Gas Sweetening Write-up - Final

Gas sweetening by amine

A Technical Report On Gas Sweetening by Amines

Subhasish Mitra, Mitra, Sr. Process Engineer Petrofac Engineering India Ltd

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Content 1. Introduction 2. Gas sweetening basics 3. Alkanolamine gas treatment basics 4. Alkanolamine gas treating chemistry 5. Alkanolamine processes-Strengths & Weakness/Solvent selection 6. Amine System Description 7. Operational Issues of Amine Sweetening System 8. Troubleshooting guide 9. Prevention of BTEX emission 10. Bulk CO2 removal technology by membrane unit 11. New developments Appendix - 1: Typical process specification for gas sweetening package Appendix - 2: Typical process flow sheet for amine absorption unit prepared in 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 microporous solid. The mixture of adsorbed components is called the adsorbate, and the microporous 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 noncatalytic 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 Page 6 of 57


partial pressure is low because the capacity of physical solvents is a strong function of partial 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 highpressure 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 MDEA’s 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) C2H4OH - NH2

DIGLYCOLAMINE Agent (DGA) 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+ + OHIonization of dissolved H2S: H2S = H+ + HSHydrolysis 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, Henry’s law can be applied to relate the vapor and physically dissolved liquid concentrations: φiyiP = γimiHi (i = H2S, CO2)

- Equation 9

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where φi = fugacity coefficient of component i yi = mole fraction of component i in vapor phase P = total pressure of the system γi = activity coefficient of component i mi = concentration of component i in liquid phase Hi = Henry’s 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 (γ) 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 ¼ 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)oxazolidone2(HEOD), N,N.bis(2-hydroxyethyl) piperazine (BHEP) and N,N,N’-tris(2-hydroxyethyl) ethylenediamine (THEED). An explanation for DEA’s wide utilization within the gas treating industry is due to DEA’s 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, ¼ 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 (DiEthanol Amine)

Solvent Cost

Low Solvent Cost

Solvent Loss

High solvent vapor pressure results in higher solvent loss.

Low Solvent Cost Low solvent vapor pressure results potentially lower solvent loss.

Selectivity Thermal Stability

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

Reactivity

High reactivity due to its primary amine characteristics.

Corrosion

Higher Corrosion potential

Recovery (Reclaimation )

Requires continuous reclaiming.

Non-selective removal in a mixed acid gas system.

Low reactivity compared to MEA and DGA Agent. Reduces corrisive nature compared to MEA. Non-reclaimable by conventional reclaiming techniques.

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DGA (Di-Glycol Amine Agent)

MDEA (Methy Di Ethanol Amine)

Relatively high solvent cost

Highest Solvent Cost

Non-selective removal in a mixed acid gas system. Removal of COS and CS2. Excellent Thermal Stability High reactivity, 1/4 grain H2S specification can generally be obtained for applications with low operating pressures & high operating temperatures.

Low vapor pressure which results in potentially low solvent loss. 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.

Lower comparative reactivity

Less corrosive Non-reclaimable by conventional reclaiming techniques.


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 2½micron 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 ½ 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 (<2 % C2 +) or Syn-Gas streams with very little hydrocarbon content can utilize a lower flash vessel residence time of 5 minutes if a flash vessel is incorporated into the amine unit design.

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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 aminecontacting 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 highperformance 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 1½ 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 ¼ 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 prefilter 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 nondisposable/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 510% HSS limit corresponds to 2 ½ 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. Page 34 of 57


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 [98.9104.4 °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]. Longradius 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 aminebase 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-),

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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 a high level of HSS contamination. The brine effluent stream typically contains some amine. This effluent stream may contain enough amine to cause problems in the refinery wastewater treatment facility.

•

Vacuum Distillation Reclamation - A proprietary vacuum distillation process utilizes a caustic neutralization pretreatment step to neutralize the bound acids followed by distillation under a vacuum to remove the converted salts. The principal advantage of the vacuum distillation process is derived from the unit’s ability to concentrate the amine solution in the effluent, thus effectively; de-watering the HSS contaminated solution as well as reducing the HSS anion content. The process is energy intensive and thus can be expensive. Disposal of waste streams may also be a potential obstacle.

•

Ion Exchange Reclamation - Ion exchange uses anion and cation resins to replace the HSS ions and sodium/potassium with water. The process produces no solid waste. Ion exchange does require a substantial amount of regeneration chemicals, but the waste streams are generally compatible with existing wastewater treatment facilities. Various amine/service vendors offer both on-line and off-site amine reclaiming services. Several factors must be considered when evaluating the need for reclaiming the amine solution for HSS removal however proper amine reclaiming option for HSS removal frequently balances a) economic consideration & b) disposal of waste/byproduct stream. 7.5.4

HSS incursion rate:

Low HSS incursion rate amine systems are not likely candidates for frequent online reclaiming but amine systems with a high HSS incursion rate require a closer review of the system reclaiming requirements and frequency of treatment. Philosophy regarding HSS management should focus heavily on preventative measures to reduce the HSS incursion into the amine system. Where possible, wash water systems to reduce the HSS incursion rate should be considered. Proper operation of the upstream process units feeding sour gas to the amine system and the control of oxygen into the system can also reduce the HSS incursion

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rate, especially in tail gas units. Failing prevention, neutralization should be considered but it must be recognized that alkali neutralization may only be a short-term fix. Low to moderate HSS incursion rate amine systems can be effectively managed using neutralization. However, if the HSS incursion rate is high, eventually, HSS removal is the last and final option available after all others fail. Recognizing the economic realities of the addition of water wash systems (they rarely can be justified though few doubt the benefits) and the fact that the refinery amine system operates in the real world, HSS removal is a necessary requirement for many refinery operators. In order to enhance the amine system operational efficiency, the HSS should be reduced/controlled to a level of between 1-2 wt % when expressed as amine (3 wt. % maximum).

8.0 Troubleshooting guide Successfully troubleshooting the common operational problems encountered in amine gas treating requires the unit operator to detect and identify symptoms, interpret the symptoms, determine likely causes and finally correct the root cause of the problem. The following is a troubleshooting checklist of the most common operational problems and likely root causes. 8.1 Off-spec product: If the H2S content of the treated hydrocarbon stream is high, the amine unit operator should check the following: 8.1.1

Proper Solvent Circulation Rate:

Rich amine solution loading is to be checked by lab analysis or material balance. If the loading is greater than the typical operating limits or the amine unit process design basis, solvent circulation rate is to be increased. Alternatively, if the circulation rate is already at a maximum, the solution amine strength can be increased to the operating limit. 8.1.2

Lean Amine Solution Loading:

The recommended maximum rich loading for amines is dependent upon the acid gas breakdown in the feed to the absorber. For systems treating gas with much higher H2S and CO2, it is acceptable to run the rich loading of 0.55 moles/mole. This is possible because of the protective iron sulphide layer formed by the corrosion reaction of iron and H2S. This protective film, while not impenetrable, does lay down a protective film which retards further corrosion at the location with the film. Solids flowing along the pipe walls can scour off the film, so it is imperative to keep the solution clean to avoid continuously exposing fresh metal to corrosion. High CO2 : H2S ratio plants cannot maintain rich loading much above 0.45 moles/mole because the predominate corrosion species (iron carbonate) does not form a protective film. CO2 attack is much more insidious than H2S attack. H2S attack tends to corrode a larger area in a kind of wash pattern, while CO2 attack tends to be pitting type corrosion. Pitting is much more severe than general corrosion because of the potential for breaching of a vessel wall or piping. While these are general guidelines, it is also true that amines are capable of absorbing acid gases to a given equilibrium value, depending upon the acid gas partial pressure and the

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system operating temperature. For systems with very low acid gas partial pressures, the maximum equilibrium loading may only be around 0.5 moles/mole. High acid gas partial pressure will result in solution loadings in excess of 1 mole/mole. For whatever operating conditions are present in your facility, it is recommended not exceed 80% of the equilibrium absorption capacity of the solvent. Too much acid gas in the rich solvent will result in easy liberation of acid gases along the rich loop, resulting in high corrosion potentials, especially after the lean/rich cross exchanger. Acid gas will also be liberated in the exchanger itself and where there are flow direction changes or pipe diameter changes. The guiding limit for circulation rate reduction will be one of two things, both equally important: • •

Rich solution loading must remain below recommended guidelines (< 0.55M/M) Absorber maximum bulge temperature must remain below 85 °C.

The first one can be determined by titration or material balance around the amine unit while the second one can be determined via simulation or thermal gun (assuming insulation can be bypassed). Whenever either of these important parameters is reached, that is the limit for circulation rate reduction. Lean solution loading is to be checked and compared to available historical norms, typical guidelines and the process design basis. If the loading is abnormally high, followings need to be checked, a) Lean/Rich Cross-Exchanger Leak: Lean solution loading across the lean/rich cross exchanger is to be checked. If the cross exchanger outlet solution loading is significantly higher than the inlet lean amine loading, cross exchanger leak is a suspect. A heat balance around the cross exchanger typically will confirm the magnitude of the cross exchanger leak and the resulting cross contamination b) Improperly Stripped Solvent: Regenerator operation needs to be checked for this scenario. If the reflux flow rate is below historical norms or the regenerator overhead temperature is low (< 210 °F [98.9 °C]), increase the heat to the reboiler (increase steam rate, hot oil flow or fuel gas flow). Recheck the lean loading and compare to historical norms. If the solution loading is still abnormal with the reflux flow/regenerator overhead temperature on the high side of the historical data, mechanical damage or foaming in the regenerator is suspected. 8.1.3

Regenerator Foaming:

Foaming in the regenerator affects the solvent stripping efficiency and can therefore impact the lean solution loading. Regenerator system requires checking for evidence of foaming such as increased tower differential pressure or erratic levels. Spot use of anti-foam may be required to regain product spec until the root source of the foaming can be determined.

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8.1.4

Lean Amine Solution Temperature:

Absorber feed solution temperature should be checked and compared to historical norms and the process design basis. In most applications, the lean amine solution temperature is typically 110 to 120 °F [43.3 to 48.9 °C]. For most gas treating applications, as a practical maximum, the lean amine temperature should not exceed 135 to 140 °F [57.5 to 60 °C]. If the solution temperature is excessive, check the lean amine cooler for proper operation and the lean solution cooler bypass (if applicable). 8.1.5

Absorber Foaming:

Absorber is to be checked for foaming symptoms, increased absorber pressure differential or solvent carryover and preventive measures are to be taken accordingly. 8.1.6

Absorber Operating Below Turndown:

Inlet process conditions need to be checked and compared to the original process design basis and available historical data to see if there has been a significant change in the operating conditions. 8.1.7

Absorber Mechanical Damage:

If the absorber has multiple feed locations, check that the top feed location is in service. A scan of the absorber to determine mal-distribution due to damaged trays or plugging may be required. For applications utilizing MDEA and formulated MDEA solvents which have a CO2 specification such as high-pressure applications with a CO2 spec of 2-3 %, if the CO2 content of the treated gas is high, following need to be checked, •

Lean Amine Solution Temperature:

The CO2-MDEA absorption mechanism is temperature dependent. CO2 removal is reduced at lower absorber temperature profiles. Therefore, the lean amine solution temperature should be checked to ensure it is not too low, typically less than 90 °F [32.2 °C]. Raise the lean amine solution temperature to increase the absorber temperature profile and increase CO2 removal. Otherwise, the troubleshooting checklist for high H2S in the residue gas is applicable as described above. If the CO2 content in the treated gas for selective applications is too low, check the following: 1. Excessive Solvent Circulation Rate. 2. Lean Amine Solution or Inlet Absorber Feed Gas is too hot. 3. Too Many Absorber Trays. Lower the feed point to the absorber. 8.2 Foaming: Amine solution foaming is probably the most persistent and troubling operational problem encountered in gas sweetening operations. Foaming has a direct impact on capacity due to the loss of proper vapor-liquid contact, solution holdup and poor solution distribution. Foaming

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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, an erratic/drastic drop in liquid levels and a sharp increase in flash gas flow. If a foaming condition is suspected, the amine unit operator should check the following: 8.2.1

Lean Amine Solvent Temperature:

The lean amine solvent temperature entering the absorber should be 10 to 15 °F [5.6 to 8.3 °C] warmer than the inlet feed gas temperature to prevent hydrocarbon condensation. The cooler the lean amine temperature, greater is the capacity for H2S removal and loading the solvent. If the lean amine temperature enters the absorber at a temperature lower than the feed gas temperature, there will be hydrocarbon condensation in the absorber. If the amine-inlet gas temperature approach is less than 10 °F [5.6 °C] increase the amine temperature to reduce the hydrocarbon condensation potential, if practical, i.e. the H2S overhead specification is okay. Prevention of hydrocarbon condensation can be easily achieved by simple diligence when operating lean amine coolers. Maintain a lean amine temperature as close to the feed gas temperature as possible, but never below the feed gas temperature. It is also possible to feed the lean amine at too high a temperature into the absorber. This is especially true during summer operation when most plants find themselves lacking in lean amine cooling duty. The inability to cool the lean amine to a reasonable temperature (< 49°C/120°F) can make the difference between meeting and failure to meet specification (especially if the sales gas specification is < 4 ppm). Maximum temperature of lean amine shall be 49°C. The feed gas temperature should be controlled to enter the amine unit between 27°C and 45°C. Some facilities operate with much cooler feed gas temperatures, which is acceptable, however, the lean amine temperature should still never drop below 27°C. Higher feed gas temperatures result in higher lean amine temperature requirement, which reduces the acid gas carrying capacity and increases the sales gas H2S content. In order to moderate the temperature, cross exchange between the lean amine and feed gas is a viable option. 8.2.2

Flash Tank Operations:

If hydrocarbon contamination is suspected as the cause of the solution foaming, the flash tank should be operated at minimum pressure. Flash Tank should be checked for excessive hydrocarbon and hydrocarbons are to be drained if possible. 8.2.3

Inlet Separation:

Operation of the feed gas inlet separator for possible hydrocarbon carryover due to a demister failure is to be checked. If the lean amine feed to the absorber shows no adverse foaming from a simple shake test but the rich amine foams, the problem is probably due to something bypassing the inlet separation/filtration equipment. It is true in all process operations that the process itself only works as well as the level of contaminants will allow. The most important vessel in the amine unit, besides the absorber and stripper, is the inlet separator. Contaminants carryover into the absorber can result in

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foaming, solution degradation, corrosion, plugging, heat stable salt building-up and equipment damage. At the minimum, an amine unit must have some form of bulk fluid separation device upstream of the absorbers (contactors). These are slug catchers or simple two or three phase separators, designed to catch a slug of liquid, gravity separate the solids and liquids from the gas and allow the gas to pass through a mesh pad before it enters the base of the absorber. The Primary contaminant is hydrocarbon, whether naturally occurring or from compressor oils. Hydrocarbons need to be removed because they can cause the amine solution to foam, which can be as small a consequence as increased antifoam consumption to a severe case of failure to meet specification. The more effective means of contaminant removal is the installation of a filter/coalescer, downstream of a bulk separator, to trap up to 70% of the feed gas aerosols. There are many plants that could not operate without the presence of the inlet coalescer/filter. The more effective coalescer is a vertical unit, rather than a horizontal vessel. Whenever horizontal vessels are installed, there is always the potential for the filter elements to be submerged in the liquid it is trying to coalesce. At this point the filters are no longer effective as coalescing elements. With vertical coalescers, the liquid is drained away from the filter elements and the opportunity for failure is reduced. Inlet separators should be placed within 15 meters (50 feet) of the absorber or hydrocarbon may condense in the line due to cooling. Mesh pads should always be installed on the horizontal plane. Vertically placed mesh pads may be flood with liquid, at which point they act as siphons rather than coalescing elements. 8.2.4

Carbon Filter:

If the amine solution foams as badly downstream of the carbon filter as upstream the carbon filter is probably spent and should be replaced. 8.2.5

Particulate Filter:

FeS particulates in the amine solution tend to stabilize amine solution foaming, therefore, if the circulating amine is heavily contaminated with FeS, check the operation of the amine solution filters and replace the filters as needed. •

Temporary foam control

When foaming occurs, it is usually necessary as a short term measure to utilize an antifoam agent to reduce the foaming tendency of the amine solution until a more permanent remedy can be found. Use of an antifoam agent should be considered a short-term measure rather than a long-term cure. Determining the source of the contamination causing the foaming and preventing future contamination is the best long-term fix. It may be necessary to evaluate a number of antifoams to determine the best selection. Any antifoam agent considered should be tested before use in the amine system. A simple shake test is usually adequate in screening potential antifoam agents. Antifoam agents are typically of two types, silicon based or longchained alcohols. The manufacturer guidelines for antifoam injection should be followed since excessive antifoam can aggravate the foaming problem. Check to see if the carbon filter removes the selected antifoam. Typically, silicon based antifoams are removed by the carbon

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filter. With the carbon filter in service, continuous addition of the silicon antifoam agent may be required. However, it is recommended that the antifoam addition be made in small periodic doses as needed to control the foaming. Great care should be taken to prevent overdosing the amine system with antifoam.

8.3 Excessive solution losses: Solution foaming contributes significantly to excessive solution losses through entrainment and amine carryover; therefore, excessive solution losses and solution foaming are linked. Solution losses are either physical or chemical based. The major source of amine losses is physical in the form of mechanical/physical entrainment. Losses due to chemical means (degradation and vaporization) are fairly small. In solving an excessive solution loss problem, the amine unit operator should concentrate on the following areas: 8.3.1

Reducing Absorber Foaming and Subsequent Solvent Carryover:

With the natural linkage of excessive losses and solution foaming, solving an excessive foaming problem will likely significantly reduce the amine system solution losses. 8.3.2

Optimizing Regenerator Operations:

Some amine systems must purge reflux to control ammonia contamination or maintain proper water balance. If the reflux contains an excessive amount of amine due to mechanical problems in the regenerator, the reflux purge can account for a significant amount of the amine losses. If a reflux purge is utilized to control ammonia, in addition to testing the reflux for ammonia content, the amine content of the reflux purge should be checked routinely to determine the extent of the amine losses directly related to the reflux purge. 8.3.3

Particulate Filter Change-out Procedure:

A review of the filter change-out procedure leads to a tightening of the amine system and a reduction in solution losses. A significant portion of the solution losses may be traced to lack of proper amine recovery from filters, pump seal flushes and the flash drum. •

Remedy-Solution analysis:

Regular sample collection and accurate analysis of the plant amine solution can be used as a tool to resolve operational problems and improve the operation of the gas treating system. Proper testing provides valuable information about the amine solvents physical and chemical condition. Plants should periodically review their amine solution analysis program to be sure they are obtaining the necessary information for the best control of plant operations.

9.0 Prevention of BTEX emission: The aromatic compounds including Benzene, Toluene, Ethyl-benzene, and Xylene (collectively known as BTEX), are included as hazardous factors in air pollutants. H2S and CO2 present in the feed gas are absorbed by the amine in the contactor column and the sweetened gas exits the top of the column. Rich amine exits the bottom of the column and is sent through the regeneration system to remove the acid gases and dissolved hydrocarbons,

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including BTEX. The lean solution is then circulated to the top of the absorber to continue the cycle. The sweetened gas exiting the absorber is saturated by water from its contact with the amine. The overheads, including BTEX from the amine regenerator column, are sent to a sulfur recovery unit. If the raw gas contains appreciable amounts of H2S; a sulfur plant is used to treat the overheads from the rich amine stripper. This treating normally destroys any BTEX or other hydrocarbons. Several operating parameters directly affect the amount of BTEX absorbed in an amine unit, such as inlet BTEX composition, contactor operating pressure, amine circulation rate, solvent type, and lean solvent temperature. MDEA absorbs the lowest amount of BTEX compared to DEA and MEA; therefore, it is recommended to use MDEA where BTEX is observed in the sour gas, (if it is applicable). Several operating parameters directly affect the amount of BTEX absorbed in an amine unit. These factors include the inlet BTEX composition, contactor operating pressure, amine circulation rate, solvent type, and lean solvent temperature. Following list of strategies can be followed to limit the BTEX emissions from gas plant: 1. Minimize the lean amine temperature. The amount of BTEX emissions in amine systems decreases with an increase in lean solvent temperature. 2. Use the best solvent for treating requirements. (i.e. MDEA absorbs the lowest amount of BTEX). 3. Minimize the lean circulation rate. BTEX pick up increases almost linearly with an increase in circulation rate. 4. If the stripper pressure is higher, the overall BTEX emissions are lower.

10.0 Bulk CO2 removal technology by membrane unit: Of late membranes have found their use in bulk removal of CO2 from natural gas. As high as 30% CO2 concentration in natural gas has been reported to be successfully processed in membrane units to reduce CO2 content~ 10 -15%. Many often a hybrid system of membrane and amine system is used to meet sales gas specification. CO2 removal membranes do not operate as filters, where small molecules are separated from larger ones through a medium with certain size pores in it. Instead they operate on the principle of solution-diffusion through a non-porous membrane. CO2 first dissolves into the membrane, and then diffuses through it. Since the membrane does not have pores it does not separate on the basis of molecular size; rather it separates based on how well different compounds dissolve into the membrane and then diffuse through it. Carbon dioxide, hydrogen, helium, hydrogen sulfide and water vapor, for example, are highly permeable gases and so are characterized as "fast" gases, whereas carbon monoxide, nitrogen, methane, ethane and other hydrocarbons are characterized as "slow" gases. The membranes therefore allow selective removal of fast gases from slow gases. Polymer-based membranes are used for this purpose. These currently include cellulose acetate, polyimide and polysulfone. The most widely used and tested material is cellulose acetate. Polyimide has potential in certain CO2 removal applications, but lacks the breadth of commercial experience of cellulose acetate. The properties of polymers can be modified to enhance membrane performance. For example, polyimide membranes were initially used for hydrogen recovery, but were then modified for CO2 removal. Cellulose acetate membranes were initially developed for reverse osmosis, but now are the most rugged CO2 removal membrane available. Membrane performance has been enhanced from selectivity aspect by

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modifying existing membrane materials and investigating alternate membrane materials. Of equal importance have been improvements in membrane element configuration, pretreatment design and optimization of the system’s mechanical design. Gas separation membranes are currently manufactured in one of two forms: flat sheet or hollow fiber. The flat sheets are typically combined into a spiral wound element, while the hollow fibers are combined into a bundle, similar to a shell and tube heat exchanger. In the spiral wound arrangement, two flat sheets of membrane with a permeate spacer in between are glued along three of their sides to form an envelope which is open at one end. Many of these envelopes are separated by feed spacers and wrapped around a permeate tube, with their open ends facing the permeate tube.

Fig: 11 Spiral wound membrane unit Feed gas enters along the side of the membrane, and passes through the feed spacers separating the envelopes. These feed spacers also provide mechanical strength. As the gas travels between the envelopes, CO2, H2S and other highly permeable compounds permeate into the envelope. These permeated components have only one outlet, which is to travel within the envelope to the permeate tube. The driving force for transport is the differential pressure between the low-pressure permeate and high-pressure feed gas. Once permeate gas reaches the permeate tube it enters it through holes drilled in the tube. From there it travels down the tube joining permeate from other tubes. Any gas on the feed side that does not get a chance to permeate, leaves through the side of the element opposite the feed position.

Fig: 12 Follow fiber membrane unit

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In high CO2-removal applications from natural gas a significant amount of hydrocarbons permeate the membrane and are lost. Multistage systems attempt to recover a portion of these hydrocarbons. The one stage with recycle design allows only a portion of the first stage permeate to be lost. The rest is recycled to the feed of the first stage. The portion of first stage permeate that is lost is usually taken from the first membrane modules in the system, where feed CO2, hence permeate CO2, is highest and hydrocarbons are lowest. Permeate that is recycled is at low pressure and must be re-pressurized before it can be combined with the feed gas.

Fig:13 One stage flow scheme Advantage of membrane system: • • • • • • • • • • • • •

Reduced initial capital expenditure Lower operating cost Faster & simplified installation Operational simplicity Adaptability High turn down Design efficiency Power generation from permeate gas High reliability and on-stream time Ideal for de-bottlenecking Low labor requirement Environmentally friendly Ideal for remote location

Disadvantage of membrane system: • • •

Higher hydrocarbon losses Coarse removal of Hydrogen sulfide Effective only for coarse removal of CO2

11.0 New developments: The objective of gas treating has always been removal of unwanted impurities in a safe and cost effective manner. A gradual progress has been under notice to achieve this through minimizing cost while still meeting desired product specification, minimizing corrosion and reducing operating problems. Of late there has been a shift in solvent selection from generic

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amines to proprietary amines developed with enhanced selectivity as an essential part of this progress. For highly selective H2S removal, solvents by The DOW Chemical Co. (Gas Spec), Union Carbide (Ucarsol), BASF (aMDEA), EXXON (Flexsorb), and others have been developed that exhibit greater selectivity and H2S removal to lower treated gas specifications. However, these solvents are MDEA-based solvents. These solvents have other applications; such as H2S removal from CO2 enhanced oil recovery (ROR) enrichment processes. Solvents for H2S selectivity are used for refinery systems with high CO2 slip, tail gas treating, natural gas treating, H2S removal from liquid hydrocarbon streams, natural gas scrubbing, and refinery systems with LPG streams containing olefins. Inventions of all these new amines with selective kinetics have contributed towards successful gas sweetening with increased emphasis on environment emission.

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APPENDIX-1 : TYPICAL PROCESS SPECIFICATION FOR GAS SWEETENING PACKAGE 1.

INTRODUCTION This document defines the Duty Specification for the Acid Gas Removal Unit to be installed as part of the onshore processing facilities for the XX project. The AGR unit will remove hydrogen sulphide from sour gas and will partially remove other components including CO2, mercaptans and COS. The AGR unit will also supply a small lean amine stream to treat NGL for removal of H2S and CO2 in the liquid treating area and receive the rich amine.

2.

PROCESS DESIGN SPECIFICATION

2.1

GENERAL AGR processes sour feed gas from the inlet reception facilities of the plant combined with condensate stabilisation recompressed offgas. The feed gas contains H2S, CO2 and organic sulphur compounds (mainly mercaptans). H2S and mercaptans shall be removed from the gas prior to the export to pipeline. The AGR shall remove the H2S down to the required specification. Mercaptans removal in the AGR is desirable, as it will simplify the downstream gas processing. The level of mercaptans removal will be defined by the Licensor. CO2 will also be defined by the Licensor. In cases where disagreement exists between data provided in this document and that provided in the references given in section 5, this document governs.

2.2

PLANT CAPACITY – NUMBER OF TRAINS The AGR unit is designed for a maximum flow per train (Design case) of 13036** kg moles/hr (7.4 million Sm3/day) of sour gas corresponding to Phase 2 winter operation. It handles the sour gas delivered from offshore supplemented with flashed sour gas from the oil trains recompressed to AGR inlet pressure. The AGR unit consists of one train capable of processing 100% of the total onshore sour gases for firstly Phase 1. A second identical train is required for Phase 2. There are several options for the future some of which involve the possible addition of a third identical AGR train at some future date.

2.3

SOUR GAS FEEDSTOCK CHARACTERISTICS The inlet gas composition to the AGR Unit is as follows. It is based on the KE-Revised Composition. Flows shown refer to the total feed to the Absorber. The cases below give sour gas for the following operations Phase 1 , summer Phase 1 , winter Phase 2 , summer Phase 2, winter, DESIGN case There is also a small acid gas removal service required in the NGL treating area- see below under section 3 for details. AGR shall be able to handle the composition range of these cases,

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DESIGN CASE PER TRAIN **Case No.

1

2

3

4

Case Description

Phase 1 Summer

Phase 1 Winter

Phase 2 Summer

Phase 2 Winter

Pressure, bar a Temperature, °C * Flow, kgmol/h Composition, mole % Hydrogen Sulphide Carbon Dioxide Nitrogen Methane Ethane Propane Butanes Pentanes Hexanes + Methyl Mercaptan Ethyl Mercaptan Propyl Mercaptan Butyl Mercaptan CS2 COS Water

67.5 52(44.1) 11,133

67.5 52(48.5) 11,943

67.5 52(43.8) 12,660

67.5 52(48.0) 13,036

18.96 4.86 1.05 56.36 9.30 6.07 2.56 0.53 0.19 95 ppmv 38 ppmv 8.4 ppmv 3.0 ppmv 2.4 ppmv 33 ppmv 0.11




*Temperatures shown are (inlet) and outlet of super heater –use 52°C for design. There is no liquid in the gas; hydrocarbon dew point is less than 40°C. BTX/EB level is approx 60 ppmv, max. **Figures presented under various cases are for illustration purpose only.

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These compositions are taken from the following Material Balances. Main Onshore Facilities Phase 1 Summer Phase 1 Winter Phase 2 Summer Phase 2 Winter

Doc no. xx Doc no. xx Doc no. xx Doc no. xx

Design case is Phase 2 Winter.

2.4

TURNDOWN RATIOS Each AGR unit shall be designed to operate with a turndown ratio at 30% of design capacity.

2.5

PRODUCT SPECIFICATIONS

2.5.1

Sweet Gas The treated gas shall comply with the following specifications for the whole range of feedstock characteristics. • • • •

2.5.2

3 in

To achieve H2S less than 7 mg/Nm the sales gas the treated gas from the AGR must be treated to less than 6.3 mg/Nm3. CO2 content: No specification for CO2 , expected to be in the ppm range, data to be provided by licensor Mercaptan, COS, CS2 and BTX/EB contents: Expected removal levels to be provided by Licensor Outlet condition: To be assessed by Licensor

Acid Gas The acid gas from the AGR solvent regeneration will be processed in a Claus type Sulphur Recovery Unit. The process licensor shall be responsible for definition of the range of characteristics of the acid gas for SRU particularly H2S content and hydrocarbons content. Licensor should reduce the hydrocarbon content of the acid gas to a minimum. The battery limit conditions at the AGR’s outlet shall be as follows: • Pressure: 2.1 bar abs minimum • Temperature: 50 °C (normal) The Sulphur Recovery Unit shall be designed for a normal value of 0.22 vol% and a maximum of 0.5 vol% of Hydrocarbons the overall feed acid gas.

2.5.3

Flash Gas Specification The flashed gas produced by the AGR’s, at rich amine flash drum outlet, shall meet the following specification: • Pressure: by Process Licensor • Temperature: by Process Licensor Licensor needs to minimize H2S content and advise suitably.

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2.6

BATTERY LIMITS DEFINITIONS AND CONDITIONS

Temperature C Max or min 52 max with super-heater on

Pressure baraoperating 67.5

Pressure bara Max or min 80 bar g mechanical design

*

*

50

*

*

To be 66.9 bara minimum To be 2.1 minimum

After Fuel Gas Scrubber and any control valve provided by Licensor

*

*

*

Discharge of Licensor pump

50

*

26

To be 8bara minimum to supply LP fuel gas at 5.5 bara operating and 7.2barg design *

Exit NGL scrubber

52

55

23

24

*

*

*

*

Battery Limit Definition

Temperature C operating

Inlet to Filter Coalescer provided by Licensor Exit from KO drum provided by Licensor After KO drum and any control valve provided by Licensor

See above cases 44 to 48 C with super-heater off *

Flash Gas

Lean amine to NGL treating unit

Stream Feed Gas

Product gas

Acid Gas

Rich amine from NGL treating unit Utilities -see Appendix 1 Hydrocarbon and water condensate Fresh amine

*

Downstream LC on Licensor Filter coalescer *

* Figures need to be specified by Licensor 3.

DESIGN REQUIREMENTS The following requirements must be taken into account for the design of the Acid Gas Removal Units.

3.1

GENERAL The Process Licensor shall provide a design that meets the guarantee statements as specified in section 4 (Process Guarantees). The Process Flow Diagram shall be optimised by the Process Licensor, in close cooperation with the Engineering Contractor, with regards to process design, equipment constructibility, transportation and erection as well as plant operation reliability and availability.

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Process Licensor responsibility is limited to the items within the AGR’s Battery Limit, taking however in consideration the general arrangement of the whole Plant. The Process Licensor is also responsible for the definition of the acid gas composition range to be used for the design of the Sulphur Recovery Units according to the operating cases defined for the Acid Gas Removal Units.

3.2

METALLURGY The licensor shall recommend the minimum materials of construction for the equipment and piping. The metallurgy selection shall consider corrosive process fluid (as per the design basis) and plant service life of 40 years. A corrosivity assessment, including corrosion calculations shall be carried out for all sections of plant associated with wet CO2 and H2S, as part of the materials selection justification. The corrosivity assessment shall assess only sections of the plant where internal corrosion can arise as result of the presence of CO2, H2S and chlorides in combination with a distinct water phase. For cost effectiveness, the primary choice material considered shall be LTCS. Alternative materials are selected in cases where LTCS is incapable of providing the required design life or where temperatures could be encountered during service that is below the permissible minimum temperature for LTCS.

3.3

UTILITIES AND CHEMICALS

3.3.1

Utilities Process Licensor to specify the required utilities quality and consumption/production in all operating cases, including peak requirement for start-up, shut down or specific AGR’s units operation.

3.3.2

Chemicals Process Licensor to specify all chemicals (solvent included) used in the Acid Gas Removal Units, the consumption, life time, inventory and initial load to purchase.

3.4

EQUIPMENT DESIGN REQUIREMENTS Licensor equipment recommendations shall be suitable for obtaining mechanical guarantees from equipment suppliers. The design of the AGR unit shall incorporate the following specific requirements: •

The top section of the Absorber shall include a wash water section to minimize solvent losses.

•

Three off 50% Lean Amine Pumps (plus Booster Pumps) shall be provided.

•

The Intermediate Storage Tank shall be sized to accommodate the whole amine inventory.

•

The top section of the Regenerator shall include a packed column (contact condenser). A circulating reflux system with an air cooler will be used.

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•

Lean amine shall also be provided for NGL treating. The requirement is for a pump discharge pressure of 26 bara and a temperature of 50 °C. Rich amine at a loading level not exceeding 0.05 mol H2S per mol amine will be returned to the Rich Amine Flash Drum. Compositions and flowrate requirement are given below.

Amine flows to and from COS treater

Amine flows to and from COS treater in

*Mass % Nitrogen H2S CO2 Methane Ethane Propane i-Butane n-Butane i-Pentane n-Pentane n-Hexane n-Heptane n-Octane n-Nonane n-Decane n-C11 n-C12 H2O M-Mercaptan E-Mercaptan CS2 COS DEAmine Temperature [C] Pressure [bar] Molar Flow [kgmole/h] Mass Flow [kg/h]

0.006 0.076 0.001 59.924 39.993

964.24 26,000

out 0.000 0.112 0.048 0.001 0.091 0.001 0.001 0.000 0.000 0.001 0.000 0.000 0.006 59.661 40.078

959.78 25,947

ppmwt out

10 915 6 9 2 2 6 0 0 62 -

in

out

*Mole % Nitrogen H2S CO Methane Ethane Propane i-Butane n-Butane i-Pentane n-Pentane n-Hexane n-Heptane n-Octane n-Nonane n-Decane n-C11 n-C12 H2O M-Mercaptan E-Mercaptan CS2 COS DEAmine

0.006 0.060 0.001 89.677 10.257

0.000 0.089 0.029 0.001 0.056 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.001 89.517 10.305

Temperature [C] Pressure [bar] Molar Flow [kgmole/h] Mass Flow [kg/h]

964.24 26,000

959.78 25,947

ppm mol out

9 561 3 4 1 1 2 0 0 13 -

*Figures mentioned above are for illustration purpose only. 4.

PROCESS GUARANTEES

4.1

GENERAL The performances shall be guaranteed under the following conditions:

Feed gas composition and conditions within the range defined in Par. 2.3. Unit engineered and erected according to the Process Licensor recommendations as provided in the AGR's Process Design Package and the Operating Guidelines. Unit under stable and normal operation according to the Process Licensor recommendations as provided in the AGR's Process Design Package and the Operating Guidelines

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4.2

PERFORMANCE GUARANTEES The AGR Licensor shall guarantee the following: • • • • • • • • •

•

3

AGR design capacity: 13036 kg moles/hr (7.4 Sm /d) – see Design case AGR turndown capacity: 30% Hydrocarbon content of Acid Gas not to exceed 0.5 mol% Minimum Acid Gas pressure 2.1 bara Process Gas pressure drop (Feed minus Treated Gas pressure) not to exceed 0.6 bar For the whole range of operating cases, the treated Gas specification: H2S: less than 6.3 mg/Nm3 For the whole range of operating cases, the flash gas specification: H2S: To be defined by the Licensor. Solvent losses: To be defined by the Licensor. Utility requirements Maximum steam rates: To be defined by Licensor. Maximum power consumption: To be defined by Licensor. Maximum flash gas production rate.

The AGR Licensor is also responsible to advise the data to be taken into account for the following: •

The acid gas range of characteristics (H2S, CO2, hydrocarbons including BTEX content and breakdown, etc.) for the design of the SRU’s.

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APPENDIX-2 : TYPICAL PROCESS FLOWSHEET OF AMINE ABSORTION UNIT PREPARED IN HYSYS

Amine regen tower

Amine absorption tower

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Gas Sweetening Write-up - Final

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