Evaluation of Methods for H2S
Scavenging
Master Degree in Oil and Gas Technology
9th semester – Group K9og-2-F14
Author:
Supervisor:
Ognyan Marinov
Ass. Prof Marco Maschietti
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Title page
Project from:
Aalborg University Esbjerg (AAUE)
Niels Bohrs Vej 8
6700 Esbjerg
Denmark
Study:
Master Degree in Oil and Gas Technology
Semester:
9th semester K9og-2-F14
Semester theme:
Production, Separation and Piping of Oil and Gas
Project title:
Evaluation of Methods for H2S Scavenging
Author:
Ognyan Marinov
Supervisor:
Ass. Prof. Marco Maschietti
Project period:
February 2014 – May 2014
Submission date:
27th May 2014
Front page picture:
http://www.drillingahead.com/page/lamesa-texasoilfield-worker-dies-after-contact-with-h2sgas?xg_source=activity
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Abstract
Natural gas is the most important and popular fossil fuel for the past decades and it will
continue to be in the future. Probably the main reason is that natural gas’ emissions are the
cleanest among all other sources of energy. However, before it is ready for domestic and
industrial use, unwanted non-hydrocarbon components such as Hydrogen Sulfide and Carbon
Dioxide must be removed. These impurities are undesirable due to their toxicity and lack of
heating value. Some of the problems that these acid gases can cause are corrosion and
environment pollution.
The process of their removal is called sweetening and this project focuses on H2S removal
by aqueous amine solution. This removal method is proven to be one of the most efficient
and widely used in the industry. Three amine solvents are considered for this study –
monoethanolamine (MEA), diethanolamine (DEA) and methyldiethanolamine (MDEA).
Furthermore, three different simulations are made with the help of specialized software –
Aspen HYSYS v7.3.
The chosen field for this study is Khurmala-Erbil situated in Iraq. It is an interesting field due
to the fact that there are huge quantities of H2S (5.37%) and CO2 (4.47%), as well as
relatively low concentration of Methane (63.27%) and relatively big amount of Ethane
(13.88%).
Lastly, approximate cost estimation is obtained. It includes the cost of the major items of
equipment that are used in the sweetening simulations with HYSYS, as well as the amines costs and
the used water for the process.
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Acknowledgment
The author would like to express gratitude towards:
To my direct supervisor Marco Maschietti for his professional advices and guidance
throughout the whole project.
To my family and friends, for their patience and support during the writing of this project.
Last but not least, to my colleagues from 8th and 10th semester in Aalborg University, Oil and
Gas Technology, for their help and guidance.
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Contents
ABSTRACT
ACKNOWLEDGMENT
1. INTRODUCTION
8
2. HYDROGEN SULFIDE
9
3. AMINE GAS SWEETENING PROCESS
3.1 INLET SEPARATOR
3.2 ABSORBER
3.3 TREATED GAS KO DRUM AND AMINE FLASH DRUM
3.4 RICH AMINE / LEAN AMINE HEAT EXCHANGER
3.5 REGENERATOR
3.6 REBOILER
3.7 CONDENSER AND REFLUX DRUM
3.8 MECHANICAL FILTER AND ACTIVATED CARBON FILTER
3.9 AMINE SURGE TANK
4. AMINES
4.1 MONOETHANOLAMINE (MEA)
4.1.1 ASPEN HYSYS SWEETENING SIMULATION WITH MEA
4.2 DIETHANOLAMINE (DEA)
4.2.1 ASPEN HYSYS SWEETENING SIMULATION WITH DEA
4.3 METHYLDIETHANOLAMINE (MDEA)
4.3.1 ASPEN HYSYS SWEETENING SIMULATION WITH MDEA
11
11
12
14
14
15
16
16
16
17
18
18
19
26
27
32
32
5. COST ESTIMATION OF THE PROCESS
37
6. CONCLUSION AND DISCUSSION
40
BIBLIOGRAPHY
41
APPENDIX A
44
APPENDIX B
45
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1. Introduction
Natural gas can be considered as one of the most popular and efficient energy sources in the
recent era. A well-known fact about natural gas is that its emission is the cleanest among all
other sources of energy. This specific quality increases drastically the demand of natural gas
each year. Natural gas is a hydrocarbon, therefore is made up of carbon and hydrogen.
Methane (CH4) is the simplest hydrocarbon; it is also the main component of natural gas.
There are other hydrocarbons in smaller quantities, such as ethane, propane, butane and
others. In addition to the above mentioned, the chemical composition of natural gas includes
some impurities like H2S, CO2, H2O, N2 and others. Table 1 represents an overview of a
natural gas composition, which varies depending on where in the world the gas is produced.
[1]
Table 1 Constituents of natural gas [2]
Name
Formula
Vol. %
Methane
CH4
> 85
Ethane
C2H6
3-8
Propane
C3H8
1-5
Butane
C4H10
1-2
Pentane
C5H12
1-5
Carbon Dioxide
CO2
1-2
Hydrogen Sulfide
H2S
1-2
Nitrogen
N2
1-5
Helium
He
< 0.5
After natural gas is separated from crude oil or produced from natural gas reservoirs, it still
contains all of the previously mentioned impurities. These impurities are harmful for the
equipment and must be removed in order to meet the pipeline specifications. Only then, it is
safe for natural gas to be transported from offshore rig to onshore. H2S and CO2 are
recognized as acid gases and the process of their removal is known as gas sweetening. [3]
Because of the corrosiveness of Hydrogen Sulfide and Carbon Dioxide in the presence of
water, as well as the toxicity of H2S and the lack of heating value of CO2, the gas to the end
consumer must be sweetened. The gas sales contracts limits the concentration of Hydrogen
Sulfide to 4 ppmv (equal to approximately 5.6 mg/m3 of gas), where the content of Carbon
Dioxide should not exceed 2% in natural gas stream. However, these limitations can vary
depending on the countries’ own regulations and allowable quantities of H2S and CO2 in the
natural gas. [4]
There are numerous processes that have been developed for removing acid gases from the
other natural gas components. However, amine gas sweetening is considered as one of the
most popular processes among the other natural gas sweetening methods. Therefore, this
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method will be used for this project with the three most common of them –
monoethanolamine (MEA), diethanolamine (DEA) and methyldiethanolamine (MDEA). [1]
[3] [5]
The focus of this project is set on removing H2S from the gas stream, as well as considering
the advantages and disadvantages of the selected method. In order to receive a more precise
estimation, there are other factors like efficiencies, running costs and materials used in the
construction of the gas cleaning installation, which are taken into account.
2. Hydrogen Sulfide
Hydrogen Sulfide (H2S), also known as sour gas, is colorless and toxic gas even at extremely
low concentrations. The sour gas is highly flammable and can cause combustion if the
concentration in the air is from 4.3 to 46 volume percent. The vapours from H2S are heavier
than air and therefore may migrate substantial distances to a source of ignition. In low
concentrations it also smells like “rotten eggs” and can cause irritation of the eyes, nose and
throat, while in high concentrations is odorless and lethal.
There are several reasons for removing Hydrogen Sulfide from natural gas. First of all, it is
very toxic and when natural gas is used for domestic purposes can be very harmful. The
human nose has evolved to such levels that it can detect concentrations of H2S as low as 0.02
ppmv. Table 2.1 shows how the human body reacts to the different levels of concentration of
Hydrogen Sulfide.
Table 2.1 Hydrogen Sulfide toxicity data [6]
Concentration, ppmv
Effects
0.02
Clearly detectable smell
10
Threshold limit value for prolonged exposure
10 – 100
500 – 700
Slight symptoms after several hours exposure
Maximum concentration that can be inhaled for one
hour without serious effects such as eye and
respiratory irritation
Dangerous after exposure of 30 minutes to one hour
700 – 900
Fatal in less than 30 minutes
1000 and above
Death in minutes
200 – 300
Secondly, because of H2S characteristics and other sulphur compounds it may cause
problems further downstream while processing of oil/gas, if not removed. Problems such as
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corrosion can occur, because of the corrosive abilities of the Hydrogen Sulfide. The corrosion
can cause equipment failure, which will result in stopping of the production process and loss
of money. The H2S acts differently depending on its own content and the content of water in
the gas stream. If natural gas is dry (without water in it) the corrosive abilities of the acid gas
are very low, while in wet natural gas it becomes very corrosive agent. For example, H2S
forms weak acid when it dissociates in water. This weak acid attacks the iron and forms
insoluble iron sulfide. The iron sulfide will adhere to the base metal and may provide some
protection from further corrosion, but if there is high velocity this film can easily be eroded
away, which will expose the fresh metal underneath to further attacks.
To summarize, there are many reasons for removing Hydrogen Sulfide from natural gas
therefore various techniques have been developed for that purpose.
[7] [8] [9]
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3. Amine Gas Sweetening Process
Aqueous alkanolamine processes have first been applied on a gas treating facility in 1930 by
Bottoms (U.S. Patent 1,783,901). Since then, alkanolamines have become the most widely
used solvents for sweeting natural gas streams. The first one that was used commercially for
gas treating was Triethanolamine (TEA). Nonetheless, with time TEA has been displaced by
other more efficient alkanolamines such as Monoethanolamine (MEA), Diethanolamine
(DEA) and in the recent years Methyldiethanolamine (MDEA). This project is based on the
use of these amines for the removal of acid gases from natural gas and a detailed explanation
is provided for each one, together with an Aspen HYSYS simulation of the process. [10]
Figure 3.1 shows a typical amine system for removing H2S from the feed stream. In this
chapter all units will be briefly explained – purpose of the unit, characteristics and etc.
Depending on the amine solvent that is selected for the gas sweetening process slight
modifications can appear, as well as optimizations for specific purposes.
Figure 3.1 Typical amine gas sweetening system [11]
3.1 Inlet Separator
The feed gas stream, also called sour gas stream, containing the acid gases H2S and/or CO2
must always pass through an Inlet Separator (also known as knockout drum). The purpose
of the knockout drum is to catch the entrained hydrocarbon liquids, occasional slugs and
entrained solids. The separator must be designed in such a way that it will have a demister
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pad, in order to prevent liquid droplets carry-over to the amine system. Generally, vertical
drums are used as inlet separators. [12] [13]
3.2 Absorber
The primary function of the contactor is to provide an extensive area of liquid surface which
will be in contact with the gas phase in such conditions that will favour mass transfer. After
the gas passes through the inlet separator it enters the Absorber (Contactor) from the bottom,
while the aqueous amine solution or also called lean solution, enters from the top. This way
the gas flows in upward direction through the column where it makes an intimate countercurrent contact with the lean solution. Chemical reaction between the amine and the sour gas
occurs and the amine solution absorbs the Hydrogen Sulfide and the Carbon Dioxide. The
temperature of the treated gas will rise, due to the heat reaction between the acid gas and the
lean solution. Therefore, the chemical reaction is exothermic. The treated gas leaves the
column from the top, while the rich solution (small amount of hydrocarbons with amine and
acid gas) leaves from the bottom of the column. [14] [15] [16]
The gas absorber is usually a tray column. The tray column is consisting of certain number of
trays, the most efficient number of trays used in the industry is between 15 and 20. The trays
are equipped with weirs (in most cases with a high of around 5 – 7.6 cm), which help for
maintaining a certain level of solvent on each tray, Figure 3.2. As mentioned above, the gas
goes in upward direction and passes through the trays, which have openings such as bubble
caps, perforations or valves. When the gas passes through the openings in the trays it
disperses into bubbles, forming froth. Continuing to travel upwards in the vapour space the
entrained amine solution in the gas has time to fall back on the liquid tray, while the gas will
continue through the next tray.
Figure 3.2 Bubble cap trays in an absorber [17]
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To reduce the amine losses, 2 to 5 trays can be installed at the top of the absorber. These
additional trays will function as a water wash section and it will minimize the vaporization
losses of amine. Furthermore, demister pad can be installed near the gas outlet of the
contactor, which will trap the entrained solvent in the sweetened gas. [15] [18]
The accepted operating conditions of the contactor are as follows:


The operating pressure normally matches the feed gas pressure. Hence, minimum
operating pressure of 4-5 bar is required to maintain feasible and operable process.
As far as the process is concerned there aren’t any limitations on high pressure,
except that with higher pressure the thickness of the steel plates that form the
absorber must be increased.
In order to avoid foaming in the contactor due to hydrocarbon condensation (if the
gas is saturated) the amine solution temperature must be maintained with 10 to
15°F (5.55 to 8.33°C). However, there are limitations on high temperature of the
lean amine stream and the feed stream. High temperatures will affect the
sweetening process – it will lower the acid gas pick up by the lean amine solution
and it will increase the water losses. Throughout the extensive experience of the
oil and gas industry a normal temperature range for the feed gas stream should be
between 80 - 120°F (26.7 - 48.9°C), while the temperature of the lean amine
solution – 90 to 130°F (32.2 – 54.4°C).
[19] [14]
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3.3 Treated Gas KO Drum and Amine Flash Drum
On one hand the Treated Gas KO Drum is a unit with the main purpose to reduce the losses
of amine solution. The treated gas that is coming out from the top of the absorber gets into
the drum, where the sweetened gas is dried by removing the entrained amine. Usually a
vertical drum is used and is equipped with demister pad. When the treated gas comes from
the drum it can be routed to the downstream facilities. [20]
On the other hand Amine Flash Drum is a unit which is installed to carry over the rich
amine solution coming out from the bottom of the absorber. Depending on the composition of
the rich stream, 3 phase horizontal or 2 phase vertical separator is used. The drum’s purpose
is to recover the hydrocarbons that may have dissolved or condensed in the amine solution.
To achieve that, the pressure is significantly reduced to a range between 3 and 5 bar. In this
way the lightest of the hydrocarbons flash, while the heavier hydrocarbons remain as liquid
as well as the aqueous amine solution, forming two separate liquid layers. The two layers are
formed, because of the different densities of the solutions. The hydrocarbons have lower
density; therefore they occupy the upper liquid layer and can be skimmed off the top. On the
other side the aqueous amine solutions is freed from the hydrocarbons and can be drained
from the bottom of the drum, Figure 3.3. [20] [21] [18]
Figure 3.3 Amine Flash Drum [22]
3.4 Rich Amine / Lean Amine Heat Exchanger
After the rich amine solution is freed from the dissolved hydrocarbons the next step is the
amine to be regenerated, but before entering the regenerator column the stream must be
preheated. The lean amine, coming out from the reboiler (it is explained further on) is of
higher temperature than the rich amine entering. The lean amine must be cooled therefore
14 | P a g e
there is a very good opportunity to exchange heat between the lean amine and the rich amine
stream. This way the lean stream will be cooled, while the rich stream will be preheated. By
using cross-flow heat exchanger the total artificial heat requirements for the process are
lowered, which leads to lower costs.
The types of heat exchangers that are used for this purpose are usually shell and tube
exchanger and plate and frame. Where, for the shell and tube the rich solution flows through
the tubes and the lean solution through the shell. In order to reduce the corrosion in the tubes
a velocity of 1 m/s is recommended. In the industry, the flow coming from the bottom of the
separator enters the heat exchanger with its current pressure and the temperature range for
each stream is as follows – outlet rich amine (90 – 110°C), inlet lean amine (110 - 130°C).
[23] [20] [18]
3.5 Regenerator
The Regenerator (also called stripper) as well as the absorber is a tray tower and is generally
designed with between 15 to 20 trays. This fractionation column is also equipped with
condenser (using water or air as cooling medium) and a reboiler (using hot oil or steam as
heating medium).
After the rich amine is preheated in the cross-flow heat exchanger, it enters the stripper near
the top and flows down counter current to a gas stream of steam, Hydrogen Sulfide and
Carbon Dioxide. The partial pressure of H2S and CO2 in the gas stream is lowered, due to the
steam generated in the reboiler, which leads to enhancing the driving force of the acid gases
from the amine solution. The liberated H2S and CO2 as well as the remaining steam keep its
way in upward direction until they exit the stripping section. On the other side, the rich amine
keeps making its way down the regenerator column and steadily gives up the absorbed
Hydrogen Sulfide and Carbon Dioxide and becomes leaner until it leaves the regenerator
from the bottom.
At the top of the stripper the vapour leaves and enters the reflux section where it is cooled in
a condenser. The steam is condensed and separated from the H2S and CO2 gas in the reflux
drum and then recycled back as reflux. The acid gas stream is sent for further treatment and
compression in order to meet the end use specifications.
Similarly to the absorber column, in the regenerator also can be installed a demister pad to
reduce the amine loss caused by physical entrainments. Depending on the given case the best
operating pressure for the process can be between 1.2 and 2 bar or between 2.1 and 3 bar.
[24] [20] [25]
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3.6 Reboiler
The objectives of the reboiler are “to heat up the rich solution, to produce the energy to
reverse the chemical reaction, to free acid gas from the amine solution and to generate steam
to strip acid gas from solution. The stripping efficiency is controlled by the overhead reflux
ratio.” [26]
The reboiler is usually a shell and tube exchanger, which uses hot oil or steam as heating
medium. The temperature inside the tube should not exceed 145°C, because above this
temperature will cause amine degradation. [26]
3.7 Condenser and Reflux Drum
The top product of the stripper is a water vapour and acid gas mixture. This mixture is
condensed in the Condenser to generate reflux to the regenerator. It uses ambient available
cooling medium, such as water or air. The type of exchanger used for this purpose is shell
and tube. [26]
After the mixture of water and acid gas is condensed the next step is to be collected and
separated. The Reflux Drum, or also called regenerator overhead drum, separates the
mixture into water and acid gas. The water leaves the drum from the bottom and is handled
by reflux pumps to the top of the regenerator. The acid gas leaves the drum from the top and
is routed to the acid gas disposal system. [26] [25]
The reflux drum is vertical or horizontal and it is equipped with demister pad. The
temperature must be maintained as low as possible in order to minimize the concentrations of
amine and water entrainments in the stream routed to the downstream facilities. Such
facilities could be for example – Sulfur recovery unit. [26] [25]
3.8 Mechanical Filter and Activated Carbon Filter
The filtration of the stream, which maintains good solution control, is important part of the
process. For this purpose Mechanical Filter and Activated Carbon Filter are used. They
remove solid impurities such as sand, iron sulfide, pipeline dust, and iron oxide. In order to
prevent erosion and corrosion, these impurities must be removed from the solution.
The filters can be located either in the rich amine stream or in the lean amine stream. In the
first case, locating the filters before the heat exchanger and the regenerator will protect them
from plugging. It will also reduce the erosion and corrosion rate in the exchanger. The main
problem for installing the filters in the rich line is that the solution is heavily loaded with acid
16 | P a g e
gas and gas pockets can form in the filter resulting in reduced or complete blocked flow, due
to drops in pressure, which will cause the acid gas to flash from the solvent. Taking in
consideration the safety of the workers who will inspect and clean out the filters, in most
plants the filters are located in the lean line.
Regardless of the line in which the filters are located, the mechanical filter (also called main
filter) should be positioned before the activated carbon filter. In this way the activated carbon
filter is protected from removing both particle matter and chemical contaminants, which may
lead to plug up with solid material long before the chemical capacity of the filter is
exhausted. However, this is a costly way to operate due to the need of frequent changeouts of
the filter.
[26] [25]
3.9 Amine Surge Tank
The final stage, before the lean amine is cooled and recycled to the absorber, is to pass
through an Amine Surge Tank. One of the tank’s purposes is that in event of process
shutdown it contains the entire contents of the circulating system. Moreover, it also maintains
the strength of the lean amine solution by periodic make-ups. In this way any minor or major
amine losses due to thermal degradation will be covered, as well as water losses during the
absorption and the regeneration stages.
The characteristics of the tank are as follows – atmospheric classical storage tank with fixed
or cone roof; the operating temperature should be above the freezing point of the solution; the
tank must be blanketed with nitrogen or treated gas in order to avoid amine degradation due
to contact with air.
[27] [28]
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4. Amines
Amines are compounds formed from ammonia (NH3). They may be categorized by the
number of hydrogen atoms bonded to the central nitrogen atom. For example, if one
hydrogen atom is replaced with another hydrocarbon group, the result is a primary amine. If
two hydrogen atoms are replaced the result is secondary amine and respectively, if three are
replaced tertiary amines. The primary amines are the one that form the strongest bases,
followed by the secondary and lastly the tertiary amines. Amines with stronger base form
stronger chemical bonds and are more reactive to acid gases (H2S and CO2).
The amines are used in aqueous solutions and their concentration ranges from 10 to 65 wt%
(approximately). The physical properties of the amines used in this project can be seen in
APPENDIX A. Most amines are alkanolamines, the difference between them is that the
alkanolamines are with OH groups attached to the hydrocarbon groups and this way they
reduce their volatility.
The process for removing acid gases from the sour gas can be introduced in two steps:
1. Physical absorption – the gas is dissolved in the liquid;
2. The weakly basic amines react with the weak acid (the dissolved gas).
[30] [31]
4.1 Monoethanolamine (MEA)
Monoethanolamine (abbreviated as MEA) is a primary amine, which means (as mentioned)
that one of the hydrocarbons bonded to the nitrogen is replaced. MEA is the most reactive
amine for acid gas removal and is also the most basic one. It is toxic, flammable, colourless,
corrosive and viscous liquid. The molecular structure of MEA can be seen in the following
figure [32]:
Figure 4.1 Molecular structure of MEA [30]
For many years the aqueous monoethanolamine solution has been used almost exclusively for
the removal of hydrogen sulfide and carbon dioxide from natural gas. It has been used for all
kinds of pressures (low, moderate and high), but with years it has been replaced by other
18 | P a g e
more efficient amines for threatening natural gases with high pressure. However, MEA is still
a very good aqueous solution for removing acid gases from natural gas with low to moderate
pressure. Besides the pressure criteria, when choosing a solvent another criteria, such as the
concentration of H2S and CO2 in the treated gas, must be reviewed. For low to moderate
concentrations of H2S and CO2, MEA is capable of reducing it to level as low as 4 ppmv (for
hydrogen sulfide) and 100 ppmv (for carbon dioxide). The aqueous monoethanolamine
solution is removing both acid gases, but cannot be used for selective removal of H2S or CO2,
when both of them are present in the threated gas. The concentration of MEA in the water
solution is in the range between 10 to 20 wt%. Hence, for so many years the oil and gas
industry uses the monoethanolamine, proving that 15 wt% is the most commonly used
concentration. [33] [34] [27]
One of the advantages of MEA is its low molecular weight, which leads to high solution
capacity at moderate concentrations, which makes it very useful for the complete removal of
H2S and CO2 with low to moderate concentration in the threated gas. Other advantages are its
high alkalinity and the ease with which it can reclaim from contaminated solutions. However,
there are numerous disadvantages, such as the formation of irreversible reaction products
with COS (carbonyl sulfide) and CS2 (carbon disulfide). This leads to big chemical losses if
the threated gas contains excessive amounts of these compounds. Another disadvantage is the
relatively high vapour pressure, which leads to high vapourization losses. Moreover, the
inability for selective removal of H2S or CO2, when the other compound is present; higher
corrosion rate than most other amines; and others. [32] [34]
4.1.1 Aspen HYSYS Sweetening simulation with MEA
In this subpoint a typical acid gas treating facility is simulated in Aspen HYSYS (Version
7.3). The units that are present in this simulation have already been explained in Chapter 3.
Therefore, only explanations for the chosen operating conditions, design and results will be
provided in this subpoint.
Aspen HYSYS v.7.3 is a special and valuable process simulator, which is recognized as one
of the leading platforms for simulation and optimization in the oil and gas industry. The
software provides engineers with flexible access to wide variety of tools and equipment, as
well as the opportunity to achieve better design, investment decisions and to increase the
efficiency of upstream projects. It also gives you the ability to choose between different fluid
packages, regarding the needs of the current case. Choosing the right package leads to getting
a better outcome which is closer to more realistic result. Like in the case of this project the
use of the right package is very essential for obtaining correct results. Aspen HYSYS offers
two fluid packages that can be used for amine sweetening process – amine package (KentEisenberg or Li-Mather) and DBR amine package (Kent-Eisenberg, Li-Mather or Physical
Solvent).
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In this project the DBR amine package Li-Mather is used. The reason for choosing this
package, instead of the amine package, is that it is newer, more improved thermodynamic
model and the predictions are based on newly available experimental data. Furthermore, there
are no limitations for alkanolamine concentration, acid gas partial pressure, operating
temperature (25 to 125°C) and the components used in the simulation. [35]
The field and its specifications shown in Table 4.1 are not chosen by chance for the
implementation of this project. There were some requirements for selecting the right study
case. First of all, it was very important that the composition will be from a real field, because
of the fact that the project will not be based on a typical imaginary composition. Secondly,
the requirement was for a non-typical field. Taking a look at the composition it can be
observed that the mole % of Ethane in the composition is relatively high. Also, the Methane
is far from its normal range, between 70 and 90 % (in some exceptional cases even more).
Moreover, in the natural gas of Khurmala-Erbil’s field there is a very high quantity of H2S
and CO2, respectively 5.37% and 4.47%. Everything mentioned about the chosen study case
so far, will make the project more interesting and unique, because of the fact that a nontypical design specifications will be chosen in order to obtain the needed results.
Table 4.1 Design data [36]
Location
Field
Sample no.
Test data
Sample type
Flow rate
Gas density
Gas SG (Air = 1.0)
Pressure
Temperature
Erbil, Iraq
Khurmala-Erbil
1 Cylinder (S-012)
14.03.2010
Natural Gas
337.6 m3/h
0.65 Kg/m3
0.67
70 bar
38 °C
Component
CH4
C2H6
C3H8
i- C4H10
n- C4H10
i- C5H12
n- C5H12
C6H14
N2
H2S
CO2
H2O
Mole %
63.27
13.88
6.02
1.36
2.44
1.03
0.73
1.19
0.11
5.37
4.47
0.13
The composition and specifications of the natural gas given in the above table are the one
designing the Feed Stream in the simulation. The first step of the simulation is the Feed
Stream to pass through an Inlet Separator, where the entrained hydrocarbon liquids and
solids will be removed. Furthermore, if there are slugs in the stream, when the stream hits the
inlet diverter it will prevent them.
After the feed stream is separated in the vertical separator, the product that comes from the
top of the separator is going to the bottom of the Absorber, where it will be treated for
removing the acid gases. The design of the Contactor is as follows:
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


The column is with 16 trays, as this was found to be the optimum number of
trays based on removal efficiency and cost. Moreover, as mentioned
previously the plants absorber units are usually designed with 15 to 20 trays.
Operating pressure of the contactor is chosen to be the same as the sour gas
coming from the reservoir – 70 bar. It is known that with high pressure and
low temperature, the sweetening process is more efficient.
The pressure drop from the first to the last tray is considered to be in the range
between 0.1 – 0.2 psi/tray. In this project is taken 0.2 psi/tray, which is 3 psi
(0.2 bar) pressure drop.
After the Absorber is designed the Lean MEA must enter from the top of the column. The
monoethanolamine stream design is as follows:



The temperature is with 6°C hotter than the sour gas, 44°C respectively. When
the lean MEA stream is hotter it avoids foaming and maximizes the fuel gas
cleanup.
The pressure is taken to be the same as the feed gas stream, 70 bar.
To design a correct flow value of the stream, the concentration of MEA in the
aqueous solution must be selected. The concentration of MEA should be in
range of 10 to 20 wt%. For this case 10 wt% are chosen and now the liquid
flow can be obtained from the formula – Vol. Flow = 41*(Q*y/x). Where Q is
the molar flow of the sour gas to be processed (107.5 MMscfd); y is the acid
gas concentration in the sour gas (H2S = 5.38%; CO2 = 4.48%; Total =
9.86%); and finally x is the chosen amine solution in liquid solution (10 wt%).
When all this data is known, from the formula above follows that Vol. Flow =
4346 USGPM (987 m3/h). However, with this flow for the aqueous amine
solution the removal of H2S was 100%. That means that the process is very
efficient, but not and from financial point of view. The maximum
concentration of H2S in the sweet gas, should not exceed 0.0004%, hence a
new volume flow was considered until the maximum specifications are met.
With trial and error method was found that flow of 328.8 m3/h is enough to
remove the Hydrogen Sulfide to a level of 0.0004%. As it can be observed the
optimum flow is 3 times lower than the one found for the complete removal of
the acid gas. [33]
21 | P a g e
When the Contactor, Lean MEA stream and Feed stream (Figure 4.2) are designed, the
results from the process can be obtained.
Figure 4.2 Design of the sweetening process in Aspen HYSYS
From Table 4.2 the composition of the sweet gas can be observed. As seen the acid gas,
Hydrogen Sulfide, is removed to the maximum allowed concentration and the Methane is
now 72% of the total composition and there is a very small amount of water added. From
these results it can be concluded that the amount of MEA in the aqueous solution is enough
and the calculated flow is correct. Furthermore, after the acid gas is removed from the gas
stream, the Sweetened Gas stream can be transported to upstream facilities for further
treatments.
Table 4.2 Mole fraction, molar flow and mass flow of Sweetened Gas Stream
COMPONENTS
Methane
Ethane
Propane
i-Butane
n-Butane
i-Pentane
n-Pentane
n-Hexane
Nitrogen
H2S
CO2
H2O
MEAmine
Total
MOLE FRACTION
MOLAR FLOW
(kgmole/h)
MASS FLOW
(kg/h)
0.7232
0.1531
0.0627
0.0132
0.0228
0.0085
0.0057
0.0071
0.0013
0.0004
0.0000
0.0021
0.0000
1.0000
3 301.7
699.2
286.3
60.3
104.2
38.7
26
32.2
5.8
1.7
0.1
9.5
0.04
4 565.7
52 969.2
21 023.3
12 624.9
3 506.8
6 054.3
2 790.5
1 874.7
2 778.5
162.7
56.9
4.8
170.7
2.5
104 019.6
22 | P a g e
From Table 4.3 can be observed that the product that comes from the bottom of the Absorber
is with composition mainly of water, MEA, CO2 and H2S, as well as very small amounts of
Methane and Ethane. Their composition in the rich stream is so low that it will be costly and
inefficient to design a separator for the separation of these light hydrocarbons. Therefore, the
rich amine stream is headed directly for preheating in the Cross-flow Heat Exchanger,
which follows as a next stage of the process.
Table 4.3 Mole fraction, molar flow and mass flow of Rich Gas Stream
COMPONENTS
Methane
Ethane
Propane
i-Butane
n-Butane
i-Pentane
n-Pentane
n-Hexane
Nitrogen
H2S
CO2
H2O
MEAmine
Total
MOLE FRACTION
MOLAR FLOW
(kgmole/h)
MASS FLOW
(kg/h)
0.0007
0.0001
0.0000
0.0000
0.0000
0.0000
0.0000
0.0000
0.0000
0.0155
0.0140
0.9371
0.0325
1.0000
12
1.8
0.5
0.0
0.0
0.0
0.0
0.03
0.01
268.9
243
16 233
563.2
17 322.2
191.4
53.5
20.6
0.4
0.7
0.4
0.2
2.4
0.4
9 162.6
10 690.4
292 439.2
34 399.2
346 961.3
After exiting the absorber from the bottom the RICH MEA stream is with very high pressure
(70 bar), in order to continue to the regeneration process the pressure must be reduced.
Hence, the stream passes through a valve so that the pressure of the stream is reduced to 3 bar
and enters the Cross Heat Exchanger.
The liquid stream (rich monoethanolamine stream) must be preheated before it enters the
regenerator column. Therefore, cross heat exchanger is designed, which will use the heat
from the lean amine stream coming from the bottom of the stripper.
The RICH MEA stream will be inside the tubes, while the Lean Amine stream will be in the
shell. The rich amine solution is preheated from 65.38°C to 98°C with a pressure drop of 1
bar. With these final changes, the rich monoethanolamine is ready to be regenerated.
The design of the Regenerator is as follows (Figure 4.3):

In the same way as with the absorber column, generally 15 to 20 trays are used
in the industry, for this case 20 trays are considered. There is a little difference
between the regenerator and the absorber tough, in the Stripper there are 2
additional stages which are not included in these 20 trays, Condenser and
Reboiler, therefore the final number of trays is designed to be 18.
23 | P a g e



The RICH MEA stream enters the column at 15th tray.
The Condenser is designed to be with full reflux, pressure of 1 bar, pressure
drop of 0.2 bar and temperature of 54.4°C.
The Reboiler is designed with pressure of 2 bar and a duty of 8.3e8 BTU/hr.
Figure 4.3 Cross heat exchanger and Regeneration column in HYSYS
When the rich amine solution is regenerated the product that comes from the top of the
column is composed mainly from H2S, CO2, H2O and small amounts of Methane, Ethane and
Propane. On the other hand, the product from the bottom of the stripper is lean aqueous
monoethanolamine solution (Table 4.5).
24 | P a g e
The top stream is sent for further treatment and compression, in order to meet the end use
specifications, while the bottom stream flows through the cross heat exchanger, to transfer its
temperature to the rich amine solution, and continuous its way to be recycled in the absorber
column.
Table 4.5 Conditions and mass flows of the products from the Regenerator
CONDITIONS
COND VAP
(top of the Reg.)
REG MEA
(bottom of the Reg.)
Pressure (bar)
1
2
Temperature (°C)
55
120.7
MOLE COMPOSITION
(%)
COMPONENTS
Methane
Ethane
Propane
i-Butane
n-Butane
i-Pentane
n-Pentane
n-Hexane
Nitrogen
H2S
CO2
H2O
MEAmine
Total
0.0196
0.0029
0.0008
0.0000
0.0000
0.0000
0.0000
0.0000
0.0000
0.4426
0.3793
0.1546
0.0000
1.0000
0.0000
0.0000
0.0000
0.0000
0.0000
0.0000
0.0000
0.0000
0.0000
0.0000
0.0008
0.9655
0.0337
1.0000
Knowing the mass flow of the initially used lean MEA stream and after that the mass flow of
the regenerated MEA stream, the total loss can be estimated (Table 4.6).
Table 4.6 Estimating the process losses
LEAN MEA
COMPONENTS
H2O
MEAmine
REGENERATED
MEA
MASS FLOWS
(kg/h)
292 242.1
34 401.7
290 747.8
34 399.2
LOSS
(kg/h)
1 494.3 (0.51%)
2.5 (0.007%)
25 | P a g e
From the table above can be observed that the losses of water are very low but still must be
fulfilled, while the MEA losses are very insignificant. After the REG MEA stream passes
through the cross heat exchanger it lowers its temperature from 120.7 to 77.4°C. A mixer is
installed after the exchanger with water stream makeup. This makeup stream will fulfil the
water losses so that it can reach the primary specifications. The design of the water stream is
– pressure 1 bar, temperature 21°C and the mass flow 1 494.3 kg/h.
Now, when the composition specifications are met the stream must be cooled to 44°C. This
can be achieved with the help of a cooler. After that a pump must be installed to rise up the
pressure of the regenerated aqueous amine solution to 70 bar.
The final design that should be done is to install a Recycle Unit, which will add up together
the two streams – Lean MEA and REG MEA. In this way the sweetening process is made as
a closed system, where the lean monoethanolamine is used and after that regenerated and
recycled.
In conclusion it can be said that removing the acid gas, H2S, can be efficiently removed with
aqueous monoethanolamine solution (MEA concentration 10 wt%). There is a very small and
insufficient loss of MEA during the whole process and a very low loss of water (0.51%). The
water is added with a makeup stream to the regenerated MEA stream and is recycled to the
Contactor column.
4.2 Diethanolamine (DEA)
Diethanolamine, or also abbreviated as DEA, is a secondary amine. The chemical formula
can be seen in Figure 4.4. Compared with MEA is less basic, reactive and it has lower vapour
pressure, therefore DEA has lower evaporation losses. Another comparison that can be made
with MEA is that DEA is also non selective and cannot separate only H2S or CO2 when both
of them are present in the stream. Moreover, diethanolamine is widely used for treating gases
with COS and CS2 compounds, because it forms regenerable compound with them. [33] [37]
[32]
Figure 4.4 Diethanolamine’s chemical formula [30]
DEA is also used in aqueous solution, but the concentration ranges from 25 to 35 wt%. The
most commonly used concentration in the practice is 25 wt%. With DEA the level of removal
can reach concentration as low as 4 ppmv for H2S and 100 ppmv for CO2. [37]
26 | P a g e
Some of the advantages of this amine compared to MEA are that:


Can be used for partial removal of COS and CS2, without significant solution
losses.
Because it is chemically weaker (compared to MEA) it requires less heat to
strip.
4.2.1 Aspen HYSYS Sweetening simulation with DEA
The design and the simulation are more or less the same as the sweetening with MEA.
Therefore, the design will be very briefly explained and the results will be obtained. The
composition given in Table 4.1 is used for this simulation as well. The chosen fluid package
is again DBR amine package Li-Mather.
The Feed Stream passes through an Inlet Separator and from there the top product of the
separator is forward to the bottom of the Contactor column where it will be sweetened. The
design of the Absorber is with pressure of 70 bar, pressure drop of 0.2 bar and number of
trays 20.
Figure 4.5 Design of the sweetening process with DEA in Aspen HYSYS
The Lean DEA entering the Absorber from the top is with 6°C hotter than the sour gas,
44°C respectively, pressure of 70 bar and Vol. Flow = 32*(Q*y/x) = 1357 USGPM (308.3
m3/h). With the calculated flow and the chosen number of trays (20), the removal of H2S
from the stream was 100%. Therefore, the design was revised and the number of trays was
lowered to 15. With this design change the concentration of Hydrogen Sulfide was 0.0005%.
The obtained result was just a little bit higher than the needed one (0.0004%), hence the flow
was lowered to 302.1 m3/h. With these final changes the optimum operating conditions were
found. [33]
27 | P a g e
In Table 4.7 the composition of the sweet gas is shown.
Table 4.7 Mole fraction, molar flow and mass flow of Sweet Gas Stream
COMPONENTS
Methane
Ethane
Propane
i-Butane
n-Butane
i-Pentane
n-Pentane
n-Hexane
Nitrogen
H2S
CO2
H2O
DEAmine
Total
MOLE FRACTION
MOLAR FLOW
(kgmole/h)
MASS FLOW
(kg/h)
0.7231
0.1531
0.0627
0.0132
0.0228
0.0085
0.0057
0.0071
0.0013
0.0004
0.0002
0.0020
0.0000
1.0000
3304
699.5
286.4
60.3
104.2
38.7
26
32.2
5.8
1.7
1.1
9.3
0.0
4 569.2
53 005.1
21 033.3
12 628.8
3 506.8
6 054.5
2 790.6
1 874.7
2 778.8
162.7
58.5
46.9
168.1
0.0
104 108.9
As seen from the table above, the hydrogen oxide equals the maximum allowed concentration
in the natural gas. In this way the specifications are met, as well as the cost efficiency of the
process. With this composition the sweet gas can be transported to the upstream facilities for
further treatment.
28 | P a g e
The composition of the stream, also called rich amine stream, coming out from the bottom of
the Contactor, can be seen in Table 4.8.
Table 4.8 Mole fraction, molar flow and mass flow of Rich Gas Stream
COMPONENTS
Methane
Ethane
Propane
i-Butane
n-Butane
i-Pentane
n-Pentane
n-Hexane
Nitrogen
H2S
CO2
H2O
DEAmine
Total
MOLE FRACTION
MOLAR FLOW
(kgmole/h)
MASS FLOW
(kg/h)
0.0007
0.0001
0.0000
0.0000
0.0000
0.0000
0.0000
0.0000
0.0000
0.0195
0.0174
0.9092
0.0530
1.0000
9.7
1.5
0.4
0.0
0.0
0.0
0.0
0.0
0.0
273.8
244.4
12 746.3
743.3
14 019.3
155.7
43.5
16.8
0.3
0.6
0.3
0.2
2
0.3
9 329.7
10 754.6
229 626.0
78 147.7
328 077.6
As it can be observed from the table above, the concentration of Methane and Ethane in the
stream is very insignificant. Thus, the use of Flash Drum won’t be needed, because it will
increase the cost value and the results won’t be satisfying.
Before the rich amine solution is directed to the next stage, the stream must be depressurised
from 70 bar to 3 bar. The lowering of the pressure is achieved with the help of a valve. After
the rich solution passes through the valve, it is directed to be preheated to 98°C in a Cross
Heat Exchanger which uses the heat from the regenerated amine that comes from the
Regenerator.
From the Cross Heat Exchanger the stream enters the Regenerator column near the top.
The column is designed to be with 18 trays, condenser and reboiler. When it enters, the rich
amine solution flows in downward direction until is cleaned and exits the bottom of the
stripper. The Condenser is designed to be with full reflux, pressure of 1 bar, pressure drop of
0.2 bar and temperature of 54.4°C. The Reboiler on the other hand is with pressure of 2 bar
and a duty of 7.9e7 BTU/hr.
29 | P a g e
The products from the Regenerator column can be seen in Table 4.10 where the conditions
and the mass flows of the two streams can be observed. The regenerated diethanolamine
stream is consisted almost entirely of water and DEA. In the vapour stream it can be seen
some small amounts of methane, ethane, but is mostly consisted of H2S, CO2 and H2O. The
water in the vapour stream can be considered as a loss during the process.
Table 4.10 Conditions and mass flows of the products from the Regenerator
CONDITIONS
COND VAP
(top of the Reg.)
REG DEA
(bottom of the Reg.)
Pressure (bar)
1
2
Temperature (°C)
56.25
121.5
MOLE COMPOSITION
(%)
COMPONENTS
Methane
Ethane
Propane
i-Butane
n-Butane
i-Pentane
n-Pentane
n-Hexane
Nitrogen
H2S
CO2
H2O
DEAmine
Total
0.0161
0.0024
0.0006
0.0000
0.0000
0.0000
0.0000
0.0000
0.0000
0.4457
0.3805
0.1547
0.0000
1.0000
0.0000
0.0000
0.0000
0.0000
0.0000
0.0000
0.0000
0.0000
0.0000
0.0004
0.0011
0.9431
0.0554
1.0000
Knowing the mass flow of the primary used lean MEA and after that the mass flow of the
regenerated DEA, the total loss can be estimated (Table 4.11).
Table 4.11 Estimating the process losses
LEAN DEA
COMPONENTS
H2O
DEAmine
REGENERATED
DEA
MASS FLOWS
(kg/h)
229 767.4
78 147.8
227 946.6
78 147.7
LOSS
(kg/h)
1 729.8 (0.75 %)
0.1 (0.00013 %)
30 | P a g e
From the table above can be concluded that there is 3/4 per cent loss of water during the
sweetening and regeneration process. This is with ¼ more the loss of MEA in the previous
case. Although, the losses of DEA are very insignificant and do not need much attention.
The loss of water is compensated by a makeup stream. This stream of water is installed right
after the heat exchanger. The design of this stream is – pressure 1 bar, temperature 21°C and
the mass flow 1729.8 kg/h. The makeup stream and the regenerated DEA are mixed in a
mixer and continue to be cooled to 44°C and after that to increase the pressure with a pump
to 70 bar. The last step of the design of this gas sweetening facility is to install a recycle unit,
which will help with the issue of closing the loop so that it can be a closed system.
In conclusion, the sweetening with DEA aqueous solution is a good method due to the
achieved low concentration of H2S and the very low DEA loss. The losses of water are a little
higher than the MEA case, but are still under one per cent, which is considered as very good
and efficient result.
31 | P a g e
4.3 Methyldiethanolamine (MDEA)
MDEA which stands for methyldiethanolamine is a tertiary amine. The most specific thing
about this amine, which is also the reason why it became so popular in recent years, is that it
can selectively remove H2S to specifications as low as 0.0004% in the presence of CO2. This
result is possible due to the fact that MDEA is “slipping” some of the Carbon Dioxide. “The
CO2 slippage occurs because H2S hydrolysis is much faster than that for the CO2, and the
carbamate formation reaction does not occur with a tertiary amine”. [38] As a result from this
slippage, to obtain the already mentioned selectivity a short contact time is used in the absorber.
One of the advantages of this tertiary amine is that it has low vapour pressure and hence, can
be used at concentrations as high as 60 wt% without significant vapourization losses. The
methyldiethanolamine is most commonly used in concentrations between 35 and 50 wt%.
These high concentrations of MDEA result in reduced circulation flow rates. On the other
hand, the reduced circulation flow rate will result in significant capital savings due to reduced
pump and regeneration requirements. Moreover, methyldiethanolamine also has a lower heat
requirement due to its low heat of regeneration. All these advantages of this tertiary amine
can be best used for treating gas streams with moderate to high pressure.
The molecular structure of the methyldiethanolamine can be seen in the following figure.
Figure 4.6 Molecular structure of MDEA [30]
[33] [39] [38] [40]
4.3.1 Aspen HYSYS Sweetening simulation with MDEA
The sweetening simulation with aqueous MDEA solution is using the same units and most of
the design data, as well as the same fluid package (DBR Amine Package). Therefore, only the
most important stages will be covered in this chapter as well as the obtained results.
The composition of the Feed Stream is taken from Table 4.1, as well as the operating
conditions. The sour gas enters the bottom of the Absorber, which is designed with 15 trays
(found to be the optimum number of stages), pressure of 70 bar and a pressure drop of 0.2
bar.
32 | P a g e
The Lean MDEA stream enters the top of the contactor column and is designed to be with
6°C hotter than the feed stream. The pressure of the stream is designed to meet the pressure
of the sour gas – 70 bar. Methyldiethanolamine is with 35 wt% in the aqueous solution and
the needed flow rate is calculated from the following formula [41]:
𝐺𝑃𝑀 =
0.206 ∗ 𝑀𝑀 ∗ (𝐻2 𝑆 + 𝐶𝑂2 ) ∗ 𝑀𝑊𝑇
𝑀𝐿 ∗ 𝑊𝑇
Where,
GPM – gallons per million;
0.206 – constant;
MM – gas flow, MMSCFD = 107.5;
H2S – mol% of H2S to be removed = 5.37%;
CO2 – mol% of CO2 to be removed = 4.47%
MWT – mole WT of MDEA = 119.9;
ML – mol loading, moles acid gas/mole = 0.6;
WT – amine solution weight percent circulated = 35 wt%.
When the given numbers are replaced in the formula, the calculated flow is 1246.7 GPM =
284.2 m3/h (Figure 4.7).
Figure 4.7 Design of the sweetening process with MDEA in Aspen HYSYS
The composition of the treated gas coming from the top of the Absorber is given in Table
4.12. As it can be observed the H2S concentration in the sweet gas is meeting the maximum
allowed pipeline specifications – 0.0004%. The composition of the other components is very
similar to the obtained results in the previous two cases, except for the CO2. As mentioned
above MDEA can be used for selective removal and the designed simulation and the given
results are confirming that.
33 | P a g e
Table 4.12 Mole fraction, molar flow and mass flow of Sweet Gas Stream
COMPONENTS
Methane
Ethane
Propane
i-Butane
n-Butane
i-Pentane
n-Pentane
n-Hexane
Nitrogen
H2S
CO2
H2O
MDEAmine
Total
MOLE FRACTION
MOLAR FLOW
(kgmole/h)
MASS FLOW
(kg/h)
0.7132
0.1510
0.0618
0.0130
0.0225
0.0083
0.0056
0.0070
0.0013
0.0004
0.0139
0.0021
0.0000
1.0000
3305
699.6
286.4
60.3
104.2
38.7
26
32.3
5.8
1.6
64.6
9.5
0.0
4 634
53 021
21 038.2
12 630.7
3 506.9
6 054.6
2 790.7
1 874.7
2 779.5
162.8
55.7
2 840.7
171.5
1
106 928.2
In Table 4.13 the composition of the rich amine solution coming from the bottom of the
contactor, can be seen. The stream is consisted mainly from water, methyldiethanolamine,
hydrogen sulfide and carbon dioxide. In the stream there is just a minor amount of light
hydrocarbons – methane and ethane, therefore design and installation of Flash Drum won’t
be necessary.
Table 4.13 Mole fraction, molar flow and mass flow of Rich Gas Stream
COMPONENTS
Methane
Ethane
Propane
i-Butane
n-Butane
i-Pentane
n-Pentane
n-Hexane
Nitrogen
H2S
CO2
H2O
MDEAmine
Total
MOLE FRACTION
MOLAR FLOW
(kgmole/h)
MASS FLOW
(kg/h)
0.0007
0.0001
0.0000
0.0000
0.0000
0.0000
0.0000
0.0000
0.0000
0.0237
0.0139
0.8934
0.0681
1.0000
8.7
1.3
0.3
0.0
0.0
0.0
0.0
0.0
0.0
283
166.1
10 662.3
813.1
11 934.9
139.6
38.6
14.8
0.2
0.4
0.2
0.1
1.3
0.3
9 644.1
7 307.9
192 082.7
96 892.5
306 122.7
34 | P a g e
Before the rich methyldiethanolamine enters the Regenerator column, the stream is
depressurized to 3 bar and is preheated to 98°C from the regenerated amine coming from the
Reboiler.
The Regenerator column design consists of 18 trays, where the Condenser (at the top) and
the Reboiler (at the bottom) are considered as 2 stages (trays). The operating pressure of the
condenser is 1 bar with a temperature of 54.4°C and a full reflux to the column. The reboiler
is designed to be with operating pressure of 2 bar and duty of 7.3e7 Btu/hr.
The products from the Regenerator column can be seen in Table 4.14 where the conditions
and the mass flows of the two streams can be observed. The regenerated
methyldiethanolamine solution stream is consisted almost entirely of water and MDEA. In
the vapour stream some small amounts of methane and ethane can be seen, but it is mostly
consisted of H2S, CO2 and H2O. The water in the vapour stream can be considered as a loss
during the process.
Table 4.14 Conditions and mass flows of the products from the Regenerator
CONDITIONS
COND VAP
(top of the Reg.)
REG MDEA
(bottom of the Reg.)
Pressure (bar)
1
2
Temperature (°C)
56.17
122
COMPONENTS
Methane
Ethane
Propane
i-Butane
n-Butane
i-Pentane
n-Pentane
n-Hexane
Nitrogen
H2S
CO2
H2O
MDEAmine
Total
MOLE COMPOSITION
(%)
0.0165
0.0024
0.0006
0.0000
0.0000
0.0000
0.0000
0.0000
0.0000
0.5106
0.3151
0.1546
0.0000
1.0000
0.0000
0.0000
0.0000
0.0000
0.0000
0.0000
0.0000
0.0000
0.0000
0.0013
0.0000
0.9275
0.0713
1.0000
35 | P a g e
When the regeneration stage is completed, all necessary data is available for calculating the
losses of the process. The mass flow of the preliminary stream (Lean MDEA) is compared to
the mass flow of the regenerated MDEA stream and the losses are obtained, Table 4.15.
Table 4.15 Estimating the process losses
LEAN MDEA
COMPONENTS
H2O
MDEAmine
REGENERATED
MDEA
MASS FLOWS
(kg/h)
192 136.5
96 893.5
190 616.7
96 892.5
LOSS
(kg/h)
1 519.8 (0.79 %)
1 (0.001 %)
As seen from Table 4.15 there is much bigger loss of water than of MDEA. Compared with
the previous two cases, in this case the loss is the highest (of water and amine). The water
loss is compensated by a makeup stream right after the heat exchanger, while the MDEA is
considered to be an insignificant loss and a makeup stream is not designed.
After the losses are fulfilled, the stream is cooled to 44°C and the pressure is increased to 70
bar. The last step of the design of this gas sweetening process is to install a recycle unit,
which will make a closed system.
As a conclusion, the tertiary amine – MDEA, is considered to be a good sweetening
chemical, based on the received results. The designed methyldiethanolamine aqueous
solution, 35 wt% MDEA and 65 wt% water, was able to remove the Hydrogen Sulfide to the
desired concentration of 0.0004%. The losses in this case are a little bit higher than the one in
the previous two cases, but the advantages of this process are that the flow rate of the amine
solution is the lowest as well as the reboiler duty.
36 | P a g e
5. Cost estimation of the process
This chapter will briefly review the cost of the major items of equipment that are used in the
sweetening simulations with HYSYS. In this way an approximate cost of the whole
sweetening process can be estimated. Furthermore, the equipment used in the three cases is
the same and more or less the operating conditions and specifications are the same. The main
differences that will distinguish the cases are the cost of the given amine, water, weight per
cent in the solution, flow rate of the solvent and reboiler duty.
Based on the tables, charts and examples in Chemical Engineering Design [42] cost
approximation of the following units have been made:






Absorption column – diameter = 1.5 m; height = 17 m; trays height = 14 m;
column material – steel; tray material – stainless steel; operating pressure 70
bar;
Regeneration column – diameter = 1.5; height 20 m; trays = 18; vessel and
trays – stainless steel; operating pressure – 2 bar;
Reboiler – fixed tube sheet, area – 27.1 m2 (MEA), 24.7 m2 (DEA), 23.9 m2
(MDEA), shell – carbon steel, tubes – stainless steel, operating pressure – 2
bar;
Condenser – fixed tube sheets, area – 32.6 m2 (MEA), 31.7 m2 (DEA), 29.9
m2 (MDEA), shell – carbon steel, tubes – stainless steel, operating pressure – 1
bar;
Cooler – U-tubes, area – 17.8 m2, shell and tubes – carbon steel, operating
pressure 2 bar;
Heat exchanger - U-tubes, area – 60.3 m2, shell – carbon steel, tube –
stainless steel, operating pressure 2 bar.
Now that all these units are specified an approximate calculation of each unit can be done.
All prices are time based to mid 2004.
Absorption column
Bare vessel cost - $ 60 000, material factor – 2, pressure factor – 2.2. Now that these data is
obtained from the charts, the vessel cost can be obtained.
Vessel cost = 60 000 x 2 x 2.2 = $ 264 000.
Tray’s cost = $ 1400/m3.
Volume of trays = (π/4) x 14 = 11 m3.
Cost of column’s trays = 11 x 1400 = $ 15 400.
Total cost of the column – 264 000 + 15 400 = 279 400, say $ 280 000.
37 | P a g e
Regeneration column
Bare vessel cost - $ 73 000, material factor – 2, pressure factor – 1.
Vessel cost – 73 000 x 2 x 1 = $ 146 000.
Cost of a tray – material factor – 1.7, 1.7 x $ 530 = $ 901.
Total cost of trays – 18 x 901 = 16 218, say $ 16 250.
Total cost of the column – 146 000 + 16 250 = $ 162 250.
Reboiler
Bare cost ($)
Type factor
Pressure factor
Purchased cost ($)
MEA case
41 000
1
0.8
32 800
DEA case
39 000
1
0.8
31 200
MDEA case
38 000
1
0.8
30 400
MEA case
35 000
1
0.8
28 000
DEA case
33 000
1
0.8
26 400
MDEA case
30 000
1
0.8
24 000
Condenser
Bare cost ($)
Type factor
Pressure factor
Purchased cost ($)
Cooler
Bare cost of the cooler - $ 11 000, pressure factor – 1, type factor – 0.85.
Purchased cost = 11 000 x 1 x 0.85 = $ 9 350.
Heat exchanger
Bare cost of the heat exchanger - $ 68 000, pressure factor – 1, type factor – 0.85.
Purchased cost = 68 000 x 1 x 0.85 = $ 57 800.
Now that an approximate value is calculated for the equipment used for removing H2S from
the gas stream, calculations for the value of the used water and chemical should be made for
each case. After these final calculations are done a summary can be made.
In Table 5.1 is presented the value estimation of used MEA, DEA and MDEA, as well as the
water used for each of the three cases. The prices are based on mid 2004 and the used
quantities are given for a day (24 hours) based on the simulations made with HYSYS.
38 | P a g e
Table 5.1 Chemical and water cost for all three cases
MEA
DEA
MDEA
Used water (kg/d)
7 013 810.4
5 514 417.6
4 611 276
Total cost of used water
($)
3 506 905.2
2 757 208.8
2 305 638
825 640.8
1 875 547.2
2 325 444
1 271 486.8
3 188 430.2
4 209 053
4 778 392
5 945 639
6 514 691
Amine used (kg/d)
Total cost of used amine
($)
Total costs ($)
After all calculations are made it can be concluded that MEA is the most economical amine,
this conclusion is based on the results from the equipment costs as well as the solution costs
for a day (Table 5.2). Even though MEA case has the highest prices for Condenser and
Reboiler units they are very insignificant compared to the difference of the prices for the
solution needed for a day.
Table 5.2 Summary table of all estimated cost for the three cases
MEA
DEA
MDEA
Absorption column ($)
280 000
280 000
280 000
Regeneration column ($)
162 250
162 250
162 250
Reboiler ($)
32 800
31 200
30 400
Condenser ($)
Cooler ($)
28 000
9 350
26 400
9 350
24 000
9 350
Heat exchanger ($)
57 800
57 800
57 800
Cost of the solution for a
day ($)
4 778 392
5 945 639
6 514 691
TOTAL COST ($)
5 348 592
6 512 639
7 078 491
39 | P a g e
6. Conclusion and Discussion
In conclusion, this case study was very interesting due to the non-typical field KhurmalaErbil. With its high concentrations of H2S and CO2 the design of the sweetening process was
quite challenging. Nevertheless, pipeline specifications for the Hydrogen Sulfide were met –
0.0004%, in all three cases.
Table 6.1 shows a summary of the designed lean amine stream, cost estimation of the process
and duty necessary for the reboiler, etc.
Table 6.1MEA, DEA, MDEA – design and specifications
MEA
DEA
MDEA
Temperature (°C)
44
44
44
Pressure (bar)
70
70
70
Flow (m3/h)
328.8
302.1
284.2
Reboiler Duty (Btu/h)
8.3e8
10
7.9e7
25
7.3e7
35
0.0004
0.0004
0.0004
Water losses (kg/h)
1 494.3 (0.51%)
1 729.8 (0.75 %)
1 519.8 (0.79 %)
Amine losses (kg/h)
2.5(0.007%)
5 348 592
0.1 (0.00013 %)
6 512 639
1 (0.001 %)
7 078 491
Weight %
H2S concentration
(in sweet gas) (mole%)
TOTAL COST ($)
From the table above it can be observed that all three cases are designed with the same
temperature and pressure. MEA case is with the highest flow rate and respectively with the
highest duty, but on the other hand is with the lowest weight % concentration in the aqueous
solution. DEA case is the middle and MDEA case is with the lowest flow rate and reboiler
duty, but it has the highest cost and weight % in the solution.
Based on this summary, it can be concluded that MEA case is the most efficient and the least
expensive amine that can be used for the chosen field of study.
40 | P a g e
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41 | P a g e
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42 | P a g e
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43 | P a g e
Appendix A
Physical Properties of MEA, DEA, MDEA
44 | P a g e
Appendix B
MEA Process Simulation in Aspen HYSYS
45 | P a g e