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In Oil & Gas industry gas sweetening process is inevitable when raw natural gas contains acid gases like H2S and CO2. Removal of these acid gases is essential since their presence poses severe corrosion problem to the downstream process lines and equipment. Primary, Secondary and Tertiary amines are widely used as solvent for sweetening purpose. This compilation aims at providing general insight to the sweetening process, associated operational problems and troubleshooting measures.
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A Technical Report
Gas Sweetening by Amines
Subhasish Mitra, Sr. Process Engineer
Petrofac Engineering (I) Ltd, Mumbai, India
Gas sweetening by amine
1.0 Introduction 3
2.0 Gas sweetening basics 7
3.0 Alkanolamine gas treatment basics 8
4.0 Alkanolamine gas treating chemistry 13
5.0 Alkanolamine processes-strengths and weakness/solvent selection 20
6.0 Amine system description 24
7.0 Operational issues of amine sweetening system 33
8.0 Troubleshooting guide 41
9.0 Prevention of BTEX emission 46
10.0 Bulk CO
removal technology by membrane unit 47
11.0 New developments 49
Appendix - 1:
Typical process specification for gas sweetening package 50
Appendix - 2:
Typical process flow sheet for amine absorption unit prepared in 58
Hysys simulator package
Gas sweetening by amine
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
Gas sweetening by amine
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 (H
Carbon di-oxide (CO
Water vapor (H
Sulfur di-oxide (SO
Nitrogen Oxides (NO
Volatile Chlorine Compounds (HCl,Cl
Volatile fluorine compounds (HF, SiF
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
Gas sweetening by amine
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
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..
It is the selective concentration of one or more components of a gas at the surface of a micro-
porous solid. The mixture of adsorbed components is called the adsorbate, and the micro-
porous solid is the adsorbent. The attractive forces holding the adsorbate on the adsorbent are
weaker than those of chemical bonds, and the adsorbate can generally be released (desorbed)
by raising the temperature or reducing the partial pressure of the component in the gas phase
in a manner analogous to the stripping of an absorbed component from solution. When an
adsorbed component reacts chemically with the solid, the operation is called chemisorption
and desorption is generally not possible.
Membrane permeation:
It is a relatively new technology in the field of gas purification. In this process, polymeric
membranes separate gases by selective permeation of one or more gaseous components from
one side of a membrane barrier to the other side. The components dissolve in the polymer at
one surface and are transported across the membrane as the result of a concentration gradient.
The concentration gradient is maintained by a high partial pressure of the key components in
the gas on one side of the membrane barrier and a low partial pressure on the other side.
Although membrane permeation is still a minor factor in the field of gas purification, it is
rapidly finding new applications.
Chemical conversion:
It is the principal operation in a wide variety of processes, including catalytic and non-
catalytic gas phase reactions and the reaction of gas phase components with solids. The
reaction of gaseous Species with liquids and with solid particles suspended in liquids is
considered to be a special case of absorption and is discussed under that subject.
Gas sweetening by amine
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 H
S and CO
from gas streams by using absorption technology and chemical solvents. Sour gas contains
, H
O, hydrocarbons, COS/CS
, solids, mercaptans, NH
, 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 CO
, all H
S, or mixed and in what proportion,
if CO
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
Table 3: Selection of treatment process
Both absorption in alkalime solution (e.g., aqueous diethanolamine) and absorption in a
physical solvent (e.g., polyethylene glycol dimethyl ether) are suitable process techniques for
treating high-volume gas streams containing hydrogen sulfide andor carbon dioxide.
However, physical absorption processes are not economically competitive when the acid gas
partial pressure is low because the capacity of physical solvents is a strong function of partial
Gas sweetening by amine
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
Membrane permeation is particularly applicable to the removal of carbon dioxide from high-
pressure gas. The process is based on the use of relatively small modules, and an increase in
plant capacity is accomplished by simply using proportionately more modules. As a result,
the process does not realize the economies of scale and becomes less competitive with
absorption processes as the plant size is increased.
At very high acid-gas concentrations (over about 15% carbon dioxide), a hybrid process
(amine + membrane) proved to be more economical than either type alone. The hybrid
process uses the membrane process for bulk removal of carbon dioxide and the amine process
for final cleanup.
When hydrogen sulfide and carbon dioxide are absorbed in alkaline solutions or physical
solvents, they are normally evolved during regeneration without undergoing a chemical
change. If the regenerator off-gas contains more than about 10 tons per day of sulfur (as
hydrogen sulfide), it is usually economical to convert the hydrogen sulfide to elemental sulfur
in a conventional Claus-type sulfur plant. For cases that involve smaller quantities of sulfur,
because of either a very low concentration in the feed gas or a small quantity of feed gas,
direct oxidation may be the preferred route.
Direct oxidation can be accomplished by absorption in a liquid with subsequent oxidation to
form slurry of solid sulfur particles or sorption on a solid with or without oxidation. The solid
sorption processes are particularly applicable to very small quantities of feed gas where
operational simplicity is important, and to the removal of traces of sulfur compounds for final
cleanup of synthesis gas streams. Solid sorption processes are also under development for
treating high temperature gas streams, which cannot be handled by conventional liquid
absorption processes.
Adsorption is a viable option for hydrogen sulfide removal when the amount of sulfur is very
small and the gas contains heavier sulfur compounds (such as mercaptans and carbon
disulfide) that must also be removed. For adsorption to be the preferred process for carbon
dioxide removal there must be a high CO
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 (H
S), carbon dioxide
), 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.
Gas sweetening by amine
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 (H
S and CO
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
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.
Gas sweetening by amine
A small packed tower with a lean amine wash may be installed on top of the flash drum to
remove H
S 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,
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
Gas sweetening by amine
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
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 CO
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
Gas sweetening by amine
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 CO
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 CO
quite slowly and also have a low heat of solution.
Gas sweetening by amine
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% CO
and a trace of H
S 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 (H
S) and carbon dioxide (CO
) 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 H
S and/or CO
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
Gas sweetening by amine
and regenerating the amine solution for reuse, thus removing the acid gas from the inlet gas
The alkanolamines are classified by the degree of substitution on the central nitrogen; a single
substitution denoting a primary amine, a double substitution, a secondary amine, and a triple
substitution, a tertiary amine. Each of the alkanolamines has at least one hydroxyl group and
one amino group. In general, the hydroxyl group serves to reduce vapor pressure and increase
water solubility, while the amine group provides the necessary alkalinity in water solutions to
promote the reaction with acid gases. It is readily apparent looking at the molecular structures
that the non-fully substituted alkanolamines have hydrogen atoms at the non-substituted
valent sites on the central nitrogen, whereas the tertiary amines are fully substituted on the
central nitrogen. This structural characteristic plays an important role in the acid gas removal
capabilities of the various treating solvents.
Amines which have two hydrogen atoms directly attached to a nitrogen atom, such as MEA
and DGA, are called primary amines and are generally the most alkaline. DEA and DPA have
one hydrogen atom directly attached to the nitrogen atom and are called secondary amines.
TEA and MDEA represent completely substituted ammonia molecules with no hydrogen
atoms attached to the nitrogen, and are called tertiary amines.
Primary amines:
Monoethanolamine (MEA) DIGLYCOLAMINE Agent (DGA)
- NH
Secondary amines
Diethanolamine (DEA) Diisopropanolamine (DIPA)
OH - NH - C
OH - NH- C
Tertiary amines
Triethanolamine (TEA) Methyldiethanolamine (MDEA)
OH - NH - C
OH - NH - C
Gas sweetening by amine
Figure 5: Structural formulae of Alkanolamines used in gas treating
In an aqueous solution, H
S and CO
dissociate to form a weakly acidic solution.
Ionization of water:
O = H
+ OH
Ionization of dissolved H
S = H
+ HS
Hydrolysis and ionization of dissolved CO
+ H
+ H
When a gas stream containing H
S and/or CO
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 H
S and CO
is exothermic and a considerable amount of
heat is liberated. Regardless of the structure of the amine, H
S 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 + H
S R1R2R3NH+ HS - Equation 1
The reaction between the amine and CO
is a bit more complex because CO
absorption can
occur via two different reaction mechanisms. When dissolved in water, CO
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:
+ H
(Carbonic Acid) - Equation 2
(Bicarbonate) - Equation 3
Gas sweetening by amine
+ R1R2R3N R1R2R3NH
-Equation 4
+ H
- 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 CO
reaction mechanism as shown by Equation 3
below requiring the presence of labile hydrogen in the molecular structure of the amine may
also occur.
- Equation 6
+ R1R2NH
- Equation 7
+ 2R1R2NH R1R2NH
- Equation 8
This second reaction mechanism for CO
, 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 CO
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 CO
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 CO
is limited to 0.5 mole of CO
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 CO
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 CO
absorption may occur by two
reaction mechanisms with significantly different kinetic characteristics has a great impact
upon the relative absorption rates of H
S and CO
among the different alkanolamines.
For primary and secondary amines, very little difference exists between the H
S and CO
reaction rates. This rate equivalence is due to the availability of the rapid carbamate
formation reaction for CO
absorption. Therefore, the primary and secondary amines achieve
essentially complete removal of H
S and CO
. However, because the tertiary amines are fully
substituted, they can not form the carbamate. Tertiary amines must react with CO2 via the
slow CO
hydrolysis mechanism discussed earlier. For MDEA, since the CO
reaction with
water to form bicarbonate is slow and the H
S reaction is fast, it is generally felt that the H
reaction is gas phase limited while the CO
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 H
S relative to CO
A little insight to the solubility phenomenon of acid gases (H
) 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:
P = γ
(i = H
) - Equation 9
Gas sweetening by amine
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 H
S and
are only considered weak acids, very little ionization occurs unless a basic compound
(such as an amine) is also present. Taking H
S as an example, the total equivalent H
S in the
aqueous phase will be the sum of free physically dissolved H
S, bisulfide ion (HS
), and
sulfide ion (S
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
Gas sweetening by amine
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 H
S in the
presence of a weak base such as MDEA, a dilute solution is assumed where activity
coefficients (γ) are unity. The total solution H
S 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,
Gas sweetening by amine
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 H
S present the solution shifts from free
physically dissolved H
S to bisulphide (HS
) ions and ultimately to sulphide (S
) ions. This
drives the equilibrium towards dissolving more total H
S. 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 H
S & MDEA ions v/s pH
Gas sweetening by amine
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 CO
removal to 100 ppmv for applications at low to moderate operating
Some of the disadvantages of MEA are:
High solvent vapor pressure which results in higher solvent losses than the other
• Higher corrosion potential than other alkanolamines,
• High energy requirements due to the high heat of reaction with H2S and CO
• Nonselective removal in a mixed acid gas system, and
• Formation of irreversible degradation products with CO2, COS and CS2, which requires
continuous reclaiming.
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 H
S and CO
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 CS
besides H
S and CO
, such as refinery gas streams, can generally be treated successfully.
The advantages of DEA are:
• Resistance to degradation from COS and CS
• Low solvent vapor pressure which results in potentially lower solvent losses,
• Reduced corrosive nature when compared to MEA, and
• Low solvent cost.
Gas sweetening by amine
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 CO
• Higher circulation requirements, and
• Non-reclaimable by conventional reclaiming techniques.
Degradation products resulting from the reaction of DEA and CO
at elevated temperatures
include hydroxyethyloxazolidone-1,dihydroxyethylpiperazine,3-(2-ydroxyethyl)oxazolidone-
2(HEOD), N,N.bis(2-hydroxyethyl) piperazine (BHEP) and N,N,N’-tris(2-hydroxyethyl)
ethylenediamine (THEED).
An explanation for 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 CS
• 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 CO
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.
Gas sweetening by amine
DGA Agent reacts with CO
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:
+ (CO
or COS) (R-NH)
CO + (H
O or H
+ (COS or CS
) (R-NH)
CS + (H
O or H
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 CO
or COS. A second by-product can also
form by the reaction of 1 mole of either CS
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 CO
, COS, or CS
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 H
when treating a gas stream containing both H
S and CO
while slipping a significant portion
of the CO
. This slippage of CO
can be useful in applications requiring the upgrading of H
content for sulfur plant feed gas or adjusting the CO
content of the treated gas while at the
same time removing H
S 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 H
S over CO
in mixed acid gas applications, Essentially complete H
removal while only a portion of CO
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, CS
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 SO
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
Gas sweetening by amine
concern when degradation products are formed and bicine is present. As with all
alkanolamines, the presence of oxygen increases the likelihood of product degradation and
corrosion concerns.
Table- 5: Comparative Study of Solvents:
Solvent Name
MEA (Mono
Ethanol Amine )
DEA (Di-
Ethanol Amine)
DGA (Di-Glycol
Amine Agent)
MDEA (Methy Di
Ethanol Amine)
Solvent Cost Low Solvent Cost Low Solvent Cost
Relatively high
solvent cost Highest Solvent Cost
Solvent Loss
High solvent
vapor pressure
results in higher
solvent loss.
Low solvent
vapor pressure
results potentially
lower solvent
Low vapor pressure
which results in
potentially low
solvent loss.
removal in a
mixed acid gas
Partial removal of
COS and CS2.
removal in a
mixed acid gas
removal in a mixed
acid gas system.
Removal of COS
and CS2.
Selectivity of H2S
over CO2 in mixed
acid gas applications.
Essentially complete
H2S removal while
only a portion of CO2
is removed enriching
the acid gas feed to
the sulfur recovery
Minimal COS and
CS2 removal.
Good Thermal
Excellent Thermal
High reactivity
due to its primary
Low reactivity
compared to
High reactivity, 1/4
grain H2S
specification can
generally be
obtained for
applications with
low operating
pressures & high
Lower comparative
Higher Corrosion
corrisive nature
compared to
MEA. Less corrosive
(Reclaimation )
by conventional
Non-reclaimable by
Gas sweetening by amine
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
Gas sweetening by amine
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 2½-
micron level. In applications where it is anticipated that the inlet gas may contain particulate
such as FeS, a filter-separator may be required. This equipment typically consists of a
horizontal vessel with filters in the inlet end of the vessel to remove the FeS followed by mist
pads or impingement baffles with a separator chamber to collect any separated liquids.
Aerosols, which may be as small as 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 % C
+) 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 H
and CO
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 % C
+) 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.
Gas sweetening by amine
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 H
S equilibrium problems on the top tray of the absorber or
increased solution losses due to excessive vaporization losses.
A differential pressure instrument should be installed on the absorber and stripper tower to
monitor the differential pressure across the trays or packing. The differential pressure should
be measured from just below the first tray or section of packing to just above the last tray or
section. A sharp increase in the absorber/stripper differential pressure is an excellent
indication that a foaming problem exists in the system. The typical absorber design does not
usually include a provision for several water wash trays (2-4 trays) above the last amine-
contacting tray to reduce amine entrainment/carryover into the sweet gas residue. However,
Gas sweetening by amine
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
S is preferred to CO
, then packed tower becomes the obvious choice. H
S reacts much
faster with the solvent than CO
and this aspect of the reaction kinetics is employed in packed
tower which owing to low hold up provides shorter contact time between the phases to
achieve preferential absorption.
Table 7, gives a comparison between performances of both type of towers for similar
operating conditions.
Table 7: Tray vs Packing in selective removal application
Although bubble-cap trays and raschig ring packings were once commonly used in amine
plant absorbers and strippers, modem plants are generally designed to use more effective
trays (e.g.. sieve or valve types) and improved packing shapes (e.g., Pall rings or high-
performance proprietary designs). Very high-performance structured packing is seldom used
for large commercial gas treating plants because of its high cost and sensitivity to plugging
by small particles suspended in the solution. The choice between trays and packing is
somewhat arbitrary because either can usually be designed to do an adequate job, and the
overall economics are seldom decisively in favor of one or the other. At this time, sieve tray
columns are probably the most popular for both absorbers and strippers in conventional, huge
commercial amine plants; while packed columns are often used for revamps to increase
capacity or efficiency and for special applications. Tray columns are particularly applicable
for high pressure columns, where pressure drop is not an important consideration and gas
purity specifications can readily be attained with about 20 trays. Packing is often specified for
removal columns, where a high degree of CO
removal is desired and the low efficiency
of trays may result in objectionably tall columns. Packing is also preferred for columns where
Gas sweetening by amine
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 H
S 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.
Gas sweetening by amine
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
S in the treated gas leaving these trays. Therefore, in these circumstances, the H
S 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 CO
however, more than 20 trays may be required if CO
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 CO
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
Gas sweetening by amine
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.
Gas sweetening by amine
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
Gas sweetening by amine
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 H
S and CO
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 CO
. The regeneration requirement to reach a typical lean
loading is a reflux ratio of 1.0 to 3.0. A reflux ratio of 1.0 should be considered as a practical
minimum. In some low pressure or tail gas treating applications, higher reflux ratios may be
required to meet the product specifications. In order to ensure adequate stripping while at the
same time optimizing energy utilization, control of the heat input to the reboiler should be
accomplished by monitoring the stripper overhead temperature. The overhead temperature
correlates directly with the reboiler energy input. The reboiler temperature is not affected by
the amount of stripping steam generated in the reboiler since the boiling point of the amine
solution is dependent upon the amine concentration and reboiler pressure. Therefore, the
reboiler temperature is not a controlled variable. The heat input to the reboiler should be set
to achieve a specified stripper overhead temperature, typically 210 to 230 °F depending upon
the gas treating application and amount of reflux desired. To prevent thermal degradation of
the amine solvent, steam or hot oil temperatures providing heat to the reboiler should not
exceed 350 °F. Superheated steam should be avoided. 50 psig saturated steam is
recommended. The maximum bulk solution temperature in the reboiler should be limited to
260 °F to avoid excessive degradation.
6.8 Filter:
A good filtration design includes both a particulate and a carbon filter. The cleaner the amine
solution, the better the amine system operates. The particulate filter is used to remove
accumulated particulate contaminants from the amine solution that can enhance foaming and
aggravate corrosion. Carbon filtration removes surface active contaminants and hydrocarbons
that contribute to foaming. With proper inlet gas separation and pre-treatment, filtering a 10
to 20 percent slipstream of the total lean solution has usually proven adequate. Where
practical, total stream filtration should be considered. The filtration system is typically
installed on the cool lean amine stream (absorber feed). Recirculation of a slipstream from
the discharge side of the charge pump to the filtration system with a return to the suction side
of the pump is a common arrangement. If combined in series, the particulate filter should be
installed upstream of the carbon filter to protect the carbon filter. A second post-filter or
screen should be installed downstream of the carbon filter to keep carbon fines out of the
circulating system. If the carbon filter is installed independent of the particulate filter, a pre-
filter should be installed on the carbon filter inlet to protect the carbon bed. In systems that
are extremely contaminated with particulate due to inadequate feed preparation, excessive
Gas sweetening by amine
corrosion, or if the inlet gas CO
S 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 H
S 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 H
6.9.1 Particulate filter:
The particulate filter should filter a minimum 10 to 20% slipstream of the circulating
solution. Numerous particulate filter mediums have been utilized in amine service: wound
bleached cotton disposable filter cartridges with polypropylene or metal cores, disposable
metal cartridges, pleated paper filter cartridges, sock-type disposable elements and non-
disposable/back-flushable mechanical filters with special metal etched filter elements.
Experience has shown that a 10-micron absolute filter is adequate for most amine
applications, although some MDEA applications as well as many refinery amine applications,
which are plagued by a black, shoe polish-like material consisting of iron sulfide bound with
hydrocarbon and polymerized amine, require more stringent filtration. The FeS-hydrocarbon
shoe polish-like material is very finely divided, with eighty percent of the FeS particles being
between 1 and 5 microns in size. 5-micron absolute filtration is typically recommended for
these applications.
6.9.2 Carbon filter:
Carbon filter is used in those in amine systems that experience severe emulsion problems due
to significant hydrocarbon contamination. A properly designed activated carbon (Activated
carbon with high iodine number i.e. high adsorption capacity, high abrasion number i.e.
abrasion resistance against degradation is preferred) system can reduce the need for antifoam,
reduce amine make up, reduce corrosion and improve scrubbing efficiencies and product
quality. The carbon system should treat at least 10 to 20% of the circulating lean amine
solution. A minimum contact time of 15 minutes and a superficial velocity of 2 to 4 gpm/sq ft
is considered appropriate. When the amine solution changes color or clarity or the solution
foaming tendency increases, the carbon is spent and should be changed. Typical maximum
carbon life is 6 to 9 months.
7.0 Operational Issues of Amine Sweetening System
A number of operational issues faced in amine gas treating units have been reported. Often
one operational difficulty can cause or influence another problem. Not all amine systems
experience the same degree of operating difficulties. A continual problem that afflicts one
amine system may occur only rarely in another amine system. Several of the more common
operational difficulties encountered are discussed below along with troubleshooting
recommendations and design considerations whose aim is to improve the amine unit
operations and control these common operational problems.
7.1 Failure to meet product specification
Difficulty in satisfying the product specification, typically the H
S specification whether the
treated stream is a liquid or a gas may be the result of poor contact (loss of efficiency)
Gas sweetening by amine
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.
Gas sweetening by amine
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 CO
S ratio.
7.2.3 Heat Stable Salts:
HSS, which are the reaction products of the amine and acids stronger than H
S and CO
which do not dissociate in the regenerator and are therefore heat stable, are corrosive and
increase the corrosivity of the solution. Historically, a rule of thumb has been utilized
limiting the HSS to 5-10% of the amine alkalinity (for a 50-wt. % amine solution, the 5-
10% HSS limit corresponds to 2 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
Gas sweetening by amine
7.2.4 Elevated Temperatures:
High process temperatures tend to promote acid gas breakout as well as having an effect
on the amine solution pH, as the solution pH tends to drop with increasing temperature.
The rich amine feed temperature to the stripper is typically limited to 210-220 °F [98.9-
104.4 °C] to prevent acid gas breakout. Additionally, the amine solution can be degraded
by excessive heat. Thermal degradation potential can be lessened by limiting the bulk
temperature of the reboiler amine to 260 °F [126.7 °C] and limiting the reboiler heating
medium temperature to 350 °F [176.7 °C]. Superheated steam should be avoided. 50 psig
[345 kPA] saturated steam is the preferred heating medium. Hot oil and direct fired
reboilers should be avoided if possible to avoid potential thermal degradation. If a hot oil
or direct-fired reboiler is necessary, care should be taken in the design of the reboiler.
7.2.5 High Velocities:
The velocity of the amine treating solution is limited to control corrosion/erosion caused
by the presence of solid particulates as well as acid gas flashing due to excessive pressure
drop. Amine solution velocities in the exchangers should be limited to 3 ft/sec [0.9
m/sec] while the velocity in the piping should be limited to 7 ft/sec [2.1 m/sec]. Long-
radius elbows should be utilized where practical in rich amine service.
7.3 Solution foaming:
Amine solution foaming is probably the most persistent and troubling operational problem
encountered in natural gas production and refinery sweetening operations. Solution foaming
contributes significantly to excessive solution losses through entrainment and amine
carryover, reduction in treating capacity through unstable operations and off-specification
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
Gas sweetening by amine
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. H
S 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 H
S, 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,
Gas sweetening by amine
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 H
base complex. The bound amine of the HSS can no longer react with the incoming acid gas;
thus the system capacity is reduced. The HSS content of the amine solution is determined by
an ion-exchange/titration method which determines the total equivalents of all anions present
in the solution and is reported as the amount of amine tied up in the form of amine salts. This
method does not distinguish between amine HSS and inorganic HSS (sodium [Na] or
potassium [K]). Comparison of the HSS by the titration method result with ion
chromatography (IC) results and cation analysis which determines Na and K helps determine
to what extent HSS are present as amine salts versus inorganic salts. The HSS precursors
found in many refinery applications, such as the carboxylic acids (formic, oxalic and acetic)
responsible for HSS contamination, typically come from sources such as the FCC and coker
off gas. The principal culprit in HSS formation (thiosulfate) in tail gas amine systems is due
to SO
breakthrough past the hydrogenation reactor and the tail gas quench water system. A
few of the more commonly found HSS anions are acetate (CH
), formate (HCOO
Gas sweetening by amine
thiocyanate (SCN
), sulfate (SO
) and thiosulfate (S
). 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 (Na
) as well as KOH and K
.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 + H
O + Na(K)HSS
+ 2H
O + 2 Amine H+ HSS
2 Amine + 2H
O +2CO
+ 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 K
, 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
Gas sweetening by amine
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/by-
product 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
Gas sweetening by amine
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 H
S 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 H
S and
, 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 H
S. 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 CO
: H
S 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.
attack is much more insidious than H
S attack. H
S attack tends to corrode a larger area
in a kind of wash pattern, while CO
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
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
Gas sweetening by amine
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
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.
Gas sweetening by amine
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
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
For applications utilizing MDEA and formulated MDEA solvents which have a CO
specification such as high-pressure applications with a CO
spec of 2-3 %, if the CO
of the treated gas is high, following need to be checked,
Lean Amine Solution Temperature:
The CO
-MDEA absorption mechanism is temperature dependent. CO
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 CO
removal. Otherwise, the troubleshooting checklist for high H
S in the residue gas is
applicable as described above.
If the CO
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.
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
Gas sweetening by amine
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 H
S 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 H
S 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 H
S 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
Gas sweetening by amine
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 long-
chained 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
Gas sweetening by amine
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. H
S and
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,
Gas sweetening by amine
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 H
S; 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 CO
removal technology by membrane unit:
Of late membranes have found their use in bulk removal of CO
from natural gas. As high as
30% CO
concentration in natural gas has been reported to be successfully processed in
membrane units to reduce CO
content~ 10 -15%. Many often a hybrid system of membrane
and amine system is used to meet sales gas specification.
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. CO
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 CO
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 CO
removal. Cellulose acetate membranes
were initially developed for reverse osmosis, but now are the most rugged CO
membrane available. Membrane performance has been enhanced from selectivity aspect by
Gas sweetening by amine
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, CO
, H
S 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
Gas sweetening by amine
In high CO
-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 CO
, hence permeate CO
, 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
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 CO
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
Gas sweetening by amine
amines to proprietary amines developed with enhanced selectivity as an essential part of this
progress. For highly selective H
S 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 H
S removal to lower treated gas specifications.
However, these solvents are MDEA-based solvents. These solvents have other applications;
such as H
S removal from CO
enhanced oil recovery (ROR) enrichment processes. Solvents
for H
S selectivity are used for refinery systems with high CO
slip, tail gas treating, natural
gas treating, H
S 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.
Gas sweetening by amine
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 CO
, mercaptans and COS.
The AGR unit will also supply a small lean amine stream to treat NGL for removal of H
S and
in the liquid treating area and receive the rich amine.
AGR processes sour feed gas from the inlet reception facilities of the plant combined with
condensate stabilisation recompressed offgas. The feed gas contains H
and organic
sulphur compounds (mainly mercaptans). H
S and mercaptans shall be removed from the
gas prior to the export to pipeline. The AGR shall remove the H
S 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. CO
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
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.
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,
Gas sweetening by amine
**Case No. 1 2 3 4
Case Description Phase 1 Phase 1 Phase 2 Phase 2
Summer Winter Summer Winter
Pressure, bar a 67.5 67.5 67.5 67.5
Temperature, °C * 52(44.1) 52(48.5) 52(43.8) 52(48.0)
Flow, kgmol/h 11,133 11,943 12,660 13,036
Composition, mole %
Hydrogen Sulphide 18.96 19.38 18.82 18.94
Carbon Dioxide 4.86 4.80 4.41 4.38
Nitrogen 1.05 0.99 1.12 1.09
Methane 56.36 54.54 54.85 53.88
Ethane 9.30 9.54 9.13 9.23
Propane 6.07 6.64 7.76 8.06
Butanes 2.56 3.10 2.94 3.51
Pentanes 0.53 0.67 0.56 0.54
Hexanes + 0.19 0.19 0.23 0.15
Methyl Mercaptan 95 ppmv 116 ppmv 106 ppmv 131 ppmv
Ethyl Mercaptan 38 ppmv 45 ppmv 39 ppmv 40 ppmv
Propyl Mercaptan 8.4 ppmv 9.9 ppmv 9.0 ppmv 6.9 ppmv
Butyl Mercaptan 3.0 ppmv 3.6 ppmv 3.4 ppmv 2.1 ppmv
2.4 ppmv 2.8 ppmv 2.4 ppmv 2.5 ppmv
COS 33 ppmv 36 ppmv 36 ppmv 37 ppmv
Water 0.11 0.19 0.16 0.20
*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.
Gas sweetening by amine
These compositions are taken from the following Material Balances.
Main Onshore Facilities Phase 1 Summer Doc no. xx
Phase 1 Winter Doc no. xx
Phase 2 Summer Doc no. xx
Phase 2 Winter Doc no. xx
Design case is Phase 2 Winter.
Each AGR unit shall be designed to operate with a turndown ratio at 30% of design capacity.
2.5.1 Sweet Gas
The treated gas shall comply with the following specifications for the whole range of feedstock
To achieve H
S less than 7 mg/Nm
3 in
the sales gas the treated gas from the AGR
must be treated to less than 6.3 mg/Nm
content: No specification for CO
, expected to be in the ppm range, data to
be provided by licensor
Mercaptan, COS, CS
and BTX/EB contents: Expected removal levels to be
provided by Licensor
Outlet condition: To be assessed by Licensor
2.5.2 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 H
S 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 H
S content and advise suitably.
Gas sweetening by amine
* Figures need to be specified by Licensor
The following requirements must be taken into account for the design of the Acid Gas
Removal Units.
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.
Stream Battery Limit
Temperature C
Max or min
Max or min
Feed Gas Inlet to Filter
provided by
See above
44 to 48 C with
super-heater off
52 max with
67.5 80 bar g
Product gas Exit from KO
drum provided
by Licensor
* * * To be 66.9
Acid Gas After KO drum
and any control
valve provided
by Licensor
50 * * To be 2.1
Flash Gas After Fuel Gas
Scrubber and
any control
valve provided
by Licensor
* * * To be 8bara
minimum to
supply LP
fuel gas at
5.5 bara
and 7.2barg
Lean amine to
NGL treating
Discharge of
Licensor pump
50 * 26 *
Rich amine
from NGL
treating unit
Exit NGL
52 55 23 24
Utilities -see
Appendix 1
and water
LC on Licensor
Filter coalescer
* * * *
Fresh amine *
Gas sweetening by amine
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.
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 CO
S, 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 CO
, H
S and chlorides in combination with a distinct water
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.
3.3.1 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
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.
Licensor equipment recommendations shall be suitable for obtaining mechanical guarantees
from equipment suppliers. The design of the AGR unit shall incorporate the following specific
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
The top section of the Regenerator shall include a packed column (contact
condenser). A circulating reflux system with an air cooler will be used.
Gas sweetening by amine
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 H
S 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 out ppmwt out
in out ppm mol out
*Mass % *Mole %
0.006 0.000 Nitrogen 0.006 0.000
S 0.076 0.112 H
S 0.060 0.089
0.001 0.048 CO
0.001 0.029
Methane - - Methane - -
Ethane - 0.001 10 Ethane - 0.001 9
Propane - 0.091 915 Propane - 0.056 561
i-Butane - 0.001 6 i-Butane - 0.000 3
n-Butane - 0.001 9 n-Butane - 0.000 4
i-Pentane - 0.000 2 i-Pentane - 0.000 1
n-Pentane - 0.000 2 n-Pentane - 0.000 1
n-Hexane - 0.001 6 n-Hexane - 0.000 2
n-Heptane - 0.000 0 n-Heptane - 0.000 0
n-Octane - 0.000 0 n-Octane - 0.000 0
n-Nonane - 0.006 62 n-Nonane - 0.001 13
n-Decane - - - n-Decane - - -
n-C11 - - n-C11 - -
n-C12 - - n-C12 - -
O 59.924 59.661
O 89.677
M-Mercaptan - - M-Mercaptan - -
E-Mercaptan - - E-Mercaptan - -
CS2 - - CS2 - -
COS - - COS - -
DEAmine 39.993 40.078
DEAmine 10.257
Temperature [C] Temperature [C]
Pressure [bar] Pressure [bar]
Molar Flow
[kgmole/h] 964.24 959.78
Molar Flow
[kgmole/h] 964.24
Mass Flow [kg/h] 26,000 25,947
Mass Flow [kg/h] 26,000
*Figures mentioned above are for illustration purpose only.
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
Gas sweetening by amine
The AGR Licensor shall guarantee the following:
AGR design capacity: 13036 kg moles/hr (7.4 Sm
/d) – see Design
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:
S: less than 6.3 mg/Nm3
For the whole range of operating cases, the flash gas specification:
S: 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
The acid gas range of characteristics (H
, hydrocarbons including BTEX
content and breakdown, etc.) for the design of the SRU’s.
Gas sweetening by amine
Amine absorption tower
Amine regen tower
... In processes such as natural gas dehydration [1], CO 2 capture from flue gas streams [2], natural gas and biogas upgrading [3], and in the development of renewable energies [4], acid-gas capture and sequestration is of integral importance. While initial investigations into gas sweetening were concerned with the removal of H 2 S to very low concentrations to avoid catalyst poisoning [5,6] and to reduce sulphur emissions from combustion, [6] removal of CO 2 from gas streams has become the focus of much experimental effort in later years. ...
... Relative to MEA and DEA, MDEA has gained a significant share of the market in the past decade due to several advantageous properties: selectivity towards H 2 S when treating H 2 S -CO 2 multicomponent acid-gas streams; low solvent vapour pressure; low corrosivity; high degradation resistance; and efficient energy utilisation. [1] To mitigate the research, plant, and operating costs associated with developing optimal solvent blends, thermodynamic models that accurately describe the vapour-liquid equilibrium (VLE) and thermodynamics of water-amine-gas systems is a prerequisite [10,11]. ...
Reactive absorption of CO2 and H2S in aqueous methyldiethanolamine (MDEA) solutions is considered within the ePC-SAFT equation of state. We demonstrate that ePC-SAFT can be employed in a predictive manner without regression of additional temperature-correlated terms. Mixed system predictions are tested using a consistent set experimental data covering a wide range of temperatures (313 K–413 K), partial pressures (0.001 kPa–1000 kPa), and MDEA mass fractions (0.05–wMDEA 0.75 wMDEA). Predicted partial pressures for acid gas absorption show good agreement for low MDEA fractions (wMDEA < 0.5). Absorption selectivity in binary H2S + CO2 absorption is correctly predicted, with absolute average deviations of 57.18% and 79.32% for partial pressures of CO2 and H2S. We identify a significant deterioration in ePC-SAFT predictive power for the high-MDEA regime (wMDEA > 0.5), likely originating from underlying assumptions in the Debye-Hückel electrolyte free energy treatment and representation of ionic species.
... As mentioned, sulfur, generally in terms of weight percent hydrogen sulfide, and carbon dioxide content must be limited per industry standards. If not removed before consumption, these acidic components would cause problems with freezing and corrosion and create significant environmental, health, and safety hazards [42]. To sequester CO 2 and H 2 S from valuable fuel streams, the following reactions are employed: ...
... The return requirement from ExxonMobil stated that a maximum 0.4 mol of combined H 2 S and CO 2 per mol MEA, which was used to calculate the amount of MEA (and MEA solution) needed to remove all of the H 2 S and CO 2 from the system. The computed mass flow of water and MEA was input to an Aspen HYSYS R model of a sweetening column ( Figure 7), which was iteratively sized using sensitivity testing on the number of stages in the column and the amount of MEA entering the system [7,22,52,18,46,42]. ...
Technical Report
Full-text available
Coking facilities are refineries that upgrade residual oils from crude distillation to more valuable light hydrocarbons. These processes usually involve fractionation columns, which separate the coking products into fuel gas, liquefied petroleum gas, naphtha, and heavier ends. Using Aspen Plus and Aspen HYSYS , we designed a coking gas plant to separate light gases, C3s, C4s, and heavy mixing components in a sour feed stream originating from the top of a coking process fractionator. Our finalized design recovers nearly 100% of the fed refinery fuel gas, 92% propylene, 96% propane, 92% butylene, 74% butane, and 99% gasoline mixing component at full capacity. All product purities are above 90% with the exception of refinery fuel gas and butylene which have purities of 65% and 73% respectively. Product purity meets all prescribed specifications, including successful extraction of H2S and CO2 contaminants. The sale of these products corresponds to an annual market value of $292.8M. With annual fixed costs of $49.3M and variable costs of $70.3M, our plant yields an investor rate of return of 23.6%, a 0.5% improvement on the alternative design. This is based on an estimated total permanent investment The payback period is estimated at 4 years with an annual return on investment of 27.0% The primary design also has the added benefit of improved product qualities and simplification of the sour gas sweetening system. Investment viability is most sensitive to product sales price and operating costs. The sensitivity of sales is of particular concern given the current economic landscape surrounding petroleum futures. Sour gas and the amine treatment required for H2S removal represent the primary safety concerns in the process, scoring 80/125 in a system FMEA analysis and achieving the maximum weighted environmental index factor using Biwer’s method when compared to other process components. These concerning elements were mitigated through process control implementation or on-site handling outside of the scope of this project.
... Taken together, the processes are popularly called 'gas sweetening'. The most popular options used around the world include absorption, adsorption, oxidation, and membrane permeation [33]. The absorption process driving force may be physical or chemical, the former when only physical interactions are involved and the latter in the case where a chemical reaction is present [34]. ...
The article presents research on the purification process of pyrolysis gas from hydrogen sulfide. To our best knowledge, similar studies have not been performed yet on a real pyrolysis gas obtained from waste tyres, which contains huge amounts of hydrogen sulfide - average 3.6% (up to 5.1% at 420 °C). Different sorbents were tested, among them sodium hydroxide, zinc oxide, and manganese oxide. NaOH in concentration 0.05 M appeared to be the most efficient, showing ∼94% H2S removal efficiency under the conditions studied. ZnO features a better efficiency of H2S removal from hot gas (∼55%) than MnO. Furthermore, the combination of ZnO and a 0.05 M solution of NaOH was studied. The detailed composition of the pyrolysis gas was performed, too. The main components and sulfur-containing compounds, such as methyl mercaptan, carbonyl sulfide, and ethyl mercaptan, concentrations were measured. Predominantly, the gas consists of methane, hydrogen, ethane, ethene, carbon dioxide, iso-butane, and hydrogen sulfide. Aggregated concentrations of the above-mentioned exceed 80% of the gas, which makes it a very promising gaseous fuel.
... From an engineering perspective, the anime regeneration units have high importance in the chemical absorption process [108]. The gas treatment process (P19) represents an absorber which is undoubtedly the single most crucial operation of gas purification processes [109]. Furthermore, in our case, the gas treatment process receives flows from multiple processes, and it is involved in a circular relation with the Anime (DEA) regeneration process, emerging its criticality. ...
Full-text available
Modern critical infrastructures are complex Cyber-Physical-Systems (CPS) that tightly integrate physical processes with Information and Communication Technology components. Numerous safety mishaps and security attacks in such systems have demonstrated the need to ensure their safety and security from early design stages. Research on CPS has mostly focused on securing existing, implemented industrial systems, while safety and security consideration during the design stages of modern industrial infrastructures has largely gone unnoticed. In this paper, we present a framework that extends previous, preliminary work on the integration of security in industrial engine-ering design practices, and provide an algorithmic approach that effectively reduces risk during industrial system design lifecycles. We achieve this by analyzing flows of materials and information related to physical processes using three steps: (1) Identifying critical compo-nents and flows, (2) prioritizing flows based on their ties to high risk and importance in terms of dependencies, and (3) classifying system components based on their influence on the overall industrial system. To do that, we utilize (i) material flow networks (MFN) for modell-ing/designing the physical system (ii) dependency risk graphs for analyzing networks dependencies and assessing the system, in terms of risk, (iii) graph minimum spanning trees, and (iv) network centrality metrics. To evaluate our approach, we model and assess the production chain corresponding to an oil refinery plant’s Liquefied Petroleum Gas (LPG) purification process. Preliminary findings demonstrate the complex dependencies between cybersecurity vulnerabilities and system safety.
... In the case of DES, Nano-DES, DES-MDEA, pure MDEA, and aqueous MDEA with NG, after heating step, these systems were equilibrated in the NVT ensemble for 30 ns to ensure equilibration (pressure of the systems was checked, as seen from Fig. S2) and then followed by 80 ns production run in the NPT ensemble at 310 K and 2 MPa. It should point out that since the physical absorption of acid gases from sour NG usually occurs in relatively high pressures ( > 1.4 MPa) [77], hence the pressure was set at the 2 MPa during the MD simulations. ...
Molecular dynamics simulations were performed to investigate the performance of three deep eutectic solvents (DESs) in the separation of acid gases from natural gas. To evaluate the efficiency of these DESs, their performance was compared with that of the conventional methyl diethanolamine (MDEA) system. For this purpose, the acidic DES (phenyl propionic acid and choline chloride; with 67 mole % Phpr), same acidic DES accompanied by boron-nitride nanotube (Nano-DES), and acidic DES with methyl diethanolamine (DES-MDEA) were considered. The solubility selectivity and diffusion selectivity were evaluated and compared to those of MDEA system. The Nano-DES system showed the best performance in terms of solubility selectivity of H2S over CH4, while the high solubility selectivity of CO2 over CH4 was obtained for pure DES systems. Likewise, the aqueous MDEA and the DES-MDEA systems exhibited the largest diffusivity selectivity of H2S over CH4 and CO2 over CH4, respectively. These results were rationalized and explained by analyzing the density profile and interaction energies. Besides, the radial and spatial distribution function, Gaussian normalized distribution of measured dipole moment of species, and the orientation of dipole moment species were investigated to explain the obtained results. Furthermore, dynamical properties such as mean square displacement and corrected diffusion coefficient of species indicated that CO2 molecules diffuse faster than H2S for studied systems except for aqueous MDEA. Likewise, the diffusion coefficient of H2S and CO2 in the considered systems follow the trends Nano-DES < DES-MDEA < DES < MDEA < aqueous MDEA, and DES < DES-MDEA < MDEA < Nano-DES < aqueous MDEA in the liquid phase, respectively.
... hydrogen sulfide (H 2 S) and carbon dioxide (CO 2 )). Gas sweetening is one of the vital purification processes which is applied to rid natural gas of such acidic contaminants which, if let remain in the natural gas, would cause problems with corrosion, environmental, health and safety hazards [1]. The industrially method most used to sweeten natural gas are those using aqueous alkanolamine solutions in absorption/regeneration process. ...
The solubility of acidic components at various temperatures and pressures in ionic liquids (ILs) is one of the decisive property needed for the appraisal of ILs as potential substitutes for alkanolamines in industrial natural gas sweetening processes, therefore its modeling encompasses scientific and commercial interest. To that end, in the present work, an advanced machine learning approach called stochastic gradient boosting (SGB) tree technique is employed in the calculation of hydrogen sulfide (H2S) solubility in 11 different ILs within the (303.15 to 363.15) K temperature and (0.0608 to 2.0168) MPa pressure range as a function of critical temperature, critical pressure and acentric factor of ILs accompanied with operational temperature and pressure. A collection of 465 experimental data points were assembled from the literatures. The statistical parameters including correlation coefficient (R) of 0.999543 and mean relative absolute error (MRAE), 0.022198, of the results form dataset values exhibit the high precision of the applied method. Furthermore, the prediction competence of the SGB model has been compared to two well-known equation of states (EOS) as well as Genetic Expression Programming (GEP) and least squares support vector machine (LSSVM) models. According to the results of comparative studies, it was found that the SGB model is more robust, reliable and efficient than other existing techniques for improved analysis and design of natural gas sweetening process.
Here, we developed a novel absorbent for the sustainable capture of hydrogen sulfide (H2S), which is the blend of 1,8-diazabicyclo[5.4.0]undec-7-ene (DBU)-triethylene glycol dimethyl ether (TEGDME)-H2O. The results indicated that the gravimetric absorption capacity was 0.174 g H2S/g lower phase at 101.3 kPa and 15 °C, which was 1.3 times higher than that of widely used aqueous methyldiethanolamine (MDEA) solution under the similar condition. In addition, TEGDME was mainly remained in the upper phase, while H2S, H2O, and DBU were mainly concentrated in the lower phase. Moreover, only around 44 wt% of the total absorbent needed to be delivered to the stripper for desorption, and other portion of absorbent was directly sent to the absorption tower. Most importantly, this absorption system was demonstrated to be a benign medium for absorbent regeneration through the liquid-phase Claus process. Our results suggested that this DBU-TEGDME-H2O blend was promising for the sustainable H2S capture and conversion in industrial applications.
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