gas treating, where it is advantageous to reduce the CO2 content of the treated gas to very low levels, absorbers containing more than 20 trays or the equivalent height of packing are often specified. In typical 20-tray absorbers, the bulk of the acid gas is absorbed in the bottom half of the column, while the top portion serves to remove the last traces of acid gas and reduce its concentration to the required product gas specification. With sufficient trays and amine, the ultimate purity of the product gas is limited by equilibrium with the lean solution at the product gas temperature.

When water washing is necessary to minimize amine loss (e.g., with low-pressure MEA absorbers), two to four additional trays are commonly installed above the acid gas absorption section. A high efficiency mist eliminator is recommended for the very top of the absorber to minimize carryover of amine solution or water.

Stripping columns commonly contain 12 to 20 trays below the feed point and two to six trays above the feed to capture vaporized amine. The less volatile amines, such as DEA and


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MDEA, require fewer trays above the feed point to achieve adequate recovery of amine vapors. Typical DEA and MDEA stripping columns use two to four trays, while MEA systems use four to six trays above the feed point Equilibrium conditions alone would indicate that the above numbers are overly conservative; however, the trays above the feed point serve to remove droplets of amine solution, which may be entrained by foaming or jetting action, as well as amine vapor.

6.5 Lean/Rich cross heat exchanger:

The temperature of rich amine leaving the absorber will be 130 to 160 F [54.4 to 71.1 C] and the lean amine from the reboiler will be 240 to 260 F [115.6 to 126.7 C]. The rich amine outlet from the lean/rich cross exchanger is typically designed for a temperature of 200-210 F [93.3-98.9 C], although some amine system designs based on MDEA and formulated MDEA solvents have designed around a rich amine feed temperature to the stripper of 220 F [104.4 C]. Based upon the above amine temperatures, the lean amine from the lean/rich cross exchanger will be cooled to about 180 F [82.2 C]. The most common problem encountered in the lean/rich cross exchanger is corrosion due to flashing acid gases at the outlet of the exchanger or in the rich amine feed line to the regenerator. High rich amine loading due to reduced circulation rate or low solvent concentration increases the potential for acid gas flashing. In many applications, especially for MEA and DGA Agent, a stainless steel (304 or 316) lean/rich exchanger tube bundle should be considered. Stainless steel metallurgy is also more likely to be considered in high CO2/H2S feed gas ratio applications. Adequate pressure should be maintained on the rich solution side of the lean/rich exchanger to reduce acid gas flashing and two-phase flow through the exchanger. Two-phase flow through the exchanger can be a major cause of erosion/corrosion in the cross exchanger. In order to reduce flashing and two phase flow, the final letdown valve on the rich amine, i.e. the flash tank level control valve, should be located downstream of the exchanger and as close as practical to the feed nozzle of the regenerator.

6.6 Liquid/liquid contactor:

The liquid/liquid treater is often the source of much of the losses and problems encountered in the amine system especially in refinery amine units. Amine carried out the treater with the LPG hydrocarbon can be a major source of amine losses as well as a major problem to downstream units such as the caustic treater. Additionally, losing the amine-hydrocarbon interface can introduce large amounts of hydrocarbon into the amine system, completely overwhelming downstream equipment, such as the rich amine flash tank and the carbon filtration system, causing significant problems. The amine liquid treater design criteria presented in Figure 3 and discussed further below assume the LPG/amine interface control is maintained in the top of the LPG treater.


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Figure 10: Typical design guideline for liquid hydrocarbon/amine absorption column

The general rule of thumb for determining the diameter of the absorber is that the combined LPG and amine flow should equate to 10-15 gpm/ft2 of the absorber cross sectional area. The LPG-amine treater is typically a packed tower. The LPG is the dispersed phase while the amine is the continuous phase. Ceramic or steel packing is recommended so the amine will preferentially wet the packing and reduce the coalescing of the LPG on the packing which can reduce the absorber efficiency. Aqueous solvents preferentially wet ceramic packing. Either an aqueous or organic solvent, depending upon the initial solvent exposure, preferentially wets metal packing. Plastic packing should be avoided since organic solvents preferentially wet them. Typical packing size is 1 to 2 inches with 2 to 3 sections of packing with a depth of 10 feet /section. It is recommended that the LPG distributor be below the lower packed bed with the LPG flowing through a disperser-support plate. A ladder-type distributor is a common satisfactory arrangement. The distributor velocity of both hydrocarbon and amine are important. The hydrocarbon distributor velocity is critical. The velocity must be sufficient to allow adequate mixing on the trays or packing but not so severe that an emulsion is formed and phase separation is difficult. The critical amine and hydrocarbon velocities are fairly low. The recommended design LPG distributor velocity is 70 ft/min. The hydrocarbon droplet size is also very important. If the dispersed hydrocarbon droplet is too large, poor treating is the result. Excessive LPG distributor velocities which result in smaller droplet size makes phase separation difficult due to emulsion formation especially if residence time is marginal. The LPG distributor orifice diameter is typically inch. Larger orifices produce non-uniform droplets. Distributor orifices that are too small can produce emulsions thus increasing the absorber amine carryover potential. When the hydrocarbon superficial velocity exceeds the design criteria of 130 ft/hr, the number of orifices is usually increased rather than increasing the orifice size. The entrance velocity of the amine is less critical but should be limited to 170 ft/min to reduce interference with the dispersed LPG rising through the absorber. The amine superficial velocity should be limited to 60 ft/hr. The amine-hydrocarbon interface is usually maintained by a level controller


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operating with the level above the packed section of the absorber. Thus the absorber operates full of amine, commonly referred to as amine continuous. Carryover of amine in the LPG is a common problem. In order to minimize the amine losses, additional headspace should be provided above the normal amine-LPG level for disengagement of the amine and LPG. A coalescer or settling tank is often installed downstream of the liquid treater to aid in the removal of entrained amine from the hydrocarbon. The combined residence time in the absorber and coalescer should be 20 to 30 minutes. A recirculating wash water system to aid in separation should also be considered. The water wash reduces the entrained amine viscosity and aids disengagement in the settling tank.

6.7 Stripper/Reboiler:

The purpose of the stripper is to regenerate the amine solution by stripping the rich amine of the H2S and CO2 with steam generated by the reboiler. The vast majority of the stripping should occur in the stripper rather than in the reboiler. If substantial stripping occurs in the reboiler, excessive corrosion and premature reboiler tube failure is likely, especially in applications with substantial CO2. The regeneration requirement to reach a typical lean loading is a reflux ratio of 1.0 to 3.0. A reflux ratio of 1.0 should be considered as a practical minimum. In some low pressure or tail gas treating applications, higher reflux ratios may be required to meet the product specifications. In order to ensure adequate stripping while at the same time optimizing energy utilization, control of the heat input to the reboiler should be accomplished by monitoring the stripper overhead temperature. The overhead temperature correlates directly with the reboiler energy input. The reboiler temperature is not affected by the amount of stripping steam generated in the reboiler since the boiling point of the amine solution is dependent upon the amine concentration and reboiler pressure. Therefore, the reboiler temperature is not a controlled variable. The heat input to the reboiler should be set to achieve a specified stripper overhead temperature, typically 210 to 230 F depending upon the gas treating application and amount of reflux desired. To prevent thermal degradation of the amine solvent, steam or hot oil temperatures providing heat to the reboiler should not exceed 350 F. Superheated steam should be avoided. 50 psig saturated steam is recommended. The maximum bulk solution temperature in the reboiler should be limited to 260 F to avoid excessive degradation.

6.8 Filter:

A good filtration design includes both a particulate and a carbon filter. The cleaner the amine solution, the better the amine system operates. The particulate filter is used to remove accumulated particulate contaminants from the amine solution that can enhance foaming and aggravate corrosion. Carbon filtration removes surface active contaminants and hydrocarbons that contribute to foaming. With proper inlet gas separation and pre-treatment, filtering a 10 to 20 percent slipstream of the total lean solution has usually proven adequate. Where practical, total stream filtration should be considered. The filtration system is typically installed on the cool lean amine stream (absorber feed). Recirculation of a slipstream from the discharge side of the charge pump to the filtration system with a return to the suction side of the pump is a common arrangement. If combined in series, the particulate filter should be installed upstream of the carbon filter to protect the carbon filter. A second post-filter or screen should be installed downstream of the carbon filter to keep carbon fines out of the circulating system. If the carbon filter is installed independent of the particulate filter, a pre-filter should be installed on the carbon filter inlet to protect the carbon bed. In systems that are extremely contaminated with particulate due to inadequate feed preparation, excessive


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corrosion, or if the inlet gas CO2/H2S ratio is high, particulate filtration of the rich amine exiting the absorber may be required. The concern is that FeS in the rich amine can dissociate in the regenerator under certain conditions to soluble iron products which lean side filtration will not remove. These soluble iron products can then react with H2S in the contactor to form additional FeS, fouling the absorber trays or packing. If components of the filtration system are installed on the rich amine stream, extreme care should be taken when performing maintenance to control the risk of exposure to H2S.

6.9.1 Particulate filter:

The particulate filter should filter a minimum 10 to 20% slipstream of the circulating solution. Numerous particulate filter mediums have been utilized in amine service: wound bleached cotton disposable filter cartridges with polypropylene or metal cores, disposable metal cartridges, pleated paper filter cartridges, sock-type disposable elements and non-disposable/back-flushable mechanical filters with special metal etched filter elements. Experience has shown that a 10-micron absolute filter is adequate for most amine applications, although some MDEA applications as well as many refinery amine applications, which are plagued by a black, shoe polish-like material consisting of iron sulfide bound with hydrocarbon and polymerized amine, require more stringent filtration. The FeS-hydrocarbon shoe polish-like material is very finely divided, with eighty percent of the FeS particles being between 1 and 5 microns in size. 5-micron absolute filtration is typically recommended for these applications.

6.9.2 Carbon filter:

Carbon filter is used in those in amine systems that experience severe emulsion problems due to significant hydrocarbon contamination. A properly designed activated carbon (Activated carbon with high iodine number i.e. high adsorption capacity, high abrasion number i.e. abrasion resistance against degradation is preferred) system can reduce the need for antifoam, reduce amine make up, reduce corrosion and improve scrubbing efficiencies and product quality. The carbon system should treat at least 10 to 20% of the circulating lean amine solution. A minimum contact time of 15 minutes and a superficial velocity of 2 to 4 gpm/sq ft is considered appropriate. When the amine solution changes color or clarity or the solution foaming tendency increases, the carbon is spent and should be changed. Typical maximum carbon life is 6 to 9 months.

7.0 Operational Issues of Amine Sweetening System

A number of operational issues faced in amine gas treating units have been reported. Often one operational difficulty can cause or influence another problem. Not all amine systems experience the same degree of operating difficulties. A continual problem that afflicts one amine system may occur only rarely in another amine system. Several of the more common operational difficulties encountered are discussed below along with troubleshooting recommendations and design considerations whose aim is to improve the amine unit operations and control these common operational problems.

7.1 Failure to meet product specification

Difficulty in satisfying the product specification, typically the H2S specification whether the treated stream is a liquid or a gas may be the result of poor contact (loss of efficiency)


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between the gas and the amine solvent caused by foaming or mechanical problems in the absorber or stripper. In the case of foaming, the gas remains trapped in bubbles, unable to contact the solvent, resulting in poor mass transfer of acid gas to the amine solution. In terms of mechanical damage, if trays are damaged, there may not be enough contact trays for adequate sweetening. If the trays are plugged, there may be poor contact between the gas and the amine solvent on each tray. Other explanations for off-specification product may be related to the amine solution. The amine circulation rate may be too low, the amine concentration may be low, the lean amine solution temperature may be high or the residual acid gas loading in the lean solution may be too high due to improper stripping or a leaking lean/rich cross exchanger. The regeneration requirement to reach the typical lean loading for most applications is a reflux ratio of 1.0 to 3.0. A reflux ratio of 1.0 should be considered as a practical minimum. In some applications, such as low pressure applications, higher reflux ratios may be required to meet the product specifications. A typical reflux flow may be as high as 10-14% of the rich amine solution flow.

7.2 Corrosion

Most corrosion problems in amine plants can usually be traced back to deficiencies in either the design or operation of the amine unit. However, experience has shown that even a well designed and operated amine unit will likely experience some degree of corrosion related problems during its operational life. Therefore, an understanding of the causes of amine unit corrosion is essential in troubleshooting corrosion-related problems. Some areas in an amine system are more likely to experience corrosion than other areas. The regenerator, reboiler and lean/rich cross exchanger will generally have the greater corrosion problems. There are numerous contributing factors affecting amine unit corrosion.


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These contributing factors have been mentioned below:

7.2.1 Amine Concentration:

Generally, the higher the amine concentration, more corrosive is the solution. MEA strength is typically limited to 18-20 weight percent while DEA strength is limited to 30 weight percent. DGA and MDEA solution strengths are usually limited to 50 weight percent in refinery service due to other process considerations associated with the liquid/liquid treaters. DGA has been utilized at concentrations up to 65 weight percent in gas processing service.

7.2.2 Acid Gas Loading :

Operating limits are typically placed on the rich amine acid gas loading in order to limit acid gas breakout, which plays a significant role in amine plant corrosion. The rich amine loading for DEA/MDEA refinery applications should be limited to 0.45-0.475 m/m. MEA and DGA application rich amine loading are typically limited to 0.425-0.45 m/m. Applications with rich loadings beyond these recommended ranges generally require some form of corrosion inhibition or changes in the materials of construction away from carbon steel to stainless. A key consideration when determining the maximum rich solution loading is the feed gas CO2/H2S ratio.

7.2.3 Heat Stable Salts:

HSS, which are the reaction products of the amine and acids stronger than H2S and CO2 which do not dissociate in the regenerator and are therefore heat stable, are corrosive and increase the corrosivity of the solution. Historically, a rule of thumb has been utilized limiting the HSS to 5-10% of the amine alkalinity (for a 50-wt. % amine solution, the 5-10% HSS limit corresponds to 2 to 5 wt. % HSS as amine). However, with the increasing utilization of specialty solvents, a more conservative approach is warranted. Therefore, the HSS level should be limited to 1-2 wt. % when expressed as wt. % amine (3 wt. % maximum). The individual concentration of HSS anions, especially the organic acid anions (acetate, formate and oxalate) should be monitored by routine HSS anion analysis.


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7.2.4 Elevated Temperatures:

High process temperatures tend to promote acid gas breakout as well as having an effect on the amine solution pH, as the solution pH tends to drop with increasing temperature. The rich amine feed temperature to the stripper is typically limited to 210-220 F [98.9-104.4 C] to prevent acid gas breakout. Additionally, the amine solution can be degraded by excessive heat. Thermal degradation potential can be lessened by limiting the bulk temperature of the reboiler amine to 260 F [126.7 C] and limiting the reboiler heating medium temperature to 350 F [176.7 C]. Superheated steam should be avoided. 50 psig [345 kPA] saturated steam is the preferred heating medium. Hot oil and direct fired reboilers should be avoided if possible to avoid potential thermal degradation. If a hot oil or direct-fired reboiler is necessary, care should be taken in the design of the reboiler.

7.2.5 High Velocities:

The velocity of the amine treating solution is limited to control corrosion/erosion caused by the presence of solid particulates as well as acid gas flashing due to excessive pressure drop. Amine solution velocities in the exchangers should be limited to 3 ft/sec [0.9 m/sec] while the velocity in the piping should be limited to 7 ft/sec [2.1 m/sec]. Long-radius elbows should be utilized where practical in rich amine service.

7.3 Solution foaming:

Amine solution foaming is probably the most persistent and troubling operational problem encountered in natural gas production and refinery sweetening operations. Solution foaming contributes significantly to excessive solution losses through entrainment and amine carryover, reduction in treating capacity through unstable operations and off-specification product.

Foaming has a direct effect on capacity due to the loss of proper vapor-liquid contact, solution holdup and poor solution distribution. Foaming can occur in the absorber or stripper and is typically accompanied by a sudden noticeable increase in the differential pressure across the tower. Other indications that a foaming condition exists may be high solution carryover rate, an erratic change in liquid levels, a sharp increase in flash gas flow or a sudden change in acid gas removal efficiency. Solution foaming is caused by changes in the surface chemistry of the amine solution. The factors that cause or enhance the foaming


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characteristics of the solution generally lower the surface tension or raise the viscosity of the amine solution. Foaming of amine solutions can usually be attributed to contamination by one of the following:

Suspended solids and particulate matter. Liquid hydrocarbons. Organic acids in the inlet gas, which react with the amine to form soap-like material. Surface-active agents contained in inhibitors, well treating fluids, compressor oils, pump lubricants and valve lubricants. Amine degradation and decomposition products. Heat stable salts (HSS).

These contaminants in conjunction with process conditions such as temperature and pressure interact to alter the surface layer characteristics that control the formation and stability of the foam such as elasticity of the film layer and film drainage. A clean amine will not form stable foam. Any contaminant that lowers the solution surface tension and raises viscosity can enhance foaming tendency and foam stability. H2S reacts not only with the amine but also with the metallurgy of the gas treating plant, which is typically carbon steel, to form iron sulfide. Additionally, finely divided iron sulfide can also enter the amine system with the inlet sour gas. Over a period of time, the iron sulfide will deposit throughout the plant, forming a thin protective layer that prevents further corrosion as long as it remains undisturbed. However, if the velocity of the amine solution is excessive the thin protective layer of iron sulfide is continually removed which exposes the metal for further corrosive attack. Iron sulfide is a very fine particulate and tends to accumulate on the surface of the treating solution increasing the solution surface viscosity and retarding the migration of liquid along the bubble walls when foam forms. The finely divided iron sulfide particulate tends to stabilize foam by retarding film drainage of the film layer encapsulating the gas bubbles that make up the foam. Iron sulfide is the most common particulate found in amine solutions. However, in systems containing no H2S, iron carbonates and oxides can be formed. Additionally, particulate can enter the amine system with the feed gas or makeup water. Solids that may enter via the inlet feed gas include rust particles, dirt, pipe scale, salts and iron sulfide as mentioned earlier. Iron sulfide entering with the inlet gas is a particular problem in many natural gas plants that normally can be corrected by installing a filter separator on the inlet feed to the amine contactor. Solution foaming is the most common operational problem caused by high particulate levels but high solid levels can also plug contactor trays or packing and foul heat exchangers. Removal of particulate matter can best be accomplished by continuous filtration of a side stream of the circulating amine solution. With proper inlet gas separation and preparation, filtering a 10 to 20 percent slipstream of the lean amine solution has proven successful in reducing particulate contamination that contributes to foaming problems. Additionally, a carbon filter should be installed downstream of the particulate filters. Carbon filtration has been shown to remove surface-active contaminants such as hydrocarbons that also contribute to foaming.

7.4 Excessive solution losses:

The most common ranking of solvent loss categories from highest to lowest is 1) mechanical, 2) entrainment due to foaming and solubility, 3) vaporization and 4) degradation. The majority of solvent loss is due to mechanical and entrainment due to foaming/emulsions and solubility. Vaporization and degradation losses constitute a small portion of the overall solvent losses. For a 30 wt% DEA solution, operating at 500 psia system pressure and 140 F,


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losses due to vaporization and degradation are estimated to be about 0.10 lb. DEA/MMSCF. Actual makeup requirement losses may range from 1-3 lbs/MMSCF, dependent on the application. Therefore, vaporization and degradation account for as little as 3% of the overall solution losses. Amine solution losses for gas plant applications are typically much lower than refinery applications. It is not uncommon for refinery amine losses to be several times gas processing amine makeup rates. When reviewing excessive solution loss problems, the two areas to focus on are A) Entrainment and B) Mechanical. Entrainment losses are a direct function of the gas and liquid hydraulics in the absorber or regenerator. Excessive solution foaming can also contribute to losses due to mechanical entrainment as described earlier. Losses due to entrainment of the amine in the absorber outlet gas by way of a mist or spray can be reviewed by confirming the tray design of the absorber to determine the actual load on the absorber trays compared to the original design. Operating trays near or above flooding can cause increased formation of droplets, which may entrain in the gas as a mist or spray. The mechanical integrity as well as the capacity and design of the absorber mist eliminator or downstream knockout equipment should be verified. The mechanical integrity of the absorber trays themselves must also be verified. In amine systems that have a liquid/liquid treater present, entrainment of the amine solution in the hydrocarbon due to emulsions also becomes an issue. Liquid treaters are designed for low velocities for both the amine and hydrocarbon phases in order to prevent small amine droplet formation and reduce emulsion formation. The observation of an emulsion "rag" layer between the hydrocarbon and amine phase in the liquid absorber level glass is an indication of small-droplet formation. Solving liquid treater entrainment losses requires careful evaluation of the treater design specifications. High absorber velocities due to poor design or damage should be corrected if possible. If the entrainment persists, downstream separation equipment such as a wash water system is required to remove the entrained amine.

7.5 Heat Stable Salt (HSS) Management:

The principal problems associated with HSS contamination of the amine system include: (1) Decreased amine system capacity, (2) Excessive corrosion (3) Operational problems caused by foaming and corrosion by-products which result in excessive amine losses, high filter change-out costs and poor amine system performance.

HSS are formed in amine systems when trace acidic components (weak acids) in the sour gas react with the amine solution (a weak base) to form soluble amine salts. These HSS cannot be regenerated at stripper conditions in a fashion similar to the reversal of the H2S/CO2 amine-base complex. The bound amine of the HSS can no longer react with the incoming acid gas; thus the system capacity is reduced. The HSS content of the amine solution is determined by an ion-exchange/titration method which determines the total equivalents of all anions present in the solution and is reported as the amount of amine tied up in the form of amine salts. This method does not distinguish between amine HSS and inorganic HSS (sodium [Na] or potassium [K]). Comparison of the HSS by the titration method result with ion chromatography (IC) results and cation analysis which determines Na and K helps determine to what extent HSS are present as amine salts versus inorganic salts. The HSS precursors found in many refinery applications, such as the carboxylic acids (formic, oxalic and acetic) responsible for HSS contamination, typically come from sources such as the FCC and coker off gas. The principal culprit in HSS formation (thiosulfate) in tail gas amine systems is due to SO2 breakthrough past the hydrogenation reactor and the tail gas quench water system. A few of the more commonly found HSS anions are acetate (CH3COO-), formate (HCOO-),


39

thiocyanate (SCN-), sulfate (SO4-2) and thiosulfate (S2O3-2). The amine supplier should routinely test for the following HSS anions: 1) acetate 2) glycolate, 3) formate, 4) chloride, 5) sulfite, 6) sulfate, 7) oxalate, 8) thiosulfate and 9) thiocyanate. Additionally, the amine supplier should routinely perform cation analysis to determine Na and K levels. Faced with the problems associated with excessive HSS contamination, the amine system operator is faced with a number of possible corrective actions to control HSS contamination. Three primary courses of Heat Stable Salt Management action can be taken to control problems associated with HSS:

1) HSS Preventative Measures 2) HSS Neutralization Measures 3) HSS Removal Measures

7.5.1 HSS prevention measures:

The principal method of HSS prevention is to reduce the incursion rate of the various acidic precursors by employing a water wash on the feed gas to the absorbers. The quench water system serves this purpose in the tail gas system though the proper control of the tail gas hydrogenation section is also very important. Some refiners have reported a reduction in HSS incursion rate of up to 50% by selectively water washing the offending sour gas streams. This corrective action is typically employed only in instances where the HSS incursion rate is high and the cost of the water wash installation is more easily justified. Additionally, economics must recognize the expense of installing water wash systems as well as the increased load on the sour water stripper system; alternatively, if the source of the offending acidic precursors is identifiable, reduction in the HSS incursion rates may be obtained by altering the process operating conditions of the process unit generating the acidic precursors.

7.5.2 HSS neutralization measures:

The addition of strong bases to the circulating amine solution such as caustic (NaOH) or soda ash (Na2CO3) as well as KOH and K2CO3 .neutralizes. the amine HSS, displacing the amine, freeing the "bound" amine and restoring amine capacity.

These neutralizing agents react as follows:

NaOH (KOH) + Amine H+ HSS- Amine + H2O + Na(K)HSS Na(K)2CO3 + 2H2O + 2 Amine H+ HSS- 2 Amine + 2H2O +2CO2 + 2 Na(K)HSS

Therefore, while the addition of alkali does restore system capacity by freeing bound amine, it does not reduce the anion content, rather it simply converts the amine HSS to a sodium or potassium HSS. The use of potassium alkali, KOH and K2CO3, is preferred since the potassium HSS is typically more soluble in the aqueous amine solution than the sodium HSS. Reduction in the corrosivity of the amine solution due to neutralization has been reported in literature since the inorganic HSS is less likely to partially disassociate at reboiler conditions (generating free acid), which is a suspected HSS corrosion mechanism, especially for formate.


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7.5.3 HSS removal measures

Purging or "Bleed and Feed" - This method involves dumping a portion of the HSS contaminated amine solution from the amine system, replacing it with fresh solution and appropriately disposing of the contaminated solution. The high cost for proper disposal of the contaminated amine and the inherent cost of the amine solvent generally deter the operator from employing this method. Additionally, the practice of disposing of the contaminated amine to a wastewater treatment system is usually prohibited due to environmental considerations. Therefore, this method is typically not employed except only in the most urgent circumstances. Followings reclamation methods for removal of HSS are well practiced in industry,

Electro-dialysis Reclamation Electro-dialysis units are stacks of membranes that allow selective passage of anions and cations through the membrane media under an electrical field. These units separate the HSS contaminated amine solution into two effluent streams, (1) a reclaimed amine stream and (2) a brine waste stream containing the HSS anions and sodium (Na) or potassium (K) cations, if alkali is utilized to "neutralize" the HSS. The electro-dialysis units appear to work best in systems with

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