NREL is a national laboratory of the U.S. Department of Energy, Office of Energy 

Efficiency & Renewable Energy, operated by the Alliance for Sustainable Energy, LLC.

Contract No. DE-AC36-08GO28308

Survey and Down-Selection 

of Acid Gas Removal Systems 

for the Thermochemical 

Conversion of Biomass to 

Ethanol with a Detailed Analysis 

of an MDEA System 

Task 1: Acid Gas Removal Technology 

Survey and Screening for 

Thermochemical Ethanol Synthesis

March 2009

Task 2: Detailed MDEA Process 

Analysis

September 2009

Nexant Inc.

San Francisco, California

Subcontract Report

NREL/SR-5100-50482

May 2011 
NREL is a national laboratory of the U.S. Department of Energy, Office of Energy 

Efficiency & Renewable Energy, operated by the Alliance for Sustainable Energy, LLC.

National Renewable Energy Laboratory

1617 Cole Boulevard

Golden, Colorado 80401

303-275-3000 • www.nrel.gov

Contract No. DE-AC36-08GO28308

Survey and Down-Selection 

of Acid Gas Removal Systems 

for the Thermochemical 

Conversion of Biomass to 

Ethanol with a Detailed 

Analysis of an MDEA System 

Task 1: Acid Gas Removal Technology 

Survey and Screening for 

Thermochemical Ethanol Synthesis

March 2009

Task 2: Detailed MDEA Process 

Analysis

September 2009

Nexant Inc.

San Francisco, California

NREL Technical Monitor: Andy Aden

Prepared under Subcontract No. KAFT-8-882786-01

Subcontract Report

NREL/SR-5100-50482

May 2011 
This publication was reproduced from the best available copy

submitted by the subcontractor and received no editorial review at NREL.

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Task 1 

Acid Gas Removal Technology 

Survey and Screening for 

Thermochemical Ethanol Synthesis 

NREL Task Order No. KAFT-8-882786-01 

March 2009 

Prepared for: 

United States Department of Energy / National Renewable Energy 

Laboratory (NREL) 
 

i

Contents 

Section Page 

 

Executive Summary....................................................................................................................................... 1 

Section 1 Introduction .............................................................................................................................. 5

Section 2 Study Approach ....................................................................................................................... 7 

Section 3 Chemical Absorption Processes ............................................................................................ 13 

Section 4 Physical Absorption Processes.............................................................................................. 22 

Section 5 Hybrid Processes................................................................................................................... 29 

Section 6 Hot Carbonate-Based Process .............................................................................................. 32 

Section 7 RTI Solid Absorbent-Based Process...................................................................................... 37 

Section 8 Liquid Phase Oxidation Processes ........................................................................................ 41 

Section 9 Membrane Separation Processes.......................................................................................... 44 

Section 10 Conclusions and Recommendations...................................................................................... 47 

References ............................................................................................................................................... 49 

Figures 

Figure ES-1 Order-of-Magnitude Comparison of CO and Hydrocarbons Absorption in a Selexol and 

MDEA Process .............................................................................................................................................. 3 

Figure 1-1 Overview of NREL Thermochemical Ethanol Synthesis Process ............................................. 5 

Figure 2-1 Gas Cleanup and Conditioning Process of NREL Alcohol Synthesis Design ........................... 8 

Figure 2-2 Process Selection Chart for Simultaneous H2S and CO2 Removal ........................................ 11 

Figure 2-3 Process Selection Chart for H2S Removal.............................................................................. 12 

Figure 2-4 Process Selection Chart for CO2 Removal ............................................................................. 12 

Figure 3-1 Typical Amine Process Flow Diagram.................................................................................... 17 

Figure 4-1 Typical Selexol Process Flow Diagram .................................................................................. 24 

Figure 4-2 Typical Selective Selexol Process Flow Scheme ................................................................... 25 

Figure 4-3 Order-of-Magnitude Comparison of CO and Hydrocarbons Absorption in a Selexol and 

MDEA Process ............................................................................................................................................ 27 

Figure 5-1 Typical Sulfinol Process Flow Diagram .................................................................................. 30 

Figure 6-1 Typical Hot Potassium Carbonate Single-Stage Process....................................................... 33 

Figure 6-2 Typical Hot Potassium Carbonate Split-Flow Process............................................................ 34 

Figure 6-3 Typical Hot Potassium Carbonate Two-Stage Process .......................................................... 34 

Figure 7-1 Simplified Overall WGCU Process Scheme ........................................................................... 39 

Figure 7-2 High Temperature Desulfurization (HTDS) Process Flow Scheme ........................................ 39 

Figure 8-1 Conventional LO-CAT System ............................................................................................... 42 

Figure 9-1 One-Stage Membrane System............................................................................................... 45 

Figure 9-2 Two-Stage Membrane System............................................................................................... 46 
 

ii

Tables 

Table ES-1 Cost Components of the Gas Cleanup and Conditioning Process Area .................................... 1 

Table 2-1 Design Basis for AGR Process Screening................................................................................. 8 

Table 3-1 Chemical Absorption Processes Summary Table.................................................................... 19 

Table 3-2 Estimated Amine Circulation Rates and Regeneration Energy for NREL Design.................... 21 

Table 4-1 Selexol Solvent Relative Solubility of Gases ........................................................................... 24 

Table 4-2 Physical Absorption Processes Summary Table ..................................................................... 28 

Table 5-1 Hybrid Processes Summary Table .......................................................................................... 31 

Table 6-1 Hot Potassium Carbonate Processes Summary Table............................................................ 36 

Table 8-1 Liquid Phase Oxidation Processes Summary Table................................................................ 43 
 

iii

Acronyms 

AGR acid gas removal 

BFW boiler feed water 

DEA diethanolamine 

DGA diglycolamine 

DIPA diisopropanolamine 

DSRP direct sulfur recovery process 

HCs hydrocarbons 

HTDS high temperature desulfurization 

IGCC integrated gasification combined cycle 

MDEA methyldiethanolamine 

MEA monoethanolamine 

MM million (1x106

) 

MW molecular weight 

NREL National Renewable Energy Laboratory 

O&M operating and maintenance 

ppm parts per million 

ppmv parts per million volume 

RTI Research Triangle Institute 

SCF standard cubic feet (1 atm, 60°F) 

TIC total installed cost 

UOP Universal Oil Products 

WGCU warm gas clean up 
 

Task 1: Acid Gas Removal Technology Survey and Screening for Thermochemical Ethanol Synthesis 

United States Department of Energy/National Renewable Energy Laboratory 

1 

 Executive Summary

NREL’s recent published report, “Thermochemical Ethanol via Indirect Gasification and Mixed 

Alcohol Synthesis of Lignocellulosic Biomass” (NREL/TP-510-41168), evaluated the technoeconomics of a biomass gasification process to produce synthesis gas (syngas) for ethanol 

synthesis. The thermochemical ethanol production process consists of five major process 

areas: feed handling and preparation, gasification, gas cleanup and conditioning, alcohol 

synthesis, and product separation. The report showed that the syngas cleanup and conditioning 

section, which consists of tar reforming and catalyst regeneration, syngas cooling and 

compression, and acid gas removal (AGR) operations, has the highest cost component of all the 

process areas, accounting for approximately 39% of the total installed plant cost (Table ES-1). 

As such, this process section has a significant impact on the overall thermochemical ethanol 

production economics. In NREL’s current design, acid gas removal is performed by an MEA 

system (a chemical absorption process). NREL wishes to examine the applicability of other acid 

gas removal processes (e.g. physical absorption systems of Selexol and Rectisol) to the 

thermochemical ethanol synthesis design, and to develop a detailed acid gas removal process 

design in order to better assess its technical and economic impacts on the overall process. 

Table ES-1 Cost Components of the Gas Cleanup and Conditioning Process Area 

Units of the gas clean up conditioning process Cost, as % of TIC 

 Tar Reforming and Catalyst Regeneration 13.4 

 SynGas Cooling 3.4 

 SynGas Compression 12.9 

 AGR 8.9 

Total 38.6 

Objectives of the current study are to conduct a survey and screening of various acid gas 

removal processes in order to evaluate their capability to meet the specific design requirements 

provided by NREL, and to recommend applicable acid gas removal options for a more detailed 

analysis in Task 2. In NREL’s design, the raw syngas enters the AGR unit at 420 psia, and 

contains 234 ppmv H2S, 11 mol% CO2, and trace amounts of hydrocarbon impurities. The 

required purity of the treated gas for alcohol synthesis is 50 ppmv H2S and 5.1 mol% CO2. This 

corresponds to 79% removal of H2S and 59% removal of CO2. 

The capability and applicability of the various acid gas removal processes are qualitatively 

evaluated based on the following criteria: 

 Feed gas compositions and conditions 

 The desired purity of the treated gas 

 Compatibility of process solvent with impurities in the feed gas 

 Selectivity of H2S over CO2

 Commercial experience of the AGR processes with synthesis gases 

 Relative costs of the AGR processes. 
 

Task 1: Acid Gas Removal Technology Survey and Screening for Thermochemical Ethanol Synthesis 

United States Department of Energy/National Renewable Energy Laboratory 

2 

The acid gas removal requirements for the NREL ethanol synthesis design can be met by a 

number of chemical and physical absorption processes. The final selection would most likely be 

based on considerations of plant costs (capital and O&M), commercial experience, and process 

reliability and operational flexibility. Other important factors that influence the design and 

selection of an AGR system must also be considered, including: 

 Syngas compression before AGR – Since syngas compression accounts for 13% of the 

total biomass-to-ethanol plant cost, the AGR unit should operate at the lowest possible 

pressure to minimize the power and equipment cost penalties associated with the 

compression of the syngas up to the operating pressure of the AGR unit. For this study, 

all AGR options are qualitatively evaluated at the fixed operating pressure of 420 psia, 

which eliminates the need to consider the syngas compression differences. In the NREL 

design, the alcohol synthesis is carried out at 990 psia. Therefore, the operating 

pressure range in which the AGR unit can theoretically be operated would be from the 

tar reformer pressure of 15 psia up to the alcohol synthesis pressure of 990 psia. For 

AGR processes that favor high operating pressures, the detailed analysis in Task 2 

requires a trade-off analysis to assess the power and cost penalties associated with the 

additional compression. 

 Absorption of product gases – The AGR feed gas contains a high concentration of CO 

and small quantities of hydrocarbons, ranging from methane to pentane, benzene and 

tar, which would be absorbed to some extent into the treating solutions. Absorption of 

CO and hydrocarbons is expected to be greater for physical absorption than chemical 

absorption systems due to the higher solubility of these compounds in a physical 

solvent. The loss of CO through absorption, if significant, would negatively affect the 

yield of the alcohol synthesis process and the overall plant efficiency. 

Equally important, the absorption of CO and hydrocarbon gases into the treating 

solvents has an environmental impact on the overall process since these gases, which 

exit the AGR unit with the acid gas stream, would be vented with the CO2 stream from 

the LO-CAT sulfur recovery unit (Figure 1-1, Section 1). The incineration of the CO2

vent gas to destroy the CO and hydrocarbons would consume a significant amount of 

product gas as fuel as well as increase the overall plant cost. 

The concentrations of CO (a criteria pollutant) and hydrocarbons (greenhouse gases) in 

the CO2 vent stream can be minimized by depressuring the rich solution and recycling 

the flashed gas to the absorber. The flash pressure can, however, influence the level of 

contaminant in the CO2 vent gas. At low flash pressures, less CO and hydrocarbons will 

be present in the CO2 vented stream, but the cost for recycling the flashed offgas to the 

absorber would increase and can negatively impact the overall process economics. In 

comparison to a chemical absorption system, a physical solvent system is expected to 

have higher concentrations of CO and hydrocarbons in the vented stream. Preliminary 

process simulation of this effect was preformed and the results are shown in Figure ES-
 

1. The allowable CO and hydrocarbons concentrations in the CO2 vent stream are sitespecific and need to be established prior to the start of Task 2 detailed analysis. 

100

1000

10000

100000

0 50 100 150 200 250 300 350 400 450 Decreasing Flash Drum Pressure

Increasing CO & Hydrocarbons Concentrations in Vent Gas

CO in vent gas - Selexol system

CO in vent gas - MDEA System

HCs in vent gas - Selexol system

HCs in vent gas - MDEA system

Low Flash Pressures High Flash Pressures

Note: The graph is intended to 

show the relative emissions of the 

Selexol vs. the MDEA systems, 

thus the units are not provided. 

Actual emissions of each system 

will be determined in Task 2.

Figure ES-1 Order-of-Magnitude Comparison of CO and Hydrocarbons Absorption in a 

Selexol and MDEA Process 

 Selective removal of H2S vs. CO2 – The alcohol synthesis catalyst requires an H2S 

concentration of up to 50 ppmv to maintain catalyst activity, and it can also tolerate a 

CO2 concentration of up to 5 mol%. The type of AGR solvents used will affect the 

degree of H2S and CO2 removal. Primary and secondary amines are known for their 

non-selectivity, and can remove comparatively more CO2 than H2S from the feed syngas 

stream. To meet the 50 ppmv H2S specification in the treated clean syngas, excess 

amount of CO2 would probably also be removed, yielding a gas with a much lower CO2 

content than 5 mole%. The exact CO2 vs. H2S content in the cleaned gas would need to 

be confirmed by process simulation. 

In comparison, a tertiary amine process such as MDEA and physical solvent systems of 

Selexol and Rectisol can selectively remove first H2S then CO2 from a feed gas stream, 

and thus in principal, has a higher degree of control with meeting the needed clean 

syngas specification. Also, MDEA has a lower circulation rate and regeneration energy 

requirement than primary and secondary amine processes of MEA and DEA 

respectively. 

Task 1: Acid Gas Removal Technology Survey and Screening for Thermochemical Ethanol Synthesis 

United States Department of Energy/National Renewable Energy Laboratory 

3 
 

Task 1: Acid Gas Removal Technology Survey and Screening for Thermochemical Ethanol Synthesis 

United States Department of Energy/National Renewable Energy Laboratory 

4 

A qualitative assessment, taking into consideration the factors mentioned above and the 

process screening criteria cited in the report, suggests that MDEA and Selexol are both viable 

AGR options for the thermochemical ethanol synthesis application to warrant of a more detailed 

design/system analysis study. The Selexol process is recommended over Rectisol, because, in 

general, it has a lower cost. The Rectisol process is generally known for its demonstrated ability 

to provide stringent purity syngas for applications such as coal-to-chemicals, synthetic natural 

gas, and Fischer-Tropsch liquids production. It can achieve complete removal of sulfur (i.e., 

<0.1 ppm H2S plus COS) as well as removal of HCN, NH3 and trace metal carbonyls, which can 

be poisonous to synthesis catalysts. Thus, unless there are trace contaminants in the raw gas 

that have not yet been identified and that can potentially damage the ethanol synthesis catalyst, 

a more costly Rectisol unit would not be justified for the NREL design. 

Limited analysis, based on in-house cost and design data available, suggests that while, in 

general, a Selexol process would have a higher plant cost than a MDEA process at the given 

operating conditions and acid gas partial pressures, it may become more economical as the 

system pressure increases. The effect must be evaluated against the cost of syngas 

compression and the impact of increased product gas absorption at high pressures with the use 

of the Selexol system. 

A detailed process design and cost estimation analysis of the MDEA system for thermochemical 

ethanol production is recommended, as Task 2 activity. In addition, a comparative study for the 

Selexol system, taking into consideration the various plant design parameters such as the 

overall syngas compression cost and the environmental impact of CO and hydrocarbons 

emissions requirement, etc., is also recommended. 
 

Section 1 Introduction

1.1 BACKGROUND 

For alcohol synthesis processes, the removal of acid gas compounds is an important process to 

prepare the syngas to achieve the required purity in order to avoid poisoning and deactivation of 

the synthesis catalyst. Acid gas generally includes H2S, CO2, HCN, and organic sulfur 

compound such as COS, CS2, and mercaptans. Figure 1-1 illustrates a simplified block flow 

diagram of the NREL thermochemical ethanol synthesis design. The feed gas to the AGR unit 

is relatively clean, with CO2 and H2S being the primary acid gas compounds. In the current 

design, an MEA system is used for the removal of acid gases from the syngas stream. A LOCAT unit is used for sulfur recovery. 

Source: Reference [1] 

Figure 1-1 Overview of NREL Thermochemical Ethanol Synthesis Process 

1.2 OBJECTIVES 

The objectives of this study, Task 1, are to conduct a survey and screening of various acid gas 

removal processes in order to evaluate their capability to meet the specific design requirements 

provided by NREL for thermochemical ethanol production, and to recommend suitable acid gas 

removal options for the detailed analysis in Task 2. As agreed with NREL, the following types of 

acid gas removal processes are to be surveyed and screened: 

1. Chemical absorption (used in current NREL design) 

2. Physical absorption 

3. Hybrid processes 

Task 1: Acid Gas Removal Technology Survey and Screening for Thermochemical Ethanol Synthesis 

United States Department of Energy/National Renewable Energy Laboratory 

5 
 

Task 1: Acid Gas Removal Technology Survey and Screening for Thermochemical Ethanol Synthesis 

United States Department of Energy/National Renewable Energy Laboratory 

6 

4. Hot carbonate-based processes 

5. RTI solid absorbent-based process 

6. Liquid phase oxidation 

7. Membrane separation 

The qualitative process screening performed to select the applicable acid gas removal 

processes for the NREL alcohol synthesis design is based on the design basis provided by 

NREL, the screening criteria outlined in the Methodology Section, and the performance data of 

the AGR process developed from the survey. 
 

Task 1: Acid Gas Removal Technology Survey and Screening for Thermochemical Ethanol Synthesis 

United States Department of Energy/National Renewable Energy Laboratory 

7 

Section 2 Study Approach

This study begins with a survey of various types of acid gas removal process, followed by a 

qualitative screening to determine the applicability of each process for thermochemical ethanol 

synthesis. The following sections discuss the types of acid gas removal technologies included 

in the survey, and the design basis and the screening criteria used for the qualitative process 

screening. 

2.1 ACID GAS REMOVAL TECHNOLOGIES SURVEYED 

This study conducts a survey of the following acid gas removal technologies. The survey 

includes an overview and performance of each process and discusses the process advantages 

and disadvantages as compared to the chemical absorption process. 

1. Chemical absorption (used in current NREL design) 

2. Physical absorption 

3. Hybrid processes 

4. Hot carbonate-based processes 

5. RTI solid absorbent-based process 

6. Liquid phase oxidation 

7. Membrane separation 

2.2 QUALITATIVE PROCESS SCREENING OF ACID GAS REMOVAL PROCESSES 

A qualitative process screening is performed to assess the suitability of the acid gas removal 

processes for the NREL thermochemical ethanol synthesis design, using the design basis and 

screening criteria below. 

2.2.1 Design Basis 

Figure 2-1 illustrates a simplified block flow diagram of the syngas cleanup and conditioning 

process in the NREL thermochemical ethanol production design. The feed gas to the AGR and 

the treated gas specifications are provided by NREL and are used as basis for the qualitative 

process screening. The design basis is shown in Table 2-1. 
 

Figure 2-1 Gas Cleanup and Conditioning Process of NREL Alcohol Synthesis Design 

Table 2-1 Design Basis for AGR Process Screening 

Stream No. 1 2 

Feed Gas to AGR Treated Gas 

Total Flow, lbmol/hr 17,942 16,740 

Temperature, °F 110 110 

Pressure, psia 420 415 

Vapor Fraction 1 1 

Compositions mole% mole% 

H2 42.9 45.9 

H2O 0.3 0.3 

CO 42.9 46.0 

N2 0.8 0.9 

CO2 11.4 5.1 

H2S 234 ppmv 50 ppmv 

NH3 82 ppmv 88 ppmv 

Methane (CH4 ) 1.4 1.5 

n-Butane (C4H10) 0.03 0.03 

Ethane (C2H6) 13 ppmv 14 ppmv 

Ethylene (C2H4) 0.1 0.1 

Acetylene (C2H2) 0.0 0.0 

Propane (C3H8) 0.2 0.2 

Pentane + 42 ppmv 46 ppmv 

Benzene (C6H6) 3 ppmv 3 ppmv 

Tar (C10H8) 0.4 ppmv 0.5 ppmv 

Source: Reference [1] 

Task 1: Acid Gas Removal Technology Survey and Screening for Thermochemical Ethanol Synthesis 

United States Department of Energy/National Renewable Energy Laboratory 

8 
 

Task 1: Acid Gas Removal Technology Survey and Screening for Thermochemical Ethanol Synthesis 

United States Department of Energy/National Renewable Energy Laboratory 

9 

2.2.2 Process Screening Criteria 

The qualitative process screening assesses the following criteria for applicability and overall 

consideration with selecting an AGR process for the NREL thermochemical ethanol synthesis 

design: 

 Feed gas compositions and conditions – The feed gas conditions and partial pressure of 

acid gas components influence the performance and selection of AGR process. 

Numerous graphs have been published showing the approximate guidelines for the 

selection of various solvent-based processes based on the acid gas partial pressure in 

the feed gas and in the treated gas[ ] 14 . Figure 2-2 shows a general guideline that can be 

used to select an appropriate process for a feed stream that contains both CO2 and H2S. 

Similar general selection guidelines for removal of only H2S or CO2 are shown in Figures 

2-3 and 2-4 respectively. Selection process depends on the acid gas partial pressure in 

the feed against the desired acid gas concentration in the treated gas. 

 The required purity of the treated gas – The required clean gas specifications usually 

depend on its end use. In the NREL thermochemical ethanol production process, the 

alcohol synthesis catalyst requires an H2S concentration of up to 100 ppm (50 ppmv) to 

maintain catalyst activity, and it can also tolerate a CO2 concentration of up to 5 mol%. 

According to NREL, the CO2 concentration in the clean gas can range from 1 to 5 mol%. 

On this basis, the acid gas removal process is not required to achieve deep H2S and 

CO2 removal. Based on the feed gas compositions and the treated gas specifications, 

the required acid gas removal is 79% of H2S and 59% of CO2. 

 Chemical compatibility – The primary concern with chemical compatibility is the 

irreversible reactions of the solvent with impurities such as HCN or organic sulfur 

compounds (COS, CS2, and mercaptans), which result in solvent degradation and 

solvent losses. As shown in the design basis, the impurities in the feed gas to the AGR 

unit in the NREL design consist of mostly CO2 and H2S, a small amount of ammonia, 

and essentially no organic sulfur compounds. Thus, chemical incompatibility of the 

treating solvent with impurities in the feed gas is unlikely to be a concern for the current 

design. 

However, if experimental data show that the biomass-derived syngas does indeed 

contain any of these impurites (HCN, COS, CS2, or mercaptans), then MEA may not be 

a suitable solvent due to the irreversible reactions of MEA with COS and CS2. All other 

amine solutions (DEA, DGA, DIPA, and MDEA) and physical solvents evaluated in this 

study are able to tolerate such impurities. 

 Selectivity of H2S over CO2 – In the NREL’s design, although some level of sulfur 

removal is required to meet the specifications for alcohol synthesis, the bulk of the acid 

gas removal is CO2. Primary and secondary amines are known for their non-selectivity, 

and can remove comparatively more CO2 than H2S from the feed syngas stream. To 

meet the 50 ppmv H2S specification in the treated clean syngas, excess amount of CO2
 

Task 1: Acid Gas Removal Technology Survey and Screening for Thermochemical Ethanol Synthesis 

United States Department of Energy/National Renewable Energy Laboratory 

10 

would probably also be removed, yielding a gas with a much lower CO2 content than 5 

mole%. The exact CO2 vs. H2S content in the cleaned gas would need to be confirmed 

by process simulation. 

In comparison, a tertiary amine process such as MDEA and physical solvent systems of 

Selexol and Rectisol can selectively remove first H2S then CO2 from a feed gas stream, 

and thus in principal, has a higher degree of control with meeting the needed clean 

syngas specification. Also, MDEA has a lower circulation rate and regeneration 

requirement than primary and secondary amine process of MEA and DEA respectively. 

 Absorption of product gases – The NREL AGR feed gas contains a high concentration of 

CO and small quantities of hydrocarbons, ranging from methane to pentane, benzene 

and tar, which would be absorbed to some extent into the treating solutions. Absorption 

of CO and hydrocarbons is expected to be greater for physical absorption than chemical 

absorption systems due to the higher solubility of these compounds in a physical 

solvent. The loss of CO through absorption, if significant, would negatively affect the 

yield of the alcohol synthesis process and the overall plant efficiency. Additionally, the 

absorption of CO and hydrocarbon gases into the treating solvents has an environmental 

impact on the overall process since these gases would be vented with the CO2 stream 

from the LO-CAT sulfur recovery unit. 

 Syngas compression before AGR – Since syngas compression accounts for 13% of the 

total biomass-to-ethanol plant cost, the AGR unit should operate at the lowest possible 

pressure to minimize the power and equipment cost penalties associated compression 

of the syngas up to the operating pressure of the AGR unit. For this study, all AGR 

options are qualitatively evaluated at the fixed operating pressure of 420 psia, which 

eliminates the need to consider the syngas compression differences. In the NREL 

design, the alcohol synthesis is carried out at 990 psia. Therefore, the operating 

pressure range in which the AGR unit can theoretically be operated would be from the 

tar reformer pressure of 15 psia up to the alcohol synthesis pressure of 990 psia. For 

AGR processes that favor high operating pressures, the detailed analysis in Task 2 

requires a trade-off analysis to assess the power and cost penalties associated with the 

additional compression.

 Commercial experience – The various acid gas removal processes surveyed in this 

study, with the exception of RTI’s Warm Gas Clean Up process, are commercially 

available. Selected chemical and physical absorption processes have been 

commercially demonstrated with coal-derived gases. Since the raw gas in the NREL 

thermochemical ethanol production is relatively clean in comparison to a typical syngas 

from coal gasification, its treatment should be relatively easier. Thus, commercially 

demonstrated acid gas removal processes for coal-gasification derived gases are a good 

starting point for the preliminary screening exercise. 

 Relative costs of processes – The acid gas removal requirement for the NREL ethanol 

synthesis can be met by a number of chemical and physical absorption processes. The 
 

final selection ultimately depends on economics of the integrated process, and not just of 

the AGR process by itself. The integration effects of the AGR system on the overall 

thermochemical ethanol production process also need to be considered. PARTIAL PRESSURE ACID GAS IN FEED, PSI

1

10

100

1000

0.1 1 10 100

LINE OF EQUAL INLET AND OUTLET PRESSURE

(DRIVING FORCE AT EQUILIBRIUM)

Figure 2-2 Process Selection Chart for Simultaneous H2S and CO2 Removal

Task 1: Acid Gas Removal Technology Survey and Screening for Thermochemical Ethanol Synthesis 

United States Department of Energy/National Renewable Energy Laboratory 

11 
 

1

10

100

1000

0.01 0.1 1 10

Figure 2-3 Process Selection Chart for H2S Removal 

1

10

100

1000

0.1 1 10 100

Figure 2-4 Process Selection Chart for CO2 Removal 

Source: reference [14] 

Task 1: Acid Gas Removal Technology Survey and Screening for Thermochemical Ethanol Synthesis 

United States Department of Energy/National Renewable Energy Laboratory 

12 
 

Task 1: Acid Gas Removal Technology Survey and Screening for Thermochemical Ethanol Synthesis 

United States Department of Energy/National Renewable Energy Laboratory 

13 

Section 3 Chemical Absorption Processes

An overview of the chemical absorption process is presented in this section. The process 

characteristics and operating parameters are provided in a summary table at the end of the 

section. 

3.1 PROCESS OVERVIEW 

In a chemical absorption process, the acid gas components react with the solvent to form a 

chemical compound. The solvent can be regenerated by the application of heat to strip the 

absorbed gases. Typical of chemical absorption is the amine process. Amines are commonly 

used in the petroleum refining and natural gas processing industries for acid gas removal. 

Additionally, amine (e.g. MDEA) has also been used to treat synthesis gas from coal 

gasification. 

The three basic types of amines that are available commercially include: primary amine, such 

as MEA or DGA, secondary amine, such as DEA or DIPA, and tertiary amines, such as MDEA. 

Primary and secondary amines react with H2S and CO2 as follows: 

H2S + RNH2 ÅÆ RNH3HS 

CO2 + RNH2 + H2O ÅÆ RNH3HCO3

where R represents the hydrocarbon and the hydroxyl groups in amine solution

 

The reactions proceed to the right at increased acid gas partial pressures and/or low 

temperatures and are reversed at low partial pressures and/or high temperatures. In most 

primary and secondary amine solutions, H2S and CO2 react simultaneously with the solution, 

making selective recovery difficult. Primary and secondary amines react with CO2 to form 

carbamate. Carbamate formation limits the solution loading capacity for CO2 to approximately 

0.5 mole of CO2 per mole of amine. Tertiary amines, however, cannot react directly with CO2 to 

form carbamate. In tertiary amine solutions, the reaction with CO2 forms bicarbonate, and the 

CO2 reaction rate is significantly slower than that of H2S with tertiary amines; therefore, tertiary 

amines generally exhibit greater selectivity for H2S in the presence of CO2. 

The removal of organic sulfur compounds (COS and CS2) is generally accomplished by the 

hydrolysis unit upstream of the amine unit. The hydrolysis unit converts the COS and CS2 into 

H2S, which can then be removed by the amine unit. COS and CS2 can be partially removed by 

some amines, especially primary and secondary ones. However, removal of such gas 

impurities subject the solution to degradation as those impurities react, either reversibly or 

irreversibly, with the amines to form degradation products. In irreversible reactions, such as 

those of MEA with COS and CS2, formation of nonregenerable degradation products result in 

excessive solvent losses if the gas contains high concentrations of those sulfur compounds. 

After prolonged use, accumulation of the degradation products in the solution reduces its 

absorption efficiency. These contaminants are commonly removed from the solution by thermal 

reclaiming. 
 

Task 1: Acid Gas Removal Technology Survey and Screening for Thermochemical Ethanol Synthesis 

United States Department of Energy/National Renewable Energy Laboratory 

14 

In addition to the basic amines, proprietary amines are also offered by many chemical 

companies, such as UOP (Amine Guard and UCARSOL), BASF (Activated MDEA), or Shell 

(ADIP-X). Proprietary formulations, usually consisting of mixtures of amines and additives, are 

used to meet specific gas treating targets such as selective H2S removal, partial or complete 

CO2 removal, high acid gas loading, COS removal[ ] 2 , etc. 

The selection of an amine solution depends on process conditions, acid gas partial pressures, 

and purity of the treated gas. Generally, amine processes are suitable for applications with low 

to medium acid gas partial pressures as the process economic advantage declines with 

increasing partial pressures of acid gases. The economy of an amine process is largely 

determined by the solvent regeneration energy requirements, which is a function of the heats of 

reaction of H2S and CO2, and the solution circulation rate, which is based on the solution 

capacity and concentration. Increasing the solution concentration can generally reduce the 

circulation rate, thereby reducing the plant cost; however, in order to avoid equipment corrosion, 

the solution concentration should not exceed the upper limit recommended for the solvent type. 

A comparison of the basic amines is provided below. Table 3-1 summarizes the characteristics 

and performance of the various amines. 

Monoethanolamine (MEA) 

MEA is generally used in natural gas processing plants to treat gases containing only H2S and 

CO2 in low concentrations, especially when maximum removal of both impurities is required. 

MEA solutions have the following advantages: 

 High alkalinity, which increases the solution effectiveness for acid gas absorption 

 High solution capacity at moderate concentrations 

 Contaminated solutions can be reclaimed with relative ease 

The MEA process also has a number of disadvantages, including: 

 The formation of irreversible reaction products with COS and CS2, which leads to 

excessive solution losses when the raw gas contains high concentrations of those 

impurities[ ] 2

 Higher corrosion rates compared to other amines, especially if the MEA concentration 

exceeds 20% and the solution acid gas loading is high[ ] 2 

 High heat of reaction with CO2 and H2S, resulting in high energy requirements for 

solvent regeneration 

 A relatively high vapor pressure, resulting in high vaporization losses in low-pressure 

operations 

Diethanolamine (DEA) 
 

Task 1: Acid Gas Removal Technology Survey and Screening for Thermochemical Ethanol Synthesis 

United States Department of Energy/National Renewable Energy Laboratory 

15 

DEA has been used for many years to treat refinery gases which often contain COS and CS2 in 

addition to CO2 and H2S. The advantages and disadvantages of DEA as compared to MEA are 

as followed: 

 The acid gas loading is typically higher for DEA than MEA (see Table 3-1)[ ] 5

 DEA forms regenerable compounds with COS and CS2, thus, partial removal of COS 

and CS2 can be achieved without significant solution degradation[ ] 2

 Requires lower energy for solvent regeneration than MEA due to lower heats of 

reaction 

 Low solvent vapor pressure makes DEA suitable for low-pressure operations 

 The reclaiming of contaminated solutions is more complex since vacuum distillation 

may be required[ ] 2

Diglycolamine (DGA) 

DGA has been used for the treatment of both natural and refinery gases due to its capability to 

remove not only H2S and CO2, but also COS and mercaptans. The DGA process is similar to 

that of the MEA process in many respects, except that the low vapor pressure of DGA allows its 

use in relatively high concentrations (40-60 wt%), which results in lower circulation rates and 

steam consumption. 

Diisopropanolamine (DIPA) 

DIPA has been used primarily in Europe for the treatment of refinery gases and liquids that 

contain COS in addition to H2S and CO2. DIPA solutions are reported to have low steam 

consumption, to be noncorrosive, and to be capable of removing considerable amounts of COS 

without causing solution degradation[ ] 2 . Additionally, DIPA also exhibits some selectivity for 

H2S over CO2, although H2S selectivity is not as great as that of tertiary amines. Because of its 

H2S selectivity, DIPA has been used for Claus plant tail gas treating. 

Methyldiethanolamine (MDEA) 

MDEA selectively removes H2S while allowing a large fraction of CO2 to slip through 

unabsorbed. MDEA has been used for selective H2S removal in various applications, including 

natural gas processing, Claus tail gas treating, and synthesis gas treating for integrated 

gasification combined cycle (IGCC) process. MDEA offers a number of advantages, including: 

 Selective removal results in a reduction of the amount of acid gas removed, thus, 

compared to other amine systems, MDEA system is more economical with respect to 

solvent circulation rates and energy requirements[ ] 2 

 Low heats of reaction result in low regeneration energy 

 MDEA’s low corrosion rates—the least corrosive of the amines—and its low vapor 

pressure permit its use in high concentrations (up to 60 wt%), which results in lower 

circulation rates, and thereby, smaller plant size and lower plant costs[ ] 2

 MDEA has high solution capacity, and excellent thermal and chemical stability 
 

Task 1: Acid Gas Removal Technology Survey and Screening for Thermochemical Ethanol Synthesis 

United States Department of Energy/National Renewable Energy Laboratory 

16 

 MDEA is only sparingly miscible with hydrocarbons, which is an advantage over physical 

absorption processes (discussed in Section 4) when treating gases rich in hydrocarbons 

since hydrocarbon losses can be minimized 

Due to the many favorable attributes of MDEA, its use as a nonselective solvent for the removal 

of high concentrations of CO2 is of considerable interest. In order to enhance MDEA’s capability 

for CO2 removal, proprietary additives are added in various proportions to the MDEA solution to 

achieve the desired removal target. The formulation of these propriety amines can be designed 

from high H2S/CO2 selectivity to complete H2S and CO2 removal. 

3.1.1 Flow Scheme 

Figure 3-1 illustrates a typical amine process. The feed gas enters a knock-out drum to remove 

any condensate. The dry feed gas then enters at the bottom of the absorber and is contacted 

countercurrently with the lean amine solution flowing down from the top. As the solution 

contacts the gas stream, the acid gases react with the amine. The sweet gas stream exits at 

the top of the absorber. The rich amine solution from the bottom of the absorber enters a flash 

drum where any dissolved hydrocarbons are removed. The recovered hydrocarbons are 

generally used as plant fuel. The rich amine solution is heated by heat exchange with the lean 

solvent stream and enters the stripper where the acid gases are stripped from the solution as it 

flows down the column to the reboiler. The lean solution exits the stripper, exchanges heat with 

the rich solution stream, and is further cooled before it enters at the top of the absorber. The 

stripped acid gas stream is cooled to recover water and is then sent to a sulfur recovery unit for 

processing. 

An important design consideration for amine processes is corrosion. In water, H2S dissociates 

to form a weak acid while CO2 forms carbonic acid. These acids attack and corrode metal. 

Therefore, process equipment that is exposed to CO2/H2S evolution (i.e. stripper overhead) is 

normally clad with stainless steel. 
 

Source: Adapted from reference [5] 

Figure 3-1 Typical Amine Process Flow Diagram 

3.2 PROCESS APPLICABILITY TO THERMOCHEMICAL ETHANOL SYNTHESIS 

Table 3-1 provides a performance summary of the various amine processes surveyed. The 

applicability of these amine processes to the NREL ethanol synthesis design is qualitatively 

screened and assessed, based on the syngas design basis and criteria outlined in Section 2. 

In NREL’s current design, the AGR feed gas has an H2S partial pressure of 0.1 psia and a CO2

partial pressure of 48 psia. The treated gas has an H2S partial pressure of 0.02 psia and a CO2

partial pressure of 21 psia. The acid gas partial pressure is high enough that a number of AGR 

processes can be applicable. General selection guideline of Figure 2-2 shows that amine 

processes, along with hot potassium carbonate and physical solvent systems, can potentially be 

used for the NREL design. 

Amine solvents – In general, primary and secondary amines are more reactive (i.e., form 

stronger bond with) than tertiary amines to acid gases and hence more cost effective at low 

pressures, whereas MDEA being less reactive, is more competitive at high pressures. The 

NREL design feed gas is at sufficiently high pressure such that the MDEA process should be 

given a serious design consideration, given its known advantages: 

- Less corrosive, 

- More resistant to irreversible degradation by impurities as such COS and CS2 normally 

presence in syngas, of which NREL syngas doesn’t seem to contain them, 

Task 1: Acid Gas Removal Technology Survey and Screening for Thermochemical Ethanol Synthesis 

United States Department of Energy/National Renewable Energy Laboratory 

17 
 

Task 1: Acid Gas Removal Technology Survey and Screening for Thermochemical Ethanol Synthesis 

United States Department of Energy/National Renewable Energy Laboratory 

18 

- Lower heat of regeneration, 

- Relatively lower vapor pressure loss, 

- Can be designed to selectively remove H2S and CO2, if needed. 

Both MEA and MDEA are widely used in the natural gas and refinery industries, and the 

processes are amenable to analysis via several commercial process simulators. Commercial 

licensed processes normally contain proprietary additives to control degradation and/or 

augment the solvents’ absorption rate. 

Preliminary PROMAX simulation using NREL design syngas - Process modeling indicates that

the CO2 content in the clean gas from an MEA system may be below the 5 mol% level specified 

in the NREL design. Since the CO2 content in the clean gas is lower than the level that the 

synthesis catalyst can tolerate, the excess CO2 removal is acceptable, but its cost is the 

additional regeneration energy and operating cost. Tertiary amines, on the other hand, absorb 

more H2S and less CO2 compared to primary and secondary amines due to the solvents’ 

selectivity for H2S. Preliminary process modeling of an MDEA system indicates that in order to 

meet the specified 5 mol% CO2 in the clean syngas, the amount of H2S removed may be more 

than necessary, resulting in a clean gas stream with an H2S concentration below the 50 ppmv 

level required for catalyst activity. In such case, the clean gas stream could be injected with a 

small amount of dimethyl disulfide in order to control the H2S content. The cost of dimethyl 

disulfide injection is not expected to be significant due to the small quantities required; thus, the 

economic impact on the overall plant should be marginal. 

Lower MDEA circulation rate and heat of regeneration – Preliminary analysis, using data shown 

in Table 3-1, was also carried out to evaluate the circulation rate and heat of regeneration of the 

MDEA system to achieve the required acid gas removal for the NREL design (i.e., removal of 

3.4 lbmol/hr H2S and 1,195 lbmol/hr CO2). An average acid gas pickup and solution 

concentration from Table 3-1 is used, and the estimated regeneration energy is based only on 

the heat of reaction. Results are shown in Table 3-2. Of the various amine solvents compared, 

MDEA has the lowest circulation rate and regeneration energy, which should result in the most 

economical plant design. A more detailed study would be needed to clearly define its 

comparative advantages. 

 

Demonstrated MDEA commercial experience - none of the amine processes has been 

demonstrated on a commercial scale with biomass-derived syngas, but MDEA has been 

successfully used in commercial IGCC application to treat coal-derived gases. This, coupled 

with the other factors cited above, would suggest that it should be a good candidate for the 

NREL biomass-derived syngas clean up application. 
Task 1: Acid Gas Removal Technology Survey and Screening for Thermochemical Ethanol Synthesis 

United States Department of Energy/National Renewable Energy Laboratory 

Table 3-1 Chemical Absorption Processes Summary Table 

Monoethanolamine 

(MEA) 

Diethanolamine 

(DEA) 

Diglycolamine 

(DGA) 

Diisopropanolamine 

(DIPA) 

Methyldie

Amine MW 61.08 105.14 105.14 133.19 Solvent Capabilities for Gas Treating Minimum Acid Gas in Treated Gas, ppmv  H2S[ ] 3 CO2[ , ] 3 4

1 

15 

4 

200 

1-4 

 15 

10

30%-90% removal 

Bulk 

Selective H2S Removal[ ] 5 No No No Some Removal of COS, CS2, and 

Mercaptans[ , ] 4 5 Partial (Note 1) Partial Partial COS only S

Solution Subject to Degradation[ , ] 4 5(Degrading Species) 

Yes 

(CO2, COS, CS2) 

Some 

(CO2, COS, CS2) 

Yes 

(CO2, COS, CS2) 

Some 

(CO2, CS2) 

Approximate Operating Parameters Solvent Concentration in H20, wt% 15-25[ ] 5 25-35[ ] 5 40-60[ ] 2 20-40[ ] 3 Acid Gas Pickup,  SCF/gal @ 100°F  Moles/mol amine 

3.1-4.3[ ] 5

0.33-0.40[ ] 5

3.8-5.0[ ] 5

0.35-0.65[ ] 5

4.7-6.6[ ] 5

0.25-0.30[ ] 5

2.4-8.1[3] 

0.38-1.1 (Note 2) 

Average Heats of Reaction, Btu/lb  H2S[ ] 2 CO2[ ] 2

615 825 

510 700 

675 820 

475 720 

Typical Utility Demands  LP Steam, lb/gal[ ] 3 0.8-1.5 0.7-1.1 1.5 0.7-0.8 Typical Solvent Losses,  lb/MMSCF Sweet Gas[ ] 3 2-4 1-2 2-4 2-4 Notes: (1) Primary amines are the most reactive with COS and CS2 while tertiary amines are the least reactive. Comparatively

would be the least with tertiary amines. (2) Moles AG/mol amine values for DIPA are calculated from the SCF/gal acid gas pickup data using a 30 wt% solution. pickup rate is used for the preliminary circulation rate estimation in Section 3.2. The acid gas pickup rate at the uapplicable for applications with high H2S partial pressures. For CO2 removal applications, the formation of carbamate lim

thus, the acid gas pickup rate is limited to 0.5 mol CO2/mol amine. 
Task 1: Acid Gas Removal Technology Survey and Screening for Thermochemical Ethanol Synthesis 

United States Department of Energy/National Renewable Energy Laboratory 

Table 3-1 (Continued) 

Monoethanolamine 

(MEA) 

Diethanolamine 

(DEA) 

Diglycolamine 

(DGA) 

Diisopropanolamine 

(DIPA) 

Methyldie

Commercial Experience  With Synthesis Gas  With Natural and Refinery Gas Yes Yes Yes 

Yes Yes 

Process Advantages Often the preferred 

solvent for treating gases containing only H2S and CO2 in low 

concentrations 

(1) Partial COS removal attainable without significant solution degradation; (2) Suitable for lowpressure operations due to low solvent vapor pressure 

(1) Partial COS removal; (2) Lower solvent circulation rate and steam consumptioncompared to MEA

(1) Partial COS removal attainable without significant 

solution degradation; (2) Low regeneration steam required 

(1) Highly s(2) Low heasolution, resregeneratio(3) Solvent thermal and stability and noncorrosiv

Process Disadvantages (1) Not suitable for 

treating gases with high COS and CS2

concentrations due to formation of nonregenerable products, resulting in appreciable solution losses; (2) High heats of solution, requiring high regeneration energy; (3) Solution more corrosive compared to other amine solutions 

Vacuum distillation may be necessary for reclaiming of contaminated solutions[ ] 2

High heats of solution, requiring high regeneration energy 

Not widely used in U.S. 
 

Task 1: Acid Gas Removal Technology Survey and Screening for Thermochemical Ethanol Synthesis 

United States Department of Energy/National Renewable Energy Laboratory 

21 

Table 3-2 Estimated Amine Circulation Rates and Regeneration Energy for NREL Design 

Solvent Type MEA DEA DGA DIPA MDEA

Acid Gas Pickup, 

mol AG/mol amine 0.37 0.50 0.28 

 

0.50 0.38 

Solution Concentration, wt% 20% 30% 50% 30% 50%

Circulation Rate, gpm 1,971 1,540 1,732 2,127 1,463 

Regeneration Energy, MMBtu/hr 

(based on heats of solution only) 43 37 43 

 

38 30 
 

Task 1: Acid Gas Removal Technology Survey and Screening for Thermochemical Ethanol Synthesis 

United States Department of Energy/National Renewable Energy Laboratory 

22 

Section 4 Physical Absorption Processes

4.1 PROCESS OVERVIEW 

Physical absorption processes use an organic solvent to absorb the acid gases from raw gas 

streams without chemical reaction. Each component in the raw gas stream is absorbed in 

proportion to its concentration and solubility. The driving force of this process is the high 

solubility of acid gases in the organic solvent. In most cases, solubility increases with 

increasing pressure and decreasing temperature. Thus physical absorption process is most 

efficient and economical when operated at the highest pressure and lowest temperature 

possible. 

Physical absorption has three advantages compared to chemical absorption. First, unlike most 

chemical absorption processes in which selective recovery may be difficult due to the 

simultaneous reactions of CO2 and H2S with the solvent, selective H2S removal can be achieved 

with physical absorption processes since H2S has a higher solubility than CO2 in most organic 

solvents. Selective H2S removal is necessary when the H2S/CO2 ratio in the feed gas stream is 

too low—a result when the raw gas has little H2S and a large amount CO2—to produce an acid 

gas stream with a high enough H2S content to be processed in the conventional Claus sulfur 

recovery unit. The H2S/CO2 ratio can be increased by selectively removing the H2S to 

concentrate the sulfur content in the acid gas stream. 

Second, minor gas impurities such as COS and mercaptans are soluble in most organic 

solvents, making removal of such impurities in the acid gas removal unit possible without 

subjecting the solvent to degradation. Due to the chemically inert nature of physical solvents, 

solvent deterioration from irreversible reactions with gas impurities is not an operating concern. 

Third, the organic solvents can be regenerated by pressure reduction, inert gas stripping, or 

thermal regeneration. Pressure reduction and inert gas stripping require little use of energy 

while thermal regeneration requires significantly less energy compared to that of chemical 

absorption processes due to the lower heat of desorption of acid gases. When stringent H2S 

purity is required in the treated gas, thermal regeneration is used to achieve thorough stripping 

of the rich solvent.

However, the benefits of the physical absorption processes may be offset by several of their 

disadvantages. These processes usually require high operating pressures, high acid gas partial 

pressures, and low operating temperatures to obtain economic solvent loading since the solvent 

loading capacity improves at higher acid gas partial pressures and lower operating 

temperatures. As a result, the capital and operating costs of the physical absorption processes 

are generally higher compared to those of the amine processes, contributed by the highpressure equipment and the additional heat exchange equipment and refrigeration required to 

achieve reduced operating temperatures. 
 

Task 1: Acid Gas Removal Technology Survey and Screening for Thermochemical Ethanol Synthesis 

United States Department of Energy/National Renewable Energy Laboratory 

23 

Additionally, hydrocarbons are soluble in most organic solvents to some degree, with heavier 

hydrocarbons more soluble than lighter ones. For this reason, physical absorption processes 

are not economical when there are high concentrations of hydrocarbons in the raw gas due to 

appreciable product losses. 

Below are brief descriptions of the two commonly used physical solvent-based AGR processes: 

Selexol and Rectisol. Their characteristics and performance are summarized in Table 4-1. 

Selexol 

The Selexol process has been used for acid gas removal in a wide range of applications 

including natural gas processing, ammonia and hydrogen production, and petroleum and coal 

gasification[ ] 2 . The solvent used in the Selexol process is a mixture of a dimethylether of 

polyethylene glycol. The Selexol solvent has high capacity for acid gases, favorable solubility 

for acid gases versus other light gases, and high selectivity for H2S over CO2. It also has low 

vapor pressure, which minimizes solvent losses, and is chemically and thermally stable, which 

eliminates the need for reclaiming or purging. Due to the solvent’s inert chemical 

characteristics, the Selexol process uses carbon steel as the standard material of construction, 

except for areas that are exposed to CO2/H2S evolution where stainless steel is used. 

The Selexol process can be configured in a number of ways, depending on the levels of acid 

gas removal required, from selective H2S or bulk CO2 to trace acid gas removal. For selective 

H2S removal, refrigeration is normally not used, but it is necessary when CO2 removal is also 

required. Figure 4-1 illustrates a Selexol process for the removal of H2S and CO2. The feed 

gas is contacted countercurrently with a lean solvent in an absorber at high pressure and lower 

temperatures so that H2S and CO2 are absorbed into the solvent. The rich solvent is 

regenerated by pressure reduction through a series of flash drums, followed by thermal 

regeneration with steam stripping. The high-pressure flash gas vapors, consisting mostly of 

hydrogen and CO, are compressed and returned to the absorber. The low-pressure flash gas 

vapor is combined with the acid gas stream from the stripper overhead. The combined stream 

is sent to the sulfur recovery unit for processing. The lean solution is either cooled or chilled, 

depending on the level of CO2 removal, and returned to the top of the absorber. 

Figure 4-2 shows the basic flow configuration of a selective Selexol process, where removal of 

H2S and CO2 can be accomplished via two separate absorbers. 
 

Table 4-1 Selexol Solvent Relative Solubility of Gases 

Compounds Solubility Ratio

H2 (Least Soluble) 1.0 

N2 1.5 

CO 2.2 

CH4 5 

CO2 75 

COS 175 

H2S 670 

CH3SH 1,700 

SO2 7,000 

H2O 55,000 

HCN (Most Soluble) 95,000 

Source: Reference [6] 

Hydraulic Turbine

High-Pressure

Flash Drum

Low-Pressure 

Flash Drum

Rich Solvent

Absorber

Recycle Compressor

Lean Solvent

Stripper

Lean Solution 

Cooler/Chiller

Makeup

Water

Feed Gas

Clean Gas

Acid Gas to 

Sulfur Recovery

Lean Solution 

Pump

Reboiler

Steam

Acid Gas 

Condenser

Rich/Lean

Exchanger

Source: Adapted from reference [6] 

Figure 4-1 Typical Selexol Process Flow Diagram 

Task 1: Acid Gas Removal Technology Survey and Screening for Thermochemical Ethanol Synthesis 

United States Department of Energy/National Renewable Energy Laboratory 

24 
 

Source: reference [15] 

Figure 4-2 Typical Selective Selexol Process Flow Scheme 

Rectisol 

The Rectisol process, which uses chilled methanol as a solvent, is distinctive from other 

physical absorption processes in that it can remove H2S and CO2 as well as other gas impurities 

such as HCN and organic sulfur compounds to very low levels—with H2S concentration as low 

as 0.1 ppm and CO2 concentration of just a few ppm. Depending on process requirements, the 

Rectisol process can be designed in various configurations to achieve (1) deep, nonselective 

CO2 and H2S removal, (2) selective H2S removal with some degree of CO2 slippage, or (3) 

selective removal and recovery of separate CO2 and H2S products, along with the treated 

product gas stream. 

The Rectisol process operates at very low temperatures, ranging from -40°F to -100°F. Low 

temperatures enhance solvent solubility and reduce solvent losses, which can be significant due 

to the relatively high vapor pressure of methanol. However, operation at low temperatures 

requires additional heat exchange equipment and high refrigeration energy, which contributes to 

high plant costs relative to other processes. For this reason, the Recitsol process is mostly 

used when stringent purity of the treated gas is required. 

4.2 PROCESS APPLICABILITY TO THERMOCHEMICAL ETHANOL SYNTHESIS 

In NREL’s current design, the raw gas enters the AGR at 420 psia has an H2S partial pressure 

of 0.1 psia and a CO2 partial pressure of 48 psia. As shown in Figure 2-1, the total acid gas 

pressure of the feed gas is high enough that physical solvent processes of both Selexol and 

Task 1: Acid Gas Removal Technology Survey and Screening for Thermochemical Ethanol Synthesis 

United States Department of Energy/National Renewable Energy Laboratory 

25 
 

Task 1: Acid Gas Removal Technology Survey and Screening for Thermochemical Ethanol Synthesis 

United States Department of Energy/National Renewable Energy Laboratory 

26 

Rectisol can be considered for the NREL application. For these processes, the issue of 

chemical incompatibility with the impurities in the feed gas, which consist of mostly CO2 and 

H2S, and a small amount of ammonia, is not a concern due to the chemically inert nature of the 

treating solvents. The required specifications of the treated gas are 50 ppmv H2S and 5 mol% 

CO2, which correspond to a required acid gas removal of about 79% H2S and 59% CO2. Given 

that deep H2S and CO2 removal is not required, the Selexol process would be the preferable 

option on the basis of relative plant cost. The capital cost of the Rectisol process is generally 

higher relative to the Selexol process due to the need for more heat exchange equipment to 

cool the absorber feed gas and to recover the cold duty from the product gas. The operating 

costs are also generally higher than the Selexol process as a result of higher power 

requirements due to colder refrigeration temperatures. 

Similar to the MDEA system, preliminary simulation modeling of a Selexol process indicates that 

in order to meet the specified 5 mol% CO2 in the treated syngas, the amount of H2S removed 

may be more than necessary due to the H2S selectivity of the Selexol solvent. Under this 

situation, the H2S content of the treated syngas can be maintained at the required 50 ppmv level 

for ethanol synthesis catalyst activity by injection with a slip stream of dimethyl disulfide. 

Alternatively, the Selexol system can be designed for selective H2S and CO2 removal, which 

should offer a better control of the overall end product treated syngas composition. 

Physical solvents have high affinity for hydrocarbons and CO. The NREL AGR feed gas stream 

contains a high concentration of CO and small quantities of hydrocarbons, ranging from 

methane to pentane, benzene and tar, which would be absorbed to some extent into the treating 

solutions. The loss of CO through absorption, if significant, would negatively affect the yield of 

the alcohol synthesis process and the overall plant efficiency. 

Equally important, the absorption of CO and hydrocarbon gases into the treating solvents has 

an environmental impact on the overall process since these gases, which exit the AGR unit with 

the acid gas stream, would be vented with the CO2 stream from the LO-CAT sulfur recovery unit 

(Figure 1-1, Section 1). The incineration of the CO2 vent gas to destroy the CO and 

hydrocarbons would consume a significant amount of product gas as fuel as well as increase 

the overall plant cost. 

The concentrations of CO (a criteria pollutant) and hydrocarbons (greenhouse gases) in the CO2

vent stream can be minimized by depressuring the rich solution and recycling the flashed gas to 

the absorber. The flash pressure can, however, influence the level of contaminant in the CO2

vent gas. At low flash pressures, less CO and hydrocarbons will be present in the CO2 vented 

stream, but the cost for recycling the flashed offgas to the absorber would increase and can 

negatively impact the overall process economics. In comparison to an MDEA system, a 

physical solvent system is expected to have higher concentrations of CO and hydrocarbons in 

the vented stream. Preliminary process simulation of this effect was preformed and the results 

are shown in Figure 4-3. The allowable CO and hydrocarbons concentrations in the CO2 vent 

stream are site-specific and need to be established prior to the start of Task 2 detailed analysis. 
 

Task 1: Acid Gas Removal Technology Survey and Screening for Thermochemical Ethanol Synthesis 

United States Department of Energy/National Renewable Energy Laboratory 

27 

100

0 50 100 150 200 250 300 350 Decreasing Flash Drum Pressure

Increasing CO &

1000

10000

100000

400 450

 Hydrocarbons Concentrations in Vent Gas

CO in vent gas - Selexol system

CO in vent gas - MDEA System

HCs in vent gas - Selexol system

HCs in vent gas - MDEA system

Low Flash Pressures High Flash Pressures

Note: The graph is intended to 

show the relative emissions of the 

Selexol vs. the MDEA systems, 

thus the units are not provided. 

Actual emissions of each system 

will be determined in Task 2.

Figure 4-3 Order-of-Magnitude Comparison of CO and Hydrocarbons Absorption in a Selexol 

and MDEA Process 

A Selexol plant, in general, is more costly than an equivalent size MDEA unit. Two main factors 

influencing the costs of the Selexol process are the solvent loading rates and the required 

refrigeration. First, the solvent loading rates of physical solvents are generally slower due to the 

higher solvent viscosity at low temperatures; thus, the absorption of acid gases into the treating 

solvent requires longer contact time and the use of more absorber stages. Accordingly, the 

Selexol process requires taller absorber column designs as compared to those of the amine 

systems, resulting in relatively higher capital costs. Second, although the steam requirements 

of the Selexol process would be lower due to the lower regeneration energy, refrigeration is 

required, incurring additional capital and operating costs. Therefore, compared to a MDEA, the 

plant cost of the Selexol process is expected to be higher at the given operating conditions. 

Economics of the Selexol process may improve with increasing operating pressures. In NREL’s 

current design, the raw gas enters the AGR unit at about 420 psia, and the treated gas is further 

compressed to the alcohol synthesis pressure of about 990 psia. If the Selexol process is 

operated at a higher pressure, the acid gas solubility is will be enhanced, which will result in 

reduced solvent circulation rate, smaller absorber column size, and lower refrigeration load. At 

the higher operating pressure, however, the absorption of product gases will also increase, 

which would impact the CO and hydrocarbons emissions from the LO-CAT unit as well as the 

overall plant efficiency. The techno-economics analysis for the AGR process selection must 

also take this compression effect into consideration. 
Task 1: Acid Gas Removal Technology Survey and Screening for Thermochemical Ethanol Synthesis 

United States Department of Energy/National Renewable Energy Laboratory 

Selexol Rect

Solvent Capabilities for Gas Treating 

Table 4-2 Physical Absorption Processes Summary Table 

Minimum Acid Gas in Treated Gas, ppmv  H2S 

 CO2

1[6] 

300[ ] 3

0

10-10

Selective H2S Removal[ ] 2 Yes YRemoval of COS, CS2, and Mercaptans[ , ] 6 7 Partial YSolution Subject to Degradation[ , ] 4 8 No Not RepoApproximate Operating Parameters Active Ingredient Dimethylether of polyethelene glycol MethaAcid Gas Solubility,  CO2 Vol/Vol Solvent @ 1 atm[ ] 2 H2S Vol/Vol Solvent @ 1 atm[2] 

3.63 32.4 

8

41

Typical Solvent Losses, lb/MMSCF Sweet Gas[ ] 3

0.5 

Commercial Experience with Gasification (Coal & Petcoke) Yes YProcess Advantages Process used in a wide range of 

industries 

Able to handle difficult treating conditions where other processes 

may not be suitable. 

Process Disadvantages Hydrocarbon losses can be 

significant when treating hydrocarbon-rich gas streams 

(1) Low levels refrigeration require(2) Hydrocarbons heavier than C3+ 

are absorbed with CO2 and H2S a

are difficult to recover[ ] 7
 

Task 1: Acid Gas Removal Technology Survey and Screening for Thermochemical Ethanol Synthesis 

United States Department of Energy/National Renewable Energy Laboratory 

29 

Section 5 Hybrid Processes

5.1 PROCESS OVERVIEW 

The concept of the hybrid process is to combine a chemical solvent (typically an amine) and a 

physical solvent in order to take advantage of the effects of both processes. The presence of a 

physical solvent enhances solution capacity at high acid gas partial pressures. However, this 

advantage is not observed at low acid gas partial pressures. The hybrid process can achieve 

very low acid gas specifications in a single treating step through a combination of bulk acid gas 

removal by the physical solvent and stringent purification by the chemical solvent. 

A common hybrid process is the Shell’s Sulfinol process, which is often used to treat refinery, 

natural, and synthesis gases The Sulfinol solvent is a mixture of tetrahydrothiophene dioxide 

(Sulfolane), an alkanolamine (either DIPA or MDEA), and water. The solution with DIPA, 

referred to as Sulfinol-D, is usually selected when complete removal of H2S and CO2 as well as 

deep removal of COS is required. The Sulfinol-D solvent is used in the majority of the Sulfinol 

plants in operation worldwide. The solution with MDEA, referred to as Sulfinol-M, is usually 

used for selective removal of H2S in the presence of CO2 and is capable of partial removal of 

COS. Both solvents are capable of removing mercaptans to low levels[ ] 2 . 

The Sulfinol process offers several advantages, including: 

 Sulfinol systems are better suited than amine systems to treat gases containing 

mercaptans and COS since the amine systems cannot efficiently remove these 

compounds 

 Sulfinol solvents are reported to have low solubility for hydrocarbons up to pentane. 

This is an advantage over the physical absorption systems when treating gases rich in 

hydrocarbons since hydrocarbon losses can be minimized 

 Sulfinol solvents are stable. Solution degradation and losses are minimal 

 Corrosion problems are not prevalent in Sulfinol plants 

 High solution capacity at high acid gas partial pressures 

 

Some disadvantages of the Sulfinol process are: 

 At CO2 acid gas pressure below 100 psia, the solvent loading (scf gas/gal solvent) of 

Sulfinol is worse than 15 wt% MEA[ ] 13 . 

 At low H2S partial pressure, the Sulfinol solution capacity is inferior to aqueous MEA[ ] 2 , 

but solution capacity improves as the H2S partial pressure increases. A pronounced 

solution capacity enhancement is observed at H2S partial pressure above 40 psia. 

Therefore, the Sulfinol process requires high operating pressures and high acid gas 

partial pressures to obtain the benefit of high solution capacity 

 Sulfinol systems may incur royalty cost and higher solvent costs than amine systems. 
 

nd Screening for Thermochemical Ethanol Synthesis 

United States Department of Energy/National Renewable Energy Laboratory 

Task 1: Acid Gas Removal Technology Survey a 30 

5.1.1 Flow Scheme 

The Sulfinol process flow scheme, shown in Figure 5-1, is similar to that of a conventional 

amine system with the inclusion of the flash tank to recover the absorbed hydrocarbons. The 

recovered hydrocarbons can either be recycled to the absorber or used as plant fuel. 

Flash Drum

Rich Solvent

Absorber

Lean Solvent

Stripper

Lean Solution 

Cooler

Feed Gas

Clean Gas Acid Gas to 

Sulfur Recovery

Lean Solution 

Pump

Reboiler

Steam

Rich/Lean

Exchanger

Fuel

Knock-out Drum

Acid Gas

Condenser

Reclaimner

Optional Equipment

Figure 5-1 Typical Sulfinol Process Flow Diagram 

5.2 PROCESS APPLICABILITY TO THERMOCHEMICAL ETHANOL SYNTHESIS 

General guidelines, Figures 2-2 to 2-3, suggest that Sulfinol is best suited for H2S removal, but 

not CO2 (Figure 2-4). This, coupled with the reported findings of Edwards[13] that the solvent 

loading of Sulfinol is worse than 15 wt% MEA at CO2 partial pressure below 100 psi, would 

seem to suggest that there is no advantage of using a Sulfinol process for the NREL ethanol 

synthesis application. 

Hybrid system of mixed solvents tends to be highly proprietary, with its formulation heavily 

guarded by the technology developer/vendors. It is hard to carry out any design and economic 

assessment/evaluation without proper inputs from them. Furthermore, hybrid system may incur 

royalty cost and higher solvent costs than the amine systems. 

 

 
Task 1: Acid Gas Removal Technology Survey and Screening for Thermochemical Ethanol Synthesis 

United States Department of Energy/National Renewable Energy Laboratory 

Sulfinol-D Sulfino

Solvent Capabilities for Gas Treating 

Table 5-1 Hybrid Processes Summary Table

Minimum Acid Gas in Treated Gas, ppmv  H2S 

 CO2

0.5-2[3] 

15-3000[ ] 3

0

Selective H2S Removal[ ] 2 Moderate YRemoval of COS, CS2, and Mercaptans[ ] 2 Yes YSolution Subject to Degradation[ ] 2 Minimal MiniApproximate Operating Parameters Active Ingredient Sulfolane, DIPA Sulfolane, MDAcid Gas Pickup, SCF/gal 4-16[3] Typical Solvent Losses, lb/MMSCF Sweet Gas 

0.5-1.6[3] 

Typical Utility Demands  LP Steam, lb/ MSCF AG Removed 0.5-1.2 Commercial Experience  With Coal Gasification  With Other Gasification 

Yes 

Process Advantages (1) Capability for mercaptans and COS 

removal (2) Lower hydrocarbon solubility compared to physical absorption processes (3) Minimal corrosion and solution degradation (4) Solution capacity enhanced at high acid gas partial pressure 

Same as Sulfinol-D 

Process Disadvantages (1) Solution capacity inferior to amine at low 

H2S acid gas partial pressure 

(2) May incur royalty cost and higher solvent costs than amine system 

Same as Sulfinol-D 
 

Task 1: Acid Gas Removal Technology Survey and Screening for Thermochemical Ethanol Synthesis 

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32 

Section 6 Hot Carbonate-Based Process

6.1 PROCESS OVERVIEW 

The hot potassium carbonate process—a chemical absorption process similar to the aminebased systems—has been used primarily for CO2 removal from synthesis gas in ammonia and 

hydrogen plants, and also for acid gas removal in natural gas plants. The process uses an 

aqueous solution of potassium carbonate (K2CO3) to absorb CO2 and H2S. The rate of CO2

absorption can be enhanced by increasing the CO2 concentration, the solution alkalinity, or 

temperature[ ] 2 . Organic sulfur compounds (COS, CS2, and mercaptans) can also be removed to 

some extent. The potassium carbonate solution reacts with the acid gases as followed: 

H2S + K2CO3 ÅÆ KHCO3 + KHS 

CO2 + K2CO3 + H2O ÅÆ 2KHCO3

The three basic flow schemes for the potassium carbonate process are the single-stage 

process, the split-flow process, and the two-stage process, as shown in Figures 6-1, 6-2, and 6-

3, respectively. Selection of a suitable configuration depends on the required purity of the 

treated gas. In the simplest configuration, illustrated by the single-stage process, the acid gas is 

absorbed in a countercurrent absorber by a lean potassium carbonate solution. The rich 

solution is regenerated by flashing, which releases approximately one third to two thirds of the 

absorbed CO2

[ ] 4 , and steam stripping in a low-pressure stripper. The stripping steam 

requirement is reduced as a result of an appreciable solvent regeneration by pressure 

reduction. The absorber and stripper normally operate at approximately the same temperature 

of 230-240°F[ ] 5 . For this reason, the heat exchange operations between the lean and rich 

streams are eliminated. 

In the split-flow process, the lean solution from the stripper is split: a major portion of the hot 

solution is fed to the middle section of the absorber for bulk removal while the remainder is 

cooled and fed at the top of the absorber for trim acid gas removal. With this process 

modification or enhancement, the two-stage process can achieve more complete CO2 removal. 

In this configuration, a major portion of the solution is withdrawn from the stripper at a point 

above the reboiler. Only a small portion of the solution passes down through the bottom of the 

stripper to the reboiler. Since this small portion of the solution is regenerated by the total steam 

supply to the stripper, it is thoroughly stripped and thus capable of removing the CO2 to low 

levels. The thoroughly stripped solution is then cooled and fed to the top of the absorber. 

The hot carbonate process offers several advantages, including: 

 Lower regeneration energy requirements compared to that of the amine systems due to 

the low heat of solution for CO2 absorption in potassium carbonate 

 Less hydrocarbon absorption compared to the physical absorption processes due to low 

solubility of hydrocarbons in the potassium carbonate solution 
 

The disadvantages of the process are: 

 Corrosion is a problem for some portions of the plant, requiring the use of stainless steel 

 At high temperatures, potassium carbonate reacts with CO to form potassium formate, 

which requires costly solvent purge and makeup. This problem can be controlled by 

lowering the absorber temperature, but at a cost of reducing the CO2 absorption rate, 

which in turn, would lead to higher solution rate and regeneration energy. 

Source: Adapted from Engineering Data Book (1987) 

Figure 6-1 Typical Hot Potassium Carbonate Single-Stage Process 

Task 1: Acid Gas Removal Technology Survey and Screening for Thermochemical Ethanol Synthesis 

United States Department of Energy/National Renewable Energy Laboratory 

33 
 

Treated Gas

Rich Solution

Absorber

Cooled 

Lean Solution

Stripper

Lean Solution 

Cooler

Feed Gas

Acid Gas to 

Sulfur Recovery

Lean Solution 

Pump

Reboiler

Steam

Acid Gas 

Condenser

Hot

Lean Solution

Figure 6-2 Typical Hot Potassium Carbonate Split-Flow Process 

Treated Gas

Rich Solution

Absorber

Cooled 

Lean Solution

Stripper

Lean Solution 

Cooler

Feed Gas

Acid Gas to 

Sulfur Recovery

Lean Solution 

Pump

Reboiler

Steam

Acid Gas 

Condenser

Hot Semi-Lean Solution

Source: Adapted from Engineering Data Book (1987) 

Figure 6-3 Typical Hot Potassium Carbonate Two-Stage Process 

Task 1: Acid Gas Removal Technology Survey and Screening for Thermochemical Ethanol Synthesis 

United States Department of Energy/National Renewable Energy Laboratory 

34 
 

Task 1: Acid Gas Removal Technology Survey and Screening for Thermochemical Ethanol Synthesis 

United States Department of Energy/National Renewable Energy Laboratory 

35 

Table 6-1 summarizes the performance and characteristic of the two popular hot carbonate 

processes: the Benfield process, licensed by UOP, and Catacarb process, licensed by 

Eickmeyer and Associates. The Benfield process uses an activated, inhibited hot potassium 

carbonate solution. Activators, either DEA or an organic activator known as ACT-1, are added 

to the solution to accelerate the rate of CO2 absorption while inhibitors are used to minimize 

corrosion. The use of activators results in significant process improvements in terms of better 

product quality and lower capital and operating costs. The Benfield process has been widely 

used for H2S and CO2 removal from natural and ammonia synthesis gas. The process has a 

variety of configurations that permit process optimization and energy reduction. 

Similar to the Benfield process, the Catacarb process also uses activators to catalyze the CO2

absorption rate to improve product quality and lower plant costs. The two types of activators 

used in this process are organic or inorganic activators. Organic activators consist of amine 

promoters to improve acid gas absorption rate. Inorganic activators are used in oxidizing 

environments where the presence of oxygen in the feed gas may deteriorate organic 

components. Inhibitors are used to mitigate corrosion problems. The Catacarb process has 

primarily been used for CO2 removal from synthesis gas, but it is also applicable for H2S 

removal from natural gas. The process is available in a number of configurations, designed to 

provide high purity gas and greater heat economy. 

6.2 PROCESS APPLICABILITY TO THERMOCHEMICAL ETHANOL SYNTHESIS 

Hot potassium carbonate process has potential as an AGR candidate for NREL’s process. The 

design syngas specification has significant amount of CO, which can cause design problems. 

At high concentration (design CO partial pressure of ~180 psia) CO has the potential to react 

with the potassium carbonate solution to form potassium formate at high temperatures, which 

would require costly solvent purge and makeup. Typical applications of the hot potassium 

carbonate processes in the hydrogen or ammonia plants normally have a CO partial pressure of 

less than 10 psia in the gas streams (after the CO shift step). Since the CO partial pressure of 

the syngas in the NREL design is considerably higher than this typical range, the hot potassium 

carbonate processes are not recommended, without addressing this issue in more detail with 

the technology vendors. 

Furthermore, despite its seemingly simple chemistry, the hot carbonate-based AGR process is 

actually quite complicated and difficult to model with commercial process simulators. There are 

additional side reactors within the aqueous media, which are not well defined. The acid gas 

reactions involve many other ionic species. Unless there is a set of well defined and 

documented experiment and/or plant data for benchmarking, it is not certain to what extent the 

process performance can be truly captured. Any application design would most likely have to be 

involved directly with the technology vendors. 

 
Task 1: Acid Gas Removal Technology Survey and Screening for Thermochemical Ethanol Synthesis 

United States Department of Energy/National Renewable Energy Laboratory 

Benfield Catac

Solvent Capabilities for Gas Treating 

Table 6-1 Hot Potassium Carbonate Processes Summary Table

Minimum Acid Gas in Treated Gas, ppmv  H2S 

 CO2

4[3] 

500[ ] 3

50

Selective H2S Removal No Removal of COS, CS2, and Mercaptans[ ] 5 Partial PaSolution Subject to Degradation[ ] 5 Minimal MiniApproximate Operating Parameters Active Ingredient K2CO3 + Activator + Inhibitor K2CO3 + Activator + InhibAcid Gas Pickup, SCF/galTypical Utility Demands  LP Steam, lb/ MSCF AG Removed 90-160[ ] 3 35-5Commercial Experience  With Coal Gasification  With Synthesis Gas from Catalytic Steam Reforming Yes YProcess Advantages (1) Can achieve high purity 

gases (2) Lower regeneration energy than amines 

(1) Can achieve high purity gases (2) Lower regeneration enerthan amines 

Process Disadvantages Corrosion problems Corrosion problems 
 

Task 1: Acid Gas Removal Technology Survey and Screening for Thermochemical Ethanol Synthesis 

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37 

 Section 7 RTI Solid Absorbent-Based Process

7.1 PROCESS OVERVIEW 

The RTI warm gas clean up (WGCU) process uses ZnO sorbent to remove H2S from gas 

streams. The principal application of the process is for bulk removal of H2S; CO2 is not removed 

by the process. Presently, there are no significant commercial applications of this process. 

RTI’s WGCU process, being developed for high-temperature IGCC application, consists of two 

major system components: High Temperature Desulfurization process (HTDS) and Direct Sulfur 

Recovery process (DSRP). Figure 7-1 shows a simplified WGCU process scheme, located at 

the downstream of the coal gasifier. The HTDS process, consisting of a pair of fluidized bed 

reactors (identified as WGCU absorber & WGCU regenerator in Figure 7-1), is shown in Figure 

7-2 in more detail. In the HTDS process, the hot raw syngas from the gasifier radiant boiler is 

first cooled to the operating temperature of the RTI WGCU absorber at approximately 800o

F, via 

a fire tube convection boiler. In the process, high-pressure, saturated steam is generated. The 

raw syngas, mixed with a tail gas recycled stream, is then sent to a cyclone for bulk ash and 

char removal. The captured solids are recycled to the gasifier. The syngas leaving the cyclone 

is routed to the WGCU absorber where it is contacted with the circulating sorbent to remove H2S 

and COS. The following reactions are believed to take place in the absorber: 

H2S + ZnO = ZnS + H2O 

COS + ZnO = ZnS + CO2

The regenerated sorbent from the regenerator, along with the recycled sorbent from the 

absorber stand pipe (Figure 7-2) enter the absorber near the bottom of the unit. Most of the 

sulfur absorption takes place in the absorber. The mixture leaves the top of the absorber into a 

cyclone where the solid sorbent containing ZnS is separated from the sulfur free syngas. Part 

of the solids is recycled to the absorber via a standpipe. A diverter valve located in the absorber 

standpipe takes a slip stream of these solids and feeds it to the regenerator. 

Within the regenerator, the ZnS containing sorbent comes into contact with a mixture of oxygen 

and nitrogen at a pre-determined ratio. The oxygen then reacts with the ZnS and forms SO2

according to the following reactions: 

ZnS + 1.5 O2 = ZnO + SO2

The above reaction is exothermic, raising the temperature of the resulting mixture to about 

1300o

F. A cyclone is used to separate the solids and recycled it to the absorber. The 

regenerator offgas containing SO2 is heat exchanged with the incoming oxygen/nitrogen mixture 

before sending onto a companion DSRP (Direct Sulfur Recovery Process) unit for sulfur 

removal. 
 

Task 1: Acid Gas Removal Technology Survey and Screening for Thermochemical Ethanol Synthesis 

United States Department of Energy/National Renewable Energy Laboratory 

38 

SO2 and N2 from HTDS regeneration, after heat exchanged cooling and filtration to remove 

entrained solids, are sent onto the fixed bed catalytic DSRP reactor where SO2 is reduced to 

elemental sulfur according to the following reactions: 

SO2 + 2 CO = 2 CO2 + S and 

SO2 + 2 H2 = 2 H2O + S 

CO and H2 are provided by a slip stream from the HTDS absorber. Approximately 98% of the 

SO2 is converted to elemental sulfur in the DSRP reactor. The reaction is exothermic and it 

raises the reaction outlet mixture temperature to about 1,200o

F. The product stream from the 

DSRP reactor is sent onto a two stage sulfur condenser unit where the elemental sulfur is 

condensed and separated. Heat is recovered by making low pressure steam. 

The tail gas containing CO2, N2 and steam is directed to 2nd stage DSRP (Hydrogenation) 

reactor where the residual SO2 is hydrogenated to H2S in according with the following reduction 

reaction: 

SO2 + 3 H2 = H2S + 2 H2O 

The hydrogenated stream is then cooled in two stages. In the first stage, the gas is cooled by 

generating low pressure steam. In the second stage the gas is further cooled to 315o

F by heat 

exchange with boiler feed water (BFW). The cooled stream is then compressed and recycled 

as feed to the HTDS reactor. 

Companion processes are included in the overall WGCU design, to remove mercury, chloride 

and ammonia to meet environmental emissions requirement, as follows: 

 A high temperature duel candle filter system to remove entrained solids from the 

desulfurized syngas stream leaving the WGCU process before sending the stream onto 

the fixed-bed mercury removal vessel, 

 High temperature (~ 550 oF) fixed-bed RTI proprietary mercury removal process, 

 A companion high temperature (~ 550 oF) fixed-bed chloride guard bed for HCl removal, 

based on reaction with sodium bicarbonate in according with the following reaction – 

HCl + NaHCO3 = NaCl + CO2 + H2O 

 Inclusion of an SCR (Selective Catalytic Reduction) onto the power train for NOx 

emissions control; thus no online ammonia removal is provided with the WGCU design. 
 

Figure 7-1 Simplified Overall WGCU Process Scheme

Figure 7-2 High Temperature Desulfurization (HTDS) Process Flow Scheme 

Task 1: Acid Gas Removal Technology Survey and Screening for Thermochemical Ethanol Synthesis 

United States Department of Energy/National Renewable Energy Laboratory 

39 
 

Task 1: Acid Gas Removal Technology Survey and Screening for Thermochemical Ethanol Synthesis 

United States Department of Energy/National Renewable Energy Laboratory 

40 

7.2 PROCESS APPLICABILITY TO THERMOCHEMICAL ETHANOL SYNTHESIS 

The WGCU process is being developed for coal gasification IGCC application to increase 

overall plant thermal efficiency. Principal application is for bulk sulfur removal and the process 

does not yet have the capability to remove CO2. RTI is currently also developing a companion 

high temperature solid absorbent-based CO2 removal process. At the current state of 

development, this process will not be able to meet the AGR requirement of the NREL design by 

itself. 

 
 

Task 1: Acid Gas Removal Technology Survey and Screening for Thermochemical Ethanol Synthesis 

United States Department of Energy/National Renewable Energy Laboratory 

41 

 Section 8 Liquid Phase Oxidation Processes

8.1 PROCESS OVERVIEW 

In the liquid phase oxidation processes, H2S is absorbed and oxidized to elemental sulfur by 

oxygen carriers dissolved or suspended in a solution. The processes are generally used for 

H2S removal from gases containing relatively low concentrations of H2S in the presence of high 

concentrations of CO2. A high ratio of CO2 to H2S in the raw gas to a typical acid gas removal 

unit, such as an amine unit, can produce an acid gas stream with a H2S concentration that is too 

low for processing in a conventional Claus sulfur recovery plant. In these cases, the liquid 

phase oxidation processes can be employed for sulfur removal and recovery. Generally, these 

processes are suitable for small-scale applications that require less than 20 tons per day of 

sulfur recovery capacity. 

Liquid phase oxidation processes have several drawbacks. First, the process only removes H2S 

and not CO2, thus by itself, it cannot serve as an AGR process for the thermochemical ethanol 

production. Second, the low solution capacities for H2S and oxygen usually require large liquid 

circulation rates and large facilities to handle the precipitated sulfur. Third, the precipitated 

sulfur is difficult to separate from the solution. Finally, the heat generated by H2S oxidation 

cannot be recovered for steam generation. 

Popular liquid phase oxidation processes are the iron-chelate systems due to the non-toxic 

nature of the treating solutions. A common iron-chelate process presently used in commercial 

applications is the LO-CAT process. Below is a brief description of the LO-CAT process. 

LO-CAT Process 

The LO-CAT process uses iron catalyst held in a chelating agent to oxidize H2S to elemental 

sulfur. The LO-CAT catalyst solution is environmentally safe and produces no hazardous waste 

byproducts. The process can be designed to achieve up to 99.9+% H2S removal efficiency. 

The LO-CAT process has been used for H2S removal from amine acid gas streams and 

synthesis gas from municipal solid waste gasification[ ] 10 . 

Figure 8-1 illustrates a conventional LO-CAT unit. The process operates at ambient 

temperature and requires no heating or cooling of the solution[ ] 2 . In the absorber, the H2S in the 

gas stream is converted to elemental sulfur as it comes in contact with the LO-CAT solution. 

The spent catalyst and the elemental sulfur from the absorber are fed into the oxidizer where 

the spent catalyst is regenerated by contact with oxygen in air and the elemental sulfur is 

concentrated into a sulfur slurry. The sulfur slurry is then processed in the sulfur handling unit 

to recover the sulfur and any entrained catalyst. The recovered sulfur is considered low-value 

due to a small amount of entrained residual catalyst. 

The capital cost of the LO-CAT system is largely determined by the solution circulation rate, 

which is directly proportional to the gas feed rate and its H2S concentration and is inversely 
 

Task 1: Acid Gas Removal Technology Survey and Screening for Thermochemical Ethanol Synthesis 

United States Department of Energy/National Renewable Energy Laboratory 

42 

proportional to the iron content in the solution while the operating costs are determined by the 

power and catalyst make-up requirements. 

Source: Reference [11] 

Figure 8-1 Conventional LO-CAT System 

8.2 PROCESS APPLICABILITY TO THERMOCHEMICAL ETHANOL SYNTHESIS 

The LO-CAT process is not a suitable option for the NREL acid gas removal application by itself 

because the process does not have the capability for CO2 removal. However, the process is 

ideally suited for sulfur recovery as it is used in the current NREL design. As shown on Table 2-

1 in Section 2, the feed gas to the amine unit has very little H2S and a substantial amount of 

CO2. As a result, the acid gas stream has a fairly low concentration of H2S and cannot be 

processed in a conventional Claus sulfur recovery unit. Therefore, the LO-CAT system is used 

for sulfur recovery. 

 
Task 1: Acid Gas Removal Technology Survey and Screening for Thermochemical Ethanol Synthesis 

United States Department of Energy/National Renewable Energy Laboratory 

LO-CAT 

Solvent Capabilities for Gas Treating Minimum Acid Gas in Treated Gas, ppmv  H2S 

 CO2

10 

Not removed 

Removal of COS, CS2, and Mercaptans Not removed Solution Subject to Degradation Degrades by HCN[ ] 3Approximate Operating Parameters Active Ingredient Chelated Iron 

Table 8-1 Liquid Phase Oxidation Processes Summary Table

Feed Stream Acid Gas and Concentration Range, mol% 1.5-100[ ] 3Typical Utility Demands  Electricity, kWh/ MSCF AG Removed 234[ ] 3Method of Regeneration Air oxidation Commercial Experience  With Coal Gasification  With Other Gasification 

Plant under construction 

Yes 

Process Advantages (1) High sulfur removal efficiency 

(2) Non-toxic catalyst solution 

Process Disadvantages High solvent circulation rate 
 

Task 1: Acid Gas Removal Technology Survey and Screening for Thermochemical Ethanol Synthesis 

United States Department of Energy/National Renewable Energy Laboratory 

44 

 Section 9 Membrane Separation Processes

9.1 PROCESS OVERVIEW 

Membrane separation systems are used primarily for bulk CO2 removal in applications that have 

large flows, high CO2 contents, or in remote locations. Membranes have been widely used in 

natural gas sweetening plants to remove CO2 and water vapor and in enhanced oil recovery 

applications, in which CO2 is removed from a natural gas stream and then reinjected into the oil 

well to enhance oil recovery. 

The membrane systems are modular, skid mounted units that are capable of producing upward 

of 250 MMSCFD of treated gas [ ] 12 . The two types of membranes used are the spiral wound or 

hollow fiber membrane elements. The membrane elements are housed in pressure tubes that 

can be mounted in either a horizontal or vertical orientation. Selection of the type of membrane 

elements depends on the applications, and selection recommendations are normally provided 

by process vendors. In general, spiral wound elements, widely used in natural gas sweetening 

industry, can handle higher pressure and are more resistant to fouling. Hollow fiber elements 

have a higher packing density, which, in turn, lead to smaller plant size relative to the spiral 

wound-based plants. 

Figure 9-1 illustrates a simple, single-stage membrane system for CO2 removal, in which the 

feed gas diffuses through the membrane and is then separated into a permeate stream that is 

CO2-rich, and a residual stream that is hydrocarbon-rich. The CO2 flux, defined as the molar 

flow of CO2 through the membrane per unit area of membrane, is dependent on the pressure 

drop across the membrane and on the CO2 permeability. Permeability, in essence, is a 

measure of how quickly a gas can permeate through a membrane. It is a function of the 

solubility of CO2 in the membrane and the diffusion coefficient of CO2 through the membrane. 

Gases with high permeability, or fast gases (i.e. CO2, H2S, H2, and water vapor), can permeate 

quickly through a membrane while gases with lower permeability, or slow gases (i.e. CO, 

methane, ethane, and other hydrocarbons), permeate at a much slower rate. Generally, 

separation of gases with high permeability requires less membrane area and thus, lower plant 

cost. 

Another important variable of the membrane systems is selectivity, which is the ratio of the 

permeability of CO2 to other components in the gas stream. Essentially, it is a measure of how 

much better the membrane permeates CO2 compared to other compounds, such as 

hydrocarbons, in the gas stream. High selectivity of CO2 over hydrocarbons results in limited 

hydrocarbons being permeated and consequently lower hydrocarbon losses. 

Membrane systems offer several advantages, including: 

 Low capital and operating costs for single-stage membrane systems. Multistage 

systems with large recycle compressors (see Figure 9-2) have comparable energy costs 

as traditional acid gas removal technologies. 
 

 Operational simplicity. The one-stage systems have no moving parts. For the two-stage 

systems, the addition of the recycle adds some complexity. 

 Reduced space requirements 

 Ideal for remote locations 

 Can be used to debottleneck existing solvent-based plant. In such cases, the 

membrane is used for bulk acid gas removal, and the existing solvent-based system is 

used for the final cleanup. 

Disadvantages of the membrane systems include the loss of hydrocarbons and hydrogen 

products through permeation, which can be excessive in high CO2 removal applications, and 

they are not suited for applications requiring deep acid gas removal. 

9.1.1 Flow Scheme 

Membrane systems can be configured in several variations, depending on the CO2 removal 

requirements. The simplest flow scheme, a one-stage configuration, is showed in Figure 9-1. 

In high CO2 removal applications, a two-stage system is used to reduce the hydrocarbons 

losses. In this flow scheme, the permeate stream from the first stage, which contains the 

hydrocarbons permeated with the CO2, is recompressed and processed in the second stage. 

The permeate stream is at low pressure and thus requires recompression. The residue stream 

from the second stage, containing the recovered hydrocarbons, is recycled to the feed. 

Feed Gas Membrane Unit

Residue

(CO2 Reduced)

Permeate

(CO2 Enriched)

Source: Reference [12] 

Figure 9-1 One-Stage Membrane System 

Task 1: Acid Gas Removal Technology Survey and Screening for Thermochemical Ethanol Synthesis 

United States Department of Energy/National Renewable Energy Laboratory 

45 
 

Source: Reference [12] 

Figure 9-2 Two-Stage Membrane System 

 

9.2 PROCESS APPLICABILITY TO THERMOCHEMICAL ETHANOL SYNTHESIS 

Membrane separation systems are primarily used for CO2 removal applications and do not have 

the capability to meet the H2S removal requirement of the thermochemical ethanol process. 

Additionally, the raw gas in the NREL design has a high concentration of hydrogen, which has a 

high permeability; thus, the use of membrane for CO2 separation can potentially result in 

substantial hydrogen losses. 

Task 1: Acid Gas Removal Technology Survey and Screening for Thermochemical Ethanol Synthesis 

United States Department of Energy/National Renewable Energy Laboratory 

46 
 

Task 1: Acid Gas Removal Technology Survey and Screening for Thermochemical Ethanol Synthesis 

United States Department of Energy/National Renewable Energy Laboratory 

47 

 Section 10 Conclusions and Recommendations

Using the design basis provided by NREL, the screening criteria outlined in Section 2, and the 

survey performance of each AGR process, the qualitative process screening shows that the 

acid gas removal requirements for the thermochemical ethanol synthesis design can be 

achieved by the chemical and physical absorption processes. All other acid gas removal 

processes have limitations that make them unsuitable for the NREL design. Within the chemical 

absorption processes, the MDEA system appears to be the best option based its low circulation 

rate and regeneration energy requirement, which are the two primary factors that determine the 

capital and operating costs of an AGR process. Preliminary estimates of solution circulation 

rate and solvent regeneration energy requirements indicate that the MDEA system would have 

the lowest plant cost compared to the other amine systems. 

The screening of the physical absorption processes indicates that the Selexol process would be 

more preferable than the Rectisol process since deep H2S and CO2 removal is not required. 

The Rectisol process, usually selected for deep or complete acid gas removal applications, 

generally has a higher plant cost than the Selexol process because of the additional heat 

exchange equipment and the higher power consumption required due to the lower refrigeration 

temperatures. 

The screening analysis also indicates that the selectivity of the MDEA and Selexol solvents for 

H2S in the presence of CO2 may result in a clean syngas stream with an H2S content that is 

below the 50 ppmv level required for synthesis catalyst activity. As discussed in the report, the 

H2S concentration of the clean gas stream could be controlled by injection with dimethyl 

disulfide. The cost impact of such chemical injection is not expected to be significant since only 

a small quantity of dimethyl disulfide is required. 

The screening analysis also indicates that in both the MDEA and Selexol systems, the 

absorption of product gases (CO and hydrocarbons) into the treating solvent has an 

environmental impact on the overall process since these gases would be vented with the CO2

stream from the LO-CAT unit. As discussed in the report, the LO-CAT unit does not include 

CO2 vent gas incineration, which would require a significant amount of product gas as fuel and 

also increase the overall plant cost. Therefore, any CO and hydrocarbons in the acid gas 

stream would be vented with the CO2. The concentrations of CO (a criteria pollutant) and 

hydrocarbons (greenhouse gases) in the vented stream depend on the degree of product gas 

absorption and the level of product recovery achievable by flashing of the rich solution. The 

vented stream from the Selexol system would be expected to have higher concentrations of CO 

and hydrocarbons due to the higher solubility of such compounds in the Selexol solvent. 

Furthermore, in addition to the environmental impact, the loss of CO through absorption, if 

significant, would reduce the yield of the alcohol synthesis process, and thus, impacts the 

overall plant efficiency 

The qualitative cost comparison of the Selexol process with the MDEA process indicates that 

the Selexol process would have a higher plant cost at the given operating conditions and acid 
 

Task 1: Acid Gas Removal Technology Survey and Screening for Thermochemical Ethanol Synthesis 

United States Department of Energy/National Renewable Energy Laboratory 

48 

gas partial pressures. The Selexol process may become more economical at higher pressures 

as a result of enhanced acid gas solubility. In the NREL design, the AGR unit can theoretically 

be operated between the ranges from the tar reformer pressure of 15 psia up to the alcohol 

synthesis pressure of 990 psia. For the Selexol process, which favors high operating pressures, 

a trade-off analysis is required to assess the power and cost penalties associated with the 

compression of syngas to high pressures. In this study, since all AGR options are qualitatively 

evaluated at the fixed operating pressure of 420 psia, the impact of syngas compession is not 

considered. Additionally, the increased solubility of CO and hydrocarbons in the Selexol solvent 

at high pressures must also be evaluated against the advantage of higher solution loading 

capacity resulting from improved acid gas solubility. 

A detailed process design and cost estimation analysis of the MDEA system for thermochemical 

ethanol production is recommended in Task 2. In addition, a comparative study for the Selexol 

system, taking into consideration the various plant design parameters such as the overall 

syngas compression cost and the environmental impact of CO and hydrocarbons emissions 

requirement, etc., is also recommended. 
 

Task 1: Acid Gas Removal Technology Survey and Screening 

United States Department of Energy/National Renewable Energy Laboratory 

49 

 References

1. Phillips, S., Aden, A., Jechura, J., Dayton, D., and Eggeman, T., “Thermochemical 

Ethanol via Indirect Gasification and Mixed Alcohol Synthesis of Lignocellulosic 

Biomass,” NREL Report No. TP-510-41168, April 2007. 

2. Kohl, A. L., and Nielsen, R. B., Gas Purification, 5th ed., Gulf Publishing, Houston, TX, 

1997. 

3. Nexant Database. 

4. Kidnay, A. J., and Parrish, W. R., Fundamentals of Natural Gas Processing, Taylor & 

Francis Group, Boca Raton, FL, 2006. 

5. Engineering Data Book, 10th ed., Section 21, Hydrocarbon Treating, Gas Processors 

Supply Association, 1987. 

6. Kubek, D. J., Polla, E., and Wilcher, F. P., “Purification and Recovery for Gasification,” 

Gasification Technologies Conference, October 1996, San Francisco, CA. 

7. Ranke, G., “Advantages of the Rectisol-Wash Process in Selective H2S Removal from 

Gas Mixtures,” Linde Reports on Science and Technology, Vol. 18, 1973. 

8. UOP Gas Treating information, available at http://uop.com/gasprocessing/6010.html 

9. Catalyst information from Johnson Matthey Group, “Absorbent for Sulphur Polishing,” 

available at http://www.synetix.com/refineries/pdfs/manuals/669w.pdf 

10. Watson, J., Jones, K., “Removing H2S from Syngas Using Proven Technology in 

Japanese Waste Gasification Facility,” available at http://www.gtp-merichem.com/ 

support/technical_papers

11. Douglas L. Heguy, Gary J. Nagl, “The State of Iron Redox Sulfur Plant Technology New 

Developments to an Established Technology,” available at http://www.gtpmerichem.com/support/technical_papers

12. Dortmundt, D., Doshi, K., “Recent Developments in CO2 Removal Membrane 

Technology,” 1999. 

13. Edwards, M.S., “H2S Removal Processes for Low-Btu Coal Gas,” prepared by Oak 

Ridge National Laboratory, Oak Ridge, Tennessee, for the Department of Energy/Fossil 

Energy Office of Program Planning and Analysis, Contract No. W-7405-eng-26, January 

1979. 

14. Tennyson & Schaaf, “Guidelines Can Help Choose Proper Process for Gas Treating 

Plants,” Oil and Gas Journal, January 1977. 

15. W. Breckenridge, A. Holiday, J. Ong and C. Sharp “Use of SELEXOL Process in Coke 

Gasification to Ammonia Project” Laurance Reid Gas Conditioning Conference, 

February 17, 2000. 
 

 

Task 2 

Detailed MDEA Process Analysis 

NREL Task Order No. KAFT-8-882786-01 

September 2009 

Prepared for: 

United States Department of Energy / National Renewable Energy 

Laboratory (NREL) 
 

i

Contents 

Section Page 

 

Executive Summary....................................................................................................................................... 1 

Section 1 Introduction .............................................................................................................................. 4

Section 2 Methodology ............................................................................................................................ 5 

Section 3 Design Basis............................................................................................................................ 7 

Section 4 Processes Description ........................................................................................................... 12 

Section 5 Analysis Results and Discussion ........................................................................................... 16 

Section 6 Summary and Conclusions .................................................................................................... 27 

References ............................................................................................................................................... 28 

Appendix A Case 1 ................................................................................................................................... 29 

Appendix B Case 2 ................................................................................................................................... 33 

Figures 

Figure ES-1 Overview of NREL Thermochemical Ethanol Synthesis Process ............................................. 1 

Figure 4-1 Simplified MDEA Process Flow Diagram................................................................................ 15 

Figure 5-1 Case 1 - MDEA Process Flow Diagram.................................................................................. 17 

Figure 5-2 Case 1 – Estimated Annual Cost of Sulfur Injection ............................................................... 19 

Figure 5-3 Case 2 – Activated MDEA Process Flow Diagram................................................................. 22 

Figure 5-4 Case 2 – Estimated Annual Cost of Sulfur Injection ............................................................... 24 

Figure 5-5 Effects of Ammonia on Utilities Requirements........................................................................ 26 

Figure 5-6 Effects of Varying Acid Gas Concentrations on Utilities Requirements .................................. 26 

Tables 

Table ES-1 Summary of Analysis Results ................................................................................................... 2 

Table 3-1 Feed Gas to the MDEA Unit ...................................................................................................... 7 

Table 5-1 Case 1 – Overall Utilities Requirements .................................................................................. 18 

Table 5-2 Case 1 – Variable Operating Cost Estimates .......................................................................... 19 

Table 5-3 Case 1 – Summary of Results ................................................................................................. 20 

Table 5-4 Case 2 – Overall Utilities Requirements .................................................................................. 23 

Table 5-5 Case 2 – Variable Operating Cost Estimates .......................................................................... 23 

Table 5-6 Case 2 – Summary of Results ................................................................................................. 25 

Table A-1 Case 1 – Detailed Utilities Requirements ................................................................................ 30 

Table A-2 Case 1 - Equipment List.......................................................................................................... 31 

Table A-3 Case 1 - Equipment Cost Estimates........................................................................................ 32 

Table B-1 Case 2 – Detailed Utilities Requirements ................................................................................ 34 

Table B-2 Case 2 - Equipment List.......................................................................................................... 35 

Table B-3 Case 2 - Equipment Cost Estimates........................................................................................ 36 
 

ii

Acronyms 

AGR acid gas removal 

BR&E Bryan Research & Engineering

CWR cooling water return 

CWS cooling water supply 

DMDS dimethyl disulfide 

MDEA methyldiethanolamine 

MM million (1x106

) 

NREL National Renewable Energy Laboratory 

ppm parts per million 

ppmv parts per million volume 

SCF standard cubic feet (1 atm, 60°F) 
 

 Executive Summary

In March 2009 NREL engaged Nexant to conduct a survey and screening of various acid gas 

removal (AGR) technologies in order to evaluate their capability to meet the specific design 

requirements for the NREL thermochemical ethanol synthesis process (Figure ES-1), and to 

recommend applicable acid gas removal options for a more detailed analysis. The results of the 

survey, documented in a report entitled “Task 1 - Acid Gas Removal Techonology Survey and 

Screening for Thermochemical Ethanol Synthesis[ ] 1 ”, identified the MDEA and Selexol systems 

as suitable AGR options for the NREL design. As described in the Task 1 report, the gas 

cleanup conditioning process, of which the AGR unit is a part, has a significant impact on the 

overall plant economics since it has the highest cost component of all the process areas. Thus, 

the goal of this study is to provide a detailed design of an AGR unit so that NREL can better 

assess its technical and economic impacts on the overall process. NREL selected the MDEA 

system for the detailed analysis. 

The detailed analysis of an MDEA unit consists of (1) the development of an MDEA process 

model using the Promax simulation software in order to obtain the heat and energy balance of 

the MDEA unit and (2) the cost estimation of the MDEA unit using the ICARUS software. 

Source: Reference [2] 

Figure ES-1 Overview of NREL Thermochemical Ethanol Synthesis Process 

The MDEA system is designed to treat approximately 163 MMSCF of syngas containing 234 

ppmv H2S and 11.4 mol% CO2. The required purity of the treated gas, which is dictated by the 

sulfur requirement and CO2 tolerance of the alcohol synthesis catalyst, are 10 - 50 ppmv H2S 

and 1 - 5 mol% CO2. The amine solution used in the analysis is a generic 50 wt% MDEA 

solution. Additionally, since the NREL syngas has considerably higher CO2 concentration than 

H2S, a piperazine-activated MDEA solution to enhance CO2 absorption is also evaluated. 

Piperazine is a promoter that is highly reactive towards CO2; thus, when added to the MDEA 

Task 2: Detailed MDEA Process Analysis 

United States Department of Energy/National Renewable Energy Laboratory 

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Task 2: Detailed MDEA Process Analysis 

United States Department of Energy/National Renewable Energy Laboratory 

2 

solution, the solution’s capability for CO2 absorption can be significantly enhanced. The results 

of the analysis, summarized in Table ES-1, show that: 

Table ES-1 Summary of Analysis Results 

Design Feed Rates Case 1 Case 2 

Amine Type MDEA 

 Activated 

MDEA 

Amine Concentration wt% 50 

 50% MDEA/ 

2% PZ 

Amine Circulation rate gpm 1,534 1,374 

Steam Consumption klb/h 119 119 

Power Consumption kW 569 497 

Make-up Water gpm 40 43 

Dimethyl Disulfide Consumption (Note 1) gal/d 21-106 4-68 

Design Product Rates 

Treated Gas MMSCFD 153 153 

Treated Gas H2S Concentration ppmv 0.13 18.16 

Treated Gas CO2 Concentration mol % 4.95 4.92 

 

Heat Duty per lb Acid Gas Removed Btu/lb 2,053 2,053 

 

Capital & Operating Costs 

Total Installed Equipment Cost, (2Q09) million $ 11.6 9.7 

Annual Operating Costs (min) (Note 2) thousand $ 260.9 145.9 

Annual Operating Costs (max) (Note 2) thousand $ 1,044.7 736.6 

Notes: 

(1) The range of operating costs shown represents the minimum and maximum levels of DMDS injection 

required in order to maintain the H2S content of the clean gas between 10 – 50 ppmv. 

(2) The minimum and maximum operating costs shown reflect the level of DMDS consumption

o Case 1 – with the use of a generic MDEA solution, which has high selectivity for H2S, 

the level of H2S removal would have to be higher than necessary in order to meet the 

CO2 specification in the clean gas. Thus, the H2S content in the clean gas is lower 

than the required specification. In this case, the H2S content of the clean gas can be 

controlled between 10 and 50 ppmv—the concentration necessary to maintain 

catalyst activity—by injection of dimethyl disulfide (DMDS) into the clean gas stream. 

DMDS is commonly used as a sulfiding reagent in industrial applications to control 

catalyst activity. The annual operating costs of the unit can vary considerably, as 

shown in Table ES-1, depending on the consumption requirement of DMDS. Thus, it 

is desirable to maintain the H2S content of the clean gas at a concentration in which 

the operating cost can be minimized while preserving the catalyst activity. Catalyst 

vendors should be consulted for recommendations of the optimal H2S content in the 
 

Task 2: Detailed MDEA Process Analysis 

United States Department of Energy/National Renewable Energy Laboratory 

3 

clean gas. The total installed equipment cost for the MDEA unit in this case is 11.6 

million dollars. 

o Case 2 – with the use of piperazine as a promoter in the MDEA solution, the 

solution’s capability for CO2 absorption is significantly enhanced, resulting in 10% 

lower amine circulation rate and smaller absorber size for the MDEA unit as 

compared to Case 1. Consequently, due to the smaller plant size of the MDEA unit 

utilizing an activated MDEA solution, the total installed equipment cost, at 9.7 million 

dollars, is approximately 16% lower compared to the unit using a generic 50 wt% 

solution. Additionally, the study shows that as the solution’s capability for CO2

absorption increases, its capacity for H2S absorption decreases. Thus, the H2S 

content in the clean gas is higher compared to Case 1. Since the achieved H2S 

purity of 18 ppmv in the clean gas meets the minimum requirement, less DMDS 

would be required to maintain the H2S content between 10 - 50 ppmv. Depending on 

the level of DMDS consumption, the operating costs of the MDEA unit in this case 

are 30 to 45% lower compared to those in Case 1. 

o Effects of the presence of ammonia in the feed stream on the MDEA unit – The 

sensitivity analysis shows that presence of ammonia in the feed gas significantly 

increases the duties of the stripper reboiler by 25% and the condenser by 45%, as 

compared to the feed gas with no ammonia present. Equally important, the 

presence of ammonia eventually leads to buildup in the stripper overhead system 

and contributes to corrosion. To minimize ammonia buildup, purging of the reflux 

water is necessary. 

 

Thus, the results of the detailed analysis indicate that for the NREL thermochemical ethanol 

synthesis application, in which the syngas that has appreciable CO2 and low H2S 

concentrations, the MDEA unit utilizing an activated MDEA solution has more economics 

advantages over the one utilizing a generic MDEA solution. The use of a piperazine-activated 

MDEA solution enhances the rate of CO2 absorption which results in a smaller plant size, and 

consequently lower capital and operating costs. Moreover, in order to minimize the impact of 

ammonia on the unit’s operating costs, maximum removal of ammonia upstream of the AGR 

unit is critical. 
 

Task 2: Detailed MDEA Process Analysis 

United States Department of Energy/National Renewable Energy Laboratory 

4 

Section 1 Introduction

1.1 INTRODUCTION 

In March 2009 Nexant conducted a survey and screening of various acid gas removal (AGR) 

technologies to identify suitable acid gas removal options for the NREL thermochemical ethanol 

synthesis process and recommended the MDEA and Selexol systems for further detailed 

analysis. NREL selected the MDEA system for the detailed analysis in this Task 2 study. 

The objective of Task 2 study is to provide a detailed design analysis of the MDEA process for 

H2S and CO2 removal. The feed gas to the MDEA unit and the clean gas specifications are 

provided by NREL. The detailed analysis consists of (1) the development of an MDEA process 

model using the Promax simulation software in order to obtain the heat and energy balance of 

the MDEA unit and (2) the cost estimation of the MDEA unit. This study evaluated a generic 

MDEA solution as well as an activated MDEA solution, which contains piperazine to enhance 

CO2 absorption. The two cases evaluated in this study are: 

Case 1 – Generic MDEA 

Case 2 – Activated MDEA 

Additionally, a sensitivity analysis was also performed to determine the impact of ammonia and 

of varying acid gas concentrations in the feed gas on utilities requirements of the unit. 
 

Task 2: Detailed MDEA Process Analysis 

United States Department of Energy/National Renewable Energy Laboratory 

5 

Section 2 Methodology

The methodology and approaches used to develop the ProMax model and the capital cost 

estimate for the MDEA unit are described in the following sections. 

2.1 PROMAX MODELING METHODOLOGY 

ProMax simulation software was used to develop the process model for the MDEA unit. The 

ProMax simulation software, developed by Bryan Research & Engineering (BR&E), is widely 

used by the oil, gas, refining, and chemical industries for plant design and optimization. 

Capabilities of the ProMax simulation software include process areas such as amine 

sweetening, glycol dehydration, crude oil refining, and liquefied petroleum gas (LPG)/natural 

gas liquids (NGL) recovery and fractionation. 

The objective of the simulation model is to provide the heat and energy balances necessary to 

perform the equipment sizing and cost estimate task. A detailed modeling approach was used 

in developing the MDEA model. This approach requires the identification of the major 

equipment within the unit and the construction of first principal separation models and/or 

detailed kinetic models. The MDEA model was developed based on the feed gas data and 

treated gas specifications provided by NREL and the process equipment design basis 

discussed in Section 3. In the ProMax simulation, the Amine Sweetening Electrolytic property 

package was used. The TSWEET Kinetics model was used for the absorber while the 

TSWEET Alternate Stripper model was used for the regenerator. 

2.2 COST ESTIMATING APPROACH 

2.2.1 Total Installed Equipment Cost Estimate 

The total installed equipment cost was estimated using the equipment cost factoring approach. 

This approach is commonly used when detailed plant configuration of the process unit operation 

are available so that a major equipment list can be compiled and the size of each piece of 

equipment is determined. For this study, ICARUS was used to estimate the purchased (bare) 

equipment cost. Then, an appropriate installation factor was used to determine the field labor, 

piping, foundations, electrical, instrumentation, etc costs for each individual piece of equipment. 

This method is well founded both theoretically and in practice. It has been in use for many 

years in petroleum and chemical process industries for plant cost estimating. This method 

relies on the observation that the total installed cost of major equipment can be reliably 

represented as a multiple of the equipment cost. For a given type of equipment, the multiplier 

(called the installed cost factor) can vary depending on the size of the piece of equipment, 

specific process design details, site location, and other factors. 

2.2.2 Operating Costs 
 

Task 2: Detailed MDEA Process Analysis 

United States Department of Energy/National Renewable Energy Laboratory 

6 

In general, the operating costs consist of variable operating costs and fixed operating costs. 

The variable operating costs of the MDEA unit include utilities and chemicals costs. This study 

assumes that steam and power purchase would not be required for the MDEA unit since the 

steam and power requirements would be supplied by the steam/power generation cycle, as 

shown in the current NREL design. (Refer to the Steam System and Power Generation of 

NREL design, PFD-P800-A602).[ ] 2

Generally, the fixed operating costs are determined for the entire plant complex, but not for the 

individual process units within the plant. Therefore, this study does not include the fixed 

operating costs for the MDEA unit. 

Plant availability factor of 96%, which is the same basis as the current NREL design, is used for 

the operating cost estimates. 

 
 

Task 2: Detailed MDEA Process Analysis 

United States Department of Energy/National Renewable Energy Laboratory 

7 

Section 3 Design Basis

3.1 GENERAL DESIGN BASIS 

3.1.1 Design Feed Gas 

The feed gas to the MDEA unit, provided by NREL, is shown in Table 3-1. 

Table 3-1 Feed Gas to the MDEA Unit 

Feed Gas to AGR

Total Flow, lbmol/hr 17,942

Temperature, °F 110 

Pressure, psia 420

Compositions mole%

H2 42.9 

H2O 0.3 

CO 42.9 

N2 0.8 

CO2 11.4 

H2S 234 ppmv

NH3 82 ppmv 

Methane (CH4 ) 1.4 

n-Butane (C4H10) 0.03 

Ethane (C2H6) 13 ppmv

Ethylene (C2H4) 0.1 

Acetylene (C2H2) 0.0 

Propane (C3H8) 0.2 

Pentane + 42 ppmv

Benzene (C6H6) 3 ppmv

Tar (C10H8) 0.4 ppmv

3.1.2 Design Product Gas 

Approximately 153 MMSCFD of treated gas is produced from the MDEA unit. The treated gas 

is fed to the downstream alcohol synthesis reactor. The synthesis catalyst requires an H2S 

concentration of up to 50 ppmv to maintain catalyst activity, and it can also tolerate a CO2

concentration of up to 5 mol%. As agreed with NREL, acid gas concentrations of the treated 

gas should be within the following range: 

H2S: 10 to 50 ppmv 

CO2: 1 to 5 mol% 
 

Task 2: Detailed MDEA Process Analysis 

United States Department of Energy/National Renewable Energy Laboratory 

8 

3.1.3 Amine Solutions 

The two types of MDEA solutions evaluated in this study include: 

o 50 wt% generic MDEA solution. 

o Piperazine-activated MDEA solution to enhance CO2 absorption. 

3.1.4 Utilities 

The utilities design basis for Task 2 was assumed to be similar to that used in the current NREL 

design. The major utilities requirements of the MDEA unit are cooling water, steam, and power. 

The following design basis was used in Task 2 study: 

o Cooling water – supplied by the cooling water system, as shown on the current NREL 

design. CW is supplied at 90°F and returned at 110°F. 

o Steam – 50 psig steam, produced by the steam/power generation cycle per the current 

MDEA design, is supplied to the MDEA unit. 

o Electric power – supplied to the MDEA unit from the steam/power generation cycle as 

shown in the current MDEA design. Purchased power is not required. 

3.2 MAJOR PROCESS EQUIPMENT DESIGN BASIS 

Refer to Figure 4-1 for reference. 

3.2.1 Absorber 

o The lean solution inlet temperature is maintained at least 10°F higher than the 

temperature of the feed gas in order to avoid hydrocarbons condensation. 

Condensed hydrocarbons accumulated in the system may promote foaming. 

o The rich acid gas loading of the absorber is not to exceed 0.40 mol acid gas/mol 

MDEA. For this application, the feed has a significantly higher ratio of CO2 to H2S, 

resulting in the rich amine solution being loaded mostly with CO2. Since CO2 is more 

corrosive than H2S in aqueous solution, the rich acid gas loading is maintained at the 

mid-range of the recommended loading range rather than at the higher end. 

(Recommended rich acid gas loading range for MDEA: 0.2 - 0.55 mol acid gas/mol 

MDEA). 

o The absorber is a tray column, designed for a jet flooding rate of 70% and a foaming 

factor of 0.8. 

o Two minutes of surge time is provided between the HLL and LLL, based on rich 

amine flow rate. 
 

Task 2: Detailed MDEA Process Analysis 

United States Department of Energy/National Renewable Energy Laboratory 

9 

o In the ProMax simulation, the TSWEET Kinetics absorber model was used to 

determine the number of theoretical stages and the MDEA circulation rate required to 

achieve the treated gas specifications. Tray efficiency of 20% was used to calculate 

the number of actual trays. Ideally, the actual operating data from a similar unit 

would be used to estimate the tray efficiency, but since such data is not available at 

present, tray efficiency was obtained from the open literature, which reported tray 

efficiency for gas absorption applications ranging from 10 - 30%. 

3.2.2 Stripper 

o The operating pressure of the stripper is set such that the reboiler bottom 

temperature does not exceed 255°F in order to minimize thermal degradation of the 

MDEA solution, which can lead to corrosion. 

o The operating temperature of the stripper condenser is set at 120°F in order to 

maintain a 10°F approach temperature with the cooling water. 

o Lean acid gas loading range: 0.004 to 0.01 mol acid gas/mol MDEA. 

o Heating medium for the reboiler is 50 psig steam generated from the steam/power 

cycle. 

o The stripper is a tray column, designed for a jet flooding rate of 85% and a foaming 

factor of 0.85. 

o In the ProMax simulation, the TSWEET Alternate Stripper model was used to 

determine the number of theoretical stages and the reboiler duty required to achieve 

the lean acid gas loading that is within the range specified above. Tray efficiency of 

40%, obtained from literature, was used to calculate the number of actual trays. 

3.2.3 Lean/Rich Exchanger 

o The exchanger’s fluid velocities are kept at a maximum of 4 ft/s to minimize 

corrosion. 

o Minimum end approach temperature of at least 20°F (lean amine temperature out – 

rich amine temperature in). 

3.2.4 Lean Amine Treating and Cooling 

o The lean amine solution from the stripper is cooled by heat exchanged with the rich 

solution in the Lean/Rich Exchanger. The lean solution is further cooled to 120°F by 

water cooling in the Lean Solution Cooler, which has a minimum approach 

temperature of 10°F. 
 

Task 2: Detailed MDEA Process Analysis 

United States Department of Energy/National Renewable Energy Laboratory 

10 

o A filtration unit is provided for solution maintenance. The presence of particulates in 

the amine solution tends to promote solution foaming while degradation products can 

contribute to foaming and corrosion. The filtration unit is a package unit, consisting 

of a 5-micron cotton pre-filter to remove particulates, followed by an activated carbon 

bed to remove degradation products, and ends with a 5-micron cotton post-filter to 

remove any carry-over carbon particles to minimize solution foaming tendencies. 

Approximately 10% of the lean amine solution is continuously filtered. 

o Since the acid gases to the MDEA unit are primarily CO2 and H2S, and their 

reactions with MDEA form regenerable compounds that can be regenerated in the 

stripper, a reclaimer unit is not provided in this design. 

o An amine sump is provided to prepare make-up solutions and to collect and recover 

amine from drainage lines. 

3.2.5 Materials of Construction 

o Carbon steel is used, except in portions of the plant that are exposed to wet CO2 or 

CO2/H2S evolution, in which stainless steel is used. All equipment is stress relieved. 

3.3 OTHER CONSIDERATIONS 

3.3.1 Impurities in Feed Stream 

o Ammonia – As shown in Table 3-1, the feed gas to the MDEA unit contains a small 

quantity of ammonia. Due to the high solubility of ammonia in water, virtually all of 

the ammonia in the feed gas stream is absorbed by the amine solution in the 

absorber. Ammonia is then released along with the H2S and CO2 in the stripper. In 

the stripper condenser, it is condensed with the water and is returned to the column 

as reflux. Overtime, ammonia builds up in the stripper overhead system and 

contributes to corrosion. Additionally, the presence of ammonia also impacts the 

energy requirements of the stripper condenser and reboiler. The condensation of 

ammonia is highly exothermic while the vaporization of ammonia requires 

considerable amount of heat, consequently resulting in increased condenser and 

reboiler duties. Ammonia buildup can be minimized by purging a portion of the reflux 

water. 

o HCl – although the feed gas in this study does not contain HCl, but according to 

NREL, some biomass-derived gases may contain small quantities of HCl. HCl reacts 

with the amine to form heat-stable, degradation compounds that cannot be 

regenerated in the stripper. Generally in this situation, soda ash is added into the 

solution to neutralize the heat-stable compounds. Once concentration of these 

compounds reaches about 4 wt%, then the MDEA solution is reclaimed[ ] 3 . 
 

Task 2: Detailed MDEA Process Analysis 

United States Department of Energy/National Renewable Energy Laboratory 

11 

3.3.2 Emissions of Absorbed Product Gases 

The feed gas to the MDEA unit contains high concentrations of hydrogen, CO, and small 

quantities of hydrocarbons, primarily methane. These compounds are absorbed to some extent 

in the absorber. The absorption of these compounds has an environment impact since these 

gases, which exit the MDEA unit with the acid gas stream, would be vented with the CO2 stream 

from the LO-CAT sulfur recovery unit (Figure ES-1). (CO is a criteria pollutant and methane is a 

greenhouse gas). 

According to the EPA, the allowable concentrations of CO and methane are dependent on the 

specific application and plant location. Since a site has not yet been selected for the NREL 

thermochemical ethanol process, and the acceptable emissions level cannot be determined at 

this time, the MDEA unit is designed to minimize the emissions level of CO and methane by 

depressuring the rich amine solution in order to recover the dissolved gases. 
 

Task 2: Detailed MDEA Process Analysis 

United States Department of Energy/National Renewable Energy Laboratory 

12 

Section 4 Processes Description

4.1 PROCESS DESCRIPTION 

Figure 4-1 illustrates a simplified process flow diagram of the MDEA unit. Two types of MDEA 

solutions are considered for the process: a generic 50 wt% MDEA solution and a piperazine 

activated MDEA solution. Descriptions of the process and the activated MDEA solution are 

provided below. 

Amine Absorption

The sour gas enters the bottom of the absorber at 110°F and 420 psia. The lean MDEA solvent 

enters the top of the absorber at 120°F. As the solvent contacts the gas stream, the acid gases 

(H2S and CO2) react with the MDEA. As a result, the H2S and CO2 in the feed gas are 

transferred to the solvent. Additionally, small amounts of CO, methane, and hydrogen, and 

virtually all of the NH3 in the feed gas are also absorbed into the MDEA solution. Makeup water 

is added to the water trays at the top of the column. 

The treated gas exits at the top of the absorber, where a demister pad is installed near the 

overhead gas outlet to trap the entrained MDEA solvent. A knockout drum is provided for the 

treated gas to collect residual solvent carryover. If necessary, the treated gas is injected with a 

small amount of dimethyl disulfide (DMDS) to control the H2S content in order to maintain the 

catalyst activity of the downstream alcohol synthesis reactor. DMDS is commonly used as a 

sulfiding reagent in industrial applications to control catalyst activity. In this design, DMDS is 

pumped and injected into the clean gas stream after the gas is compressed and preheated to 

the operating conditions of the alcohol synthesis reactor. Injection into a hot gas line is 

preferable so that the liquid DMDS can be vaporized. 

Rich Solution Flashing

The rich amine solution from the absorber bottom is depressured through the Flash Drum to 

about 50 psia to recover the CO, methane, and hydrogen dissolved in the MDEA solution. (The 

ammonia remains in the rich solution). Depressuring the rich solution to recover the dissolved 

gases helps to minimize their emissions into the atmosphere as they would be vented along 

with the CO2 stream from the LO-CAT sulfur recovery unit. 

The flashed gas stream is to be combined with the unreacted syngas stream (operating at 34 

psia) from the alcohol synthesis area (NREL design, PFD-P800-A402, stream 426)[ ] 2 and fed to 

the Catalyst Regenerator (R-301A) as purge gas. In the Catalyst Regenerator, the CO and 

methane compounds are destroyed by incineration. 

Another option to handle the flashed gas would be to re-compress and recycle the gas to the 

absorber. However, because the flashed gas stream is fairly small—the CO, H2, and methane 

contents in the flashed gas are equivalent to just 0.1%, or less, of the respective feed gas 
 

Task 2: Detailed MDEA Process Analysis 

United States Department of Energy/National Renewable Energy Laboratory 

13 

components—it would not justify the added capital cost and complexity of the recycle loop. For 

this reason, the flashed gas recycle option is not considered for the MDEA design. 

Amine Stripping, Treating & Cooling 

The rich solvent from the flash drum is preheated before entering the stripper by exchanging 

heat with the lean MDEA stream from the stripper. The Lean/Rich Exchanger is a shell and 

tube exchanger, with the rich solvent on the tube side. In the stripper, as the rich MDEA 

solution travels down the column, the H2S and CO2 are stripped from the solution by the 

stripping steam generated in the kettle reboiler. 

The stripped acid gases exit at the top of the stripper and enter the condenser where the water 

vapor is condensed and returned to the column as reflux. A portion of the reflux water is purged 

in order to minimize ammonia buildup in the stripper. The purged stream is sent to the waste 

water treatment facility. The acid gas from the reflux drum overhead is sent to the LO-CAT 

sulfur recovery unit for processing. 

Lean MDEA solution from the bottom of the stripper, after cooling by heat exchange with the 

rich solution, is sent to the MDEA Surge Tank. The lean solution is then pumped to about 450 

pisa and cooled to120°F before it enters the absorber. 

A portion of the lean solution, approximately 10%, is continuously filtered through the MDEA 

Pre-Filter to remove solids which tend to promote solution foaming as well as plugging of 

downstream carbon filter. From the MDEA Pre-Filter, the lean solution passes through the 

MDEA Carbon Bed to remove degradation products which tend to promote foaming and 

corrosion. Carbon filter effluent then passes through the MDEA Post-Filter to remove any carryover carbon particle. 

A sump is provided for preparation of makeup solutions and for collection of drainage to recover 

amine. All amine containing lines and equipment are to drain to the amine sump to minimize 

amine losses. 

4.2 PIPERAZINE-ACTIVATED MDEA SOLUTION 

The rate of CO2 absorption in MDEA solution is considerably lower compared to that of H2S 

absorption due to the low reaction rate of MDEA with CO2. The CO2 absorption rate can be 

significantly enhanced by the addition of promoters, such as piperazine. For many years, 

piperazine-activated MDEA has been successfully used in natural gas, ammonia, and syngas 

plants for high capacity CO2 removal. The use of piperazine-activated MDEA was patented by 

BASF, until 2002, under the name aMDEA. At present, MDEA with piperazine promoter is 

offered by most solution vendors under various trade names. 

The advantages of piperazine are its high reactivity towards CO2 and its high carrying capacity 

for CO2. According to studies conducted by Bishnoi and Rochelle [4]

, the addition of piperazine 

(5 wt%) to 45 wt% MDEA increases the CO2 absorption rate by approximately two orders of 
 

Task 2: Detailed MDEA Process Analysis 

United States Department of Energy/National Renewable Energy Laboratory 

14 

magnitude at low loading and one order of magnitude at moderate loading compared to 50 wt% 

MDEA. These characteristics make piperazine-activated solution suitable for applications with 

appreciably high CO2 concentration relative to H2S. 

Typically, the piperazine concentration in the MDEA solution is specified by the solution vendors 

based on the desired removal target. BASF’s aMDEA process uses a 2.5-4.5 M (30-54 wt%) 

MDEA solution containing up to 0.8 M (7 wt%) piperazine [ ] 3 . For this study, the amine solution 

evaluated is a mixture of 50 wt% MDEA and 2 wt% piperazine. Higher piperazine 

concentrations were also evaluated, and the simulation results indicated that piperazine addition 

of about 2 wt% would be sufficient to achieve the required acid gas removal target for the NREL 

design. 
Task 2: Detailed MDEA Process Analysis 

United States Department of Energy/National Renewable Energy Laboratory 

Figure 4-1 Simplified MDEA Process Flow Diagram 

D-302

Flash Drum

C-301

Absorber

E-302

Lean/Rich

Exchanger

C-302

Stripper E-301 Lean Amine Cooler

Feed Gas

D-301

Treated Gas

KO Drum

G-301

Lean Amine Pump

To Catalyst Regenerator

E-303

Stripper Condenser

Filter

Makeup Wtr

Package unit

G-304

Amine Sump Pump

T-301

MDEA Surge Tank

MDEA Make-up

T-302

MDEA Sump

Clean Gas Compression & 

Preheat

990 psia, 570°F

Dimethyl Disulfide (DMDS) 

Injection 

G-305 

DMDS Injection Pump

NREL 

S

Task 2: 

CWS

CWS

CWR
 

Task 2: Detailed MDEA Process Analysis 

United States Department of Energy/National Renewable Energy Laboratory 

16 

Section 5 Analysis Results and Discussion

5.1 OVERVIEW 

In this study, a detailed analysis of the MDEA acid gas removal unit was performed for the 

NREL thermochemical ethanol synthesis process. The detailed analysis entails the 

development of an MDEA process model to obtain the heat and energy balance of the MDEA 

unit and the cost estimation of the unit. The MDEA solutions evaluated are generic MDEA 

(Case 1) and activated MDEA (Case 2) solutions. The following sections present the analysis 

results of the two cases. 

 

Additionally, a sensitivity analysis was also performed to determine the impact of ammonia and 

of varying acid gas concentrations in the feed gas on the utilities requirements of the unit. 

5.2 CASE 1 – GENERIC MDEA 

5.2.1 Discussion of Results 

Figure 5-1 illustrates the MDEA unit utilizing a generic 50 wt% MDEA solution to treat the NREL 

biomass-derived syngas. The unit’s material and energy balance are also shown on Figure 5-1. 

In this case, the absorber with 32 total actual trays—30 trays for acid gas absorption and 2 trays 

for makeup water addition—and an amine circulation rate of approximately 1533 gpm are 

required to treat the sour feed gas containing 234 ppmv H2S and 11.4 mol% CO2. The achieved 

purity of the clean gas is 0.13 ppmv H2S and 4.95 mol% CO2. Due to the selectivity of MDEA 

for H2S, the amount of H2S removal has to be higher than necessary in order to meet the CO2

specification in the clean gas. As a result, the H2S content in the clean gas is lower than the 

specified 10 – 50 ppmv required to maintain the alcohol synthesis catalyst activity. To increase 

the H2S content of the clean gas, DMDS solution is injected into the clean gas after the gas is 

compressed and preheated to the operating conditions of the alcohol synthesis reactor. The 

cost and the amount of DMDS required to achieve various levels of H2S content in the clean gas 

are shown in the operating cost section. 

The rich amine solution entering the amine stripper has an acid gas loading of 0.37 mol/mol 

MDEA. The stripper unit, which has 20 total actual trays plus 1 chimney tray for reboiler draw, 

can achieve a lean loading of 0.004 mol acid gas/mol MDEA by utilizing 119,000 lb/hr of 50 psig 

steam. For this unit, the heat duty per lb of acid gas removed is 2,053 Btu/lb. 

From the reflux drum, approximately 30% of the reflux water, or 35 gpm, is purged in order to 

minimize ammonia buildup in the amine stripper. The water balance of the unit is maintained by 

the addition of about 40 gpm of make-up water to the absorber.
Stream No. 100 101 102 103 105 106 107 109 113 114 119 120 125 126 130

Temperature F 110 129 161 161 203 224 120 120 255 255 120 135 80 129 298

Pressure psia 420 417 420 50 40 25 20 20 28 445 28 1,000 450 417 65

Vapor Fraction 1.00 1.00 1.00 - 0.04 1.00 1.00 1.00 - - - - 1.00 - 1.00Mass Flow lb/hr 332,660 844,571 279,645 843,892 679 113,069 17,472 115,249 54,830 771,589 791,556 19,967 279,674 40 119,040

Molar Flow lbmol/hr 17,942 16,749 26,464 26,442 22 4,537 1,312 967 6,254 24,163 25,271 1,108 16,750 0.4 6,608

H2 lbmol/hr 7,687.0 7,683.0 4.0 3.8 0.2 0.2 0.2 0.0 - - - - 7,683.0 - - 

H2O lbmol/hr 50.6 82.6 21,903.8 21,902.0 1.8 3,323.6 110.5 6,171.0 20,827.7 964.0 21,935.8 1,108.2 82.6 - 6,607.7

CO lbmol/hr 7,690.8 7,687.3 3.5 3.3 0.2 0.2 0.2 0.0 - - - - 7,687.3 - - 

N2 lbmol/hr 148.7 148.7 0.0 0.0 0.0 0.0 0.0 0.0 - - - - 148.7 - - 

CO2 lbmol/hr 2,040.9 829.8 1,224.6 12.1 1,212.5 1,203.2 1,197.1 77.8 1.8 13.5 13.5 - 829.8 - - 

H2S lbmol/hr 4.2 0.13 ppmv 4.2 0.0 4.2 4.2 4.2 0.0 0.0 0.0 0.0 - 50 ppmv - -

Ammonia lbmol/hr 1.5 0.0 1.5 0.0 1.5 4.8 0.0 1.4 0.1 0.0 0.0 - - 0.0 - 

Methane lbmol/hr 251.1 250.7 0.4 0.3 0.1 0.1 0.1 0.0 - - - - 250.7 - - 

n-Butane lbmol/hr 5.0 5.0 0.0 0.0 0.0 0.0 0.0 0.0 - - - - - 5.0 - 

Ethane lbmol/hr 0.2 0.2 0.0 0.0 0.0 0.0 0.0 0.0 - - - - - 0.2 - 

Ethylene lbmol/hr 22.6 22.5 0.1 0.1 0.0 0.0 0.0 0.0 - - - - 22.5 - - 

Acetylene lbmol/hr 2.3 2.2 0.1 0.0 0.1 0.1 0.1 0.0 - - - - - 2.2 - 

Propane lbmol/hr 36.2 36.1 0.1 0.1 0.0 0.0 0.0 0.0 - - - - 36.1 - - 

Pentane lbmol/hr 0.8 0.8 0.0 0.0 0.0 0.0 0.0 0.0 - - - - - 0.8 - 

Benzene lbmol/hr 0.1 0.0 0.0 0.0 0.0 0.0 0.0 0.0 - - - - - 0.0 - 

Naphthalene lbmol/hr 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 - - 0.0 - 

MDEA lbmol/hr - 3,321.3 0.0 3,321.3 0.0 0.0 0.0 0.0 3,321.3 5.5 3,321.3 0.0 - 0.0 - 

DMDS lbmol/hr 0.4

Enthalpy MMBtu/hr -721.97 -517.85 -3,507.36 -2.39 -3,453.15 -543.78 -213.58 -118.10 -646.77 -3,075.27 -3,303.25 -136.11 -0.01 -517.86 -675.79

Task 2: Detailed MDEA Process Analysis 

United States Department of Energy/National Renewable Energy Laboratory 

Figure 5-1 Case 1 - MDEA Process Flow Diagram 
 

Task 2: Detailed MDEA Process Analysis 

United States Department of Energy/National Renewable Energy Laboratory 

18 

5.2.1.1 Utilities Requirements 

Table 5-1 summarizes the overall utilities requirements of the MDEA unit. Detailed utilities 

requirements of individual equipment are provided in Appendix A. The unit requires 

approximately 119,000 lb/hr of 50 psig steam for the stripper reboiler, 10,497 gpm of cooling 

water circulation for the cooling and/or condensing of the stripper overhead and the lean amine 

solution, and 569 kW of electric power, mainly for the pumping of the lean solution to the 

absorber. Additionally, 40 gpm of make-up water to the absorber is also required. 

As noted in the design basis in Section 2, it is assumed that the steam and power requirements 

of the MDEA unit are supplied by the steam/power generation cycle as shown in the current 

NREL design (PFD-P800-A602)[ ] 2 . Thus the steam/power generation cycle design should be 

revised based on the utilities requirements provided from this study. 

 Table 5-1 Case 1 – Overall Utilities Requirements 

Steam (50 psig), klb/hr 119 

Electric Power, kW 569 

Cooling Water Circulation, gpm 10,497 

Make-up Water, gpm 40 

5.2.1.2 Equipment Cost Estimates 

The detailed equipment list, including equipment sizes, and the equipment cost estimates for 

the MDEA unit are shown in Appendix A. The equipment cost was estimated using ICARUS 

based on 2nd quarter 2009 dollars. For this case, the bare equipment cost is 4.4 million dollars. 

The installation factor is calculated to be 2.61 and the total installed equipment cost is 11.6 

million dollars. 

5.2.1.3 Variable Operating Costs 

The variable operating costs of the MDEA unit, which include the utilities and chemicals costs, 

are shown in Table 5-2. The variable operating costs estimates are based on 96% plant 

availability. The MDEA make-up requirement is based on the amine loss of approximately 0.91 

lb MDEA per MMSCF of treated gas. Amine loss includes losses from entrainment and 

vaporization. 

 
 

Table 5-2 Case 1 – Variable Operating Cost Estimates 

 Consumption Cost 

Utilities & Chemicals Unit Cost Initial per Day Initial Annual

Water (1000 gal) 1.03 $/1000gal - 57 - $20,662

MDEA Solution (gal) 8.38 $/gal 25,408 16 $212,923 $46,929

DMDS Solution (gal), (min) 26.25 $/gal - 21 - $193,338

DMDS Solution (gal), (max) 26.25 $/gal - 106 - $977,152

Total Variable Operating Costs (min) $260,930

Total Variable Operating Costs (max) $1,044,744

The DMDS solution consumptions shown on Table 5-2 are the quantities required per day in 

order to increase the H2S content of the clean gas from 0.13 ppmv to 10 ppmv (minimum 

requirement) and to 50 ppmv (maximum requirement), respectively. As Figure 5-2 shows, the 

annual operating cost of sulfur injection can increase considerably with increasing H2S content 

of the clean gas. Therefore, consultation with the catalyst vendor is recommended to obtain the 

optimal H2S content in the clean gas such that the catalyst activity can be maintained while 

minimizing the operating cost. 

Figure 5-2 Case 1 – Estimated Annual Cost of Sulfur Injection 

-

200,000

400,000

600,000

800,000

1,000,000

10 20 30 40 50

H2S Content in Clean Gas (ppmv)

Annual Cost of DMDS ($/year)

5.2.2 Summary of Results 

Table 5-3 summarizes the results of Case 1 discussed above. 

Task 2: Detailed MDEA Process Analysis 

United States Department of Energy/National Renewable Energy Laboratory 

19 
 

Task 2: Detailed MDEA Process Analysis 

United States Department of Energy/National Renewable Energy Laboratory 

20 

 Table 5-3 Case 1 – Summary of Results 

Design Feed Rates Case 1 

Feed Gas MMSCFD 163 

Feed Gas H2S Concentration ppmv 234 

Feed Gas CO2 Concentration mol % 11 

Amine Type MDEA 

Amine Concentration wt% 50 

Amine Circulation rate gpm 1,534 

Steam Consumption klb/h 119 

Power Consumption kW 569 

Make-up Water gpm 40 

DMDS Consumption (Note 1) gal/d 21-106 

 

Design Product Rates 

Treated Gas MMSCFD 153 

Treated Gas H2S Concentration ppmv 0.13 

Treated Gas CO2 Concentration mol % 4.95 

 

Heat Duty per lb Acid Gas Removed Btu/lb 2,053 

 

Capital & Operating Costs 

Total Installed Equipment Cost, (2Q09) million $ 11.6 

Annual Operating Costs (min) (Note 2) thousand $ 260.9 

Annual Operating Costs (max) (Note 2) thousand $ 1,044.7 

Notes: 

(1) The range of DMDS injection shown represents the minimum and maximum levels of sulfur injection 

required in order to maintain the H2S content of the clean gas between 10 – 50 ppmv. 

(2) The minimum and maximum operating costs shown reflect the level of DMDS consumption. 

 
 

Task 2: Detailed MDEA Process Analysis 

United States Department of Energy/National Renewable Energy Laboratory 

21 

5.3 CASE 2 – ACTIVATED MDEA 

5.3.1 Discussion of Results 

Figure 5-3 illustrates the MDEA unit utilizing a piperazine-activated MDEA solution (50 wt% 

MDEA and 2 wt% piperazine) to treat the NREL biomass-derived syngas. The unit’s material 

and energy balance are also shown on Figure 5-2. In this case, the absorber with 12 total 

actual trays—10 trays for acid gas absorption and 2 trays for makeup water addition—and an 

amine circulation rate of approximately 1406 gpm are required to treat the sour feed gas 

containing 234 ppmv H2S and 11.4 mol% CO2. The achieved purity of the clean gas is 18 ppmv 

H2S and 4.92 mol% CO2. In this case, the H2S content meets minimum concentration required 

for alcohol synthesis. If a higher H2S concentration in the treated gas is desirable, DMDS 

injection similar to Case 1 can be utilized.

As discussed in Section 4.2, piperazine can significantly increase the rate of CO2 absorption into 

the amine solution. As more CO2 is absorbed by the solution, its capacity for H2S absorption 

decreases. Consequently, less H2S is removed, resulting in a higher H2S content in the clean 

gas compared to Case 1. Additionally, due to the enhanced CO2 absorption capability of the 

amine solution, less amine circulation rate and fewer absorber stages are required, as 

compared to Case 1, to achieve the treated gas specifications. The results indicated that this 

case requires 10% less circulation rate and 20 fewer absorption stages compared to Case 1. 

The rich amine solution entering the amine stripper has an acid gas loading of 0.39 mol/mol 

MDEA. The stripper unit, which has 20 total actual trays plus 1 chimney tray for reboiler draw, 

can achieve a lean loading of 0.005 mol acid gas/mol MDEA by utilizing 119,000 lb/hr of 50 psig 

steam. From the reflux drum, approximately 30% of the reflux water, or 35 gpm, is purged in 

order to minimize ammonia buildup in the amine stripper. The water balance of the unit is 

maintained by the addition of about 43 gpm of make-up water to the absorber. 
Stream No. 100 101 102 103 105 106 107 109 113 114 119 120 125 126 130

Temperature F 157 161 161 204 222 120 120 254 254 120 110 80 135 157 298 

Pressure psia 417 420 420 50 40 25 20 20 28 28 445 450 1,000 417 65 Vapor Fraction 1.00 1.00 - 1.00 0.04 1.00 1.00 - 1.00 - - - 1.00 1.00 1.00 

Mass Flow lb/hr 332,660 281,043 760,912 504 760,407 113,267 55,166 17,430 115,210 687,812 709,295 21,483 281,072 119,280 25

Molar Flow lbmol/hr 17,942 16,829 23,128 17 23,111 4,540 1,323 965 6,251 20,823 22,016 1,192 0.3 16,830 6,621

H2 lbmol/hr 7,687.0 7,683.5 3.5 3.3 0.2 0.2 0.2 0.0 - - - - - 7,683.5 - H2O lbmol/hr 50.6 163.8 18,745.0 1.4 18,743.7 3,321.4 115.8 6,187.4 17,666.2 18,858.3 961.7 1,192.1 163.8 6,621.0 -

CO lbmol/hr 7,690.8 7,687.7 3.0 2.9 0.1 0.1 0.1 0.0 - - - - - 7,687.7 - N2 lbmol/hr 148.7 148.7 0.0 0.0 0.0 0.0 0.0 0.0 - - - - - 148.7 - CO2 lbmol/hr 2,040.9 827.7 1,229.6 1,221.0 1,208.9 1,202.8 8.6 1.8 45.2 16.4 16.4 - - 827.7 - 

H2S lbmol/hr 4.2 18 ppmv 3.9 0.0 3.9 3.9 3.9 0.0 0.0 0.0 0.0 - - 50 ppmv - Ammonia lbmol/hr 1.5 0.0 1.5 0.0 1.5 4.8 0.0 1.4 0.0 0.0 0.0 - - 0.0 - Methane lbmol/hr 251.1 250.8 0.4 0.3 0.1 0.1 0.1 0.0 - - - - - 250.8 - 

n-Butane lbmol/hr 5.0 5.0 0.0 0.0 0.0 0.0 0.0 0.0 - - - - - 5.0 - Ethane lbmol/hr 0.2 0.2 0.0 0.0 0.0 0.0 0.0 0.0 - - - - - 0.2 - Ethylene lbmol/hr 22.6 22.5 0.1 0.1 0.0 0.0 0.0 0.0 - - - - - 22.5 - 

Acetylene lbmol/hr 2.3 2.2 0.1 0.0 0.1 0.1 0.1 0.0 - - - - - 2.2 - Propane lbmol/hr 36.2 36.1 0.1 0.1 0.0 0.0 0.0 0.0 - - - - - 36.1 - 

Pentane lbmol/hr 0.8 0.8 0.0 0.0 0.0 0.0 0.0 0.0 - - - - - 0.8 - 

Benzene lbmol/hr 0.1 0.0 0.0 0.0 0.0 0.0 0.0 0.0 - - - - - 0.0 - Naphthalene lbmol/hr 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 - - 0.0 - MDEA lbmol/hr - 0.0 2,976.1 2,976.1 0.0 0.0 0.0 0.0 5.7 2,976.1 2,976.2 0.0 - 0.0 - 

Piperazine lbmol/hr - 0.0 164.7 164.7 0.0 0.0 0.0 0.0 12.4 164.7 164.7 0.0 - 0.0 - 

DMDS lbmol/hr 0.3

Enthalpy MMBtu/hr -721.97 -522.52 -3,064.19 -1.74 -3,017.64 -544.56 -215.09 -117.82 -642.88 -2,638.12 -2,864.74 -146.42 -0.01 -522.53 -677.15

Task 2: Detailed MDEA Process Analysis 

United States Department of Energy/National Renewable Energy Laboratory 

Figure 5-3 Case 2 – Activated MDEA Process Flow Diagram 
 

Task 2: Detailed MDEA Process Analysis 

United States Department of Energy/National Renewable Energy Laboratory 

23 

5.3.1.1 Cass 2 Utilities Requirements 

Table 5-4 summarizes the utilities requirements of the MDEA unit. The unit requires 

approximately 119,000 lb/hr of 50 psig steam for the stripper reboiler, 9,996 gpm of cooling 

water circulation for the cooling and/or condensing of the stripper overhead and the lean amine 

solution, and 497 kW of electric power, mainly for the pumping of the lean solution to the 

absorber. Additionally, 43 gpm of make-up water to the absorber is also required. 

 Table 5-4 Case 2 – Overall Utilities Requirements 

Steam (50 psig), klb/hr 119

Electric Power, kW 497

Cooling Water Circulation, gpm 9,996

Make-up Water, gpm 43

5.3.1.2 Case 2 Equipment Cost Estimates 

The detailed equipment list, including equipment sizes, and the equipment cost estimates for 

the MDEA unit are shown in Appendix B. The equipment cost was estimated using ICARUS 

based on 2nd quarter 2009 dollars. For this case, the bare equipment cost is 3.7 million dollars. 

The installation factor is calculated to be 2.61, and the total installed equipment cost is 9.7 

million dollars. 

5.3.1.3 Case 2 Variable Operating Cost 

The variable operating costs of the MDEA unit, which include the utilities and chemicals costs, 

are shown in Table 5-5. The variable operating costs estimates are based on 96% plant 

availability. The MDEA make-up requirement is based on the amine loss of approximately 1.5 

lb MDEA per MMSCF of treated gas. Amine loss includes losses from entrainment and 

vaporization. 

 Table 5-5 Case 2 – Variable Operating Cost Estimates 

 Consumption Cost 

Utilities & Chemicals Unit Cost Initial per Day Initial Annual

Water (1000 gal) 1.03 $/1000gal - 62 - $22,299

Activated MDEA Solution (gal) 9.22 $/gal 21,248 27 $195,864 $87,384

DMDS Solution (gal), (min) 26.25 $/gal - 4 - $36,263

DMDS Solution (gal), (max) 26.25 $/gal - 68 - $626,946

Total Variable Operating Costs (min) $145,946

Total Variable Operating Costs (max) $736,629

For this case, the clean gas has an H2S concentration of approximately 18 ppmv, which meets 

the minimum requirement needed to maintain the alcohol synthesis catalyst activity. Therefore, 

DMDS injection may not be necessary, depending on the catalyst vendor’s recommendation. 
 

Figure 5-4 shows the incremental annual cost of DMDS injection to increase the H2S 

concentration in the clean gas up to 50 ppmv, should this option be desirable. 

Figure 5-4 Case 2 – Estimated Annual Cost of Sulfur Injection 

-

100,000

200,000

300,000

400,000

500,000

600,000

700,000

10 20 30 40 50

H2S Content in Clean Gas (ppmv)

Annual Cost of DMDS ($/year)

5.3.2 Summary of Results 

Table 5-6 summarizes the results of Case 2 discussed above. 

 

Task 2: Detailed MDEA Process Analysis 

United States Department of Energy/National Renewable Energy Laboratory 

24 
 

Task 2: Detailed MDEA Process Analysis 

United States Department of Energy/National Renewable Energy Laboratory 

25 

Table 5-6 Case 2 – Summary of Results 

Design Feed Rates Case 2 

Feed Gas MMSCFD 163 

Feed Gas H2S Concentration ppmv 234 

Feed Gas CO2 Concentration mol % 11 

Amine Type Activated MDEA 

Amine Concentration (MDEA/PZ) wt% 50/2 

Amine Circulation rate gpm 1,374 

Steam Consumption klb/h 119 

Power Consumption kW 497 

Make-up Water gpm 43 

DMDS Consumption (Note 1) gal/d 4-68 

 

Design Product Rates 

Treated Gas MMSCFD 153 

Treated Gas H2S Concentration ppmv 18 

Treated Gas CO2 Concentration mol % 4.92 

 

Heat Duty per lb Acid Gas Removed Btu/lb 2,053 

 

Capital & Operating Costs 

Total Installed Equipment Cost, (2Q09) million $ 9.7 

Annual Operating Costs (min) (Note 2) thousand $ 145.9 

Annual Operating Costs (max) (Note 2) thousand $ 736.6 

Note: 

(1) The range of DMDS injection shown represents the minimum and maximum levels of sulfur injection 

required in order to maintain the H2S content of the clean gas between 10 – 50 ppmv. 

(2) The minimum and maximum operating costs shown reflect the level of DMDS consumption 

5.4 SENSITIVITY ANALYSIS 

The sensitivity analysis examines the impact of ammonia and of varying acid gas concentrations 

in the feed stream on the utilities requirements of the MDEA unit. 

As noted, the presence of ammonia in the feed gas can significantly increase the duties of the 

stripper unit’s condenser and reboiler because the condensation of ammonia is highly 

exothermic and the vaporization of ammonia requires considerable amount of heat. This effect 

is illustrated in Figure 5-5, which shows an increase in duties of about 25% in the reboiler and 

45% in the condenser, compared to the feed gas with no ammonia present. As the impact of 

ammonia on the unit’s utilities requirements is substantial, coupled with the potential corrosion 

problems due to ammonia buildup in the stripper overhead system, the removal of ammonia 

upstream of the AGR unit is critical to minimize operating problems as well as operating costs. 
 

Figure 5-5 Effects of Ammonia on Utilities Requirements 

0.0%

10.0%

20.0%

30.0%

40.0%

50.0%

% Increase (as compared to feed 

stream with no ammonia)

Reboiler Duty

Condenser Duty

The MDEA unit utilizing a generic MDEA solution (Case 1) was used as the reference case for 

the sensitivity analysis to determine the impact of varying acid gas concentrations in the feed 

stream on the utilities requirements of the unit. This analysis assumes that the volumetric flow 

rate of the feed gas is the same as the reference case, and only the H2S and CO2

concentrations are varied by ±10%. Figure 5-6 illustrates the percentage change in utilities 

requirements, relative to the reference case, as the H2S and CO2 concentrations in the feed gas 

are varied. The analysis results show that when the acid gas concentrations are increased by 

10%, the utilities requirements of the unit increase between 9-11%. When the acid gas 

concentrations are decreased by 10%, the impact on power consumption, with reduction of 

about 16%, is considerable; however, the impact on steam and cooling water circulation 

requirements are less pronounced. This is most likely due the effect of ammonia on the reboiler 

and condenser duties, as shown above. The results show no change in make-up water 

requirement when the acid gas concentration decreases. 

Figure 5-6 Effects of Varying Acid Gas Concentrations on Utilities Requirements 

-20.0%

-15.0%

-10.0%

-5.0%

0.0%

5.0%

10.0%

15.0%

% Change, (relative to Reference Case)

Amine Circulation

Power Consumption

Steam Consumption

Make-up Water

CW Circulation

Effects of 10% 

increase in acid 

gas conc. 

Effects of 10% 

decrease in 

acid gas conc. 

Task 2: Detailed MDEA Process Analysis 

United States Department of Energy/National Renewable Energy Laboratory 

26 
 

Task 2: Detailed MDEA Process Analysis 

United States Department of Energy/National Renewable Energy Laboratory 

27 

Section 6 Summary and Conclusions

This study conducted a detailed analysis of the MDEA system remove acid gases from the 

NREL biomass-derived syngas. The MDEA system is designed to treat approximately 163 

MMSCFD of syngas containing 234 ppmv H2S and 11.4 mol% CO2. The required purity of the 

treated gas, which is dictated by the sulfur requirement and CO2 tolerance of the alcohol 

synthesis catalyst, are 10 – 50 ppmv H2S and 1 - 5 mol% CO2. This study evaluated two types 

of solution the MDEA system: a generic 50 wt% MDEA solution (Case 1) and an activated 

MDEA solution consisting of a mixture of 50 wt% MDEA and 2% piperazine (Case 2). 

The analysis showed that the presence of ammonia in the feed stream has a major impact on 

the utilities requirements of the MDEA unit. Moreover, the presence of ammonia eventually 

leads to buildup in the stripper overhead system and contributes to corrosion. For these 

reasons, the removal of ammonia upstream of the AGR unit is critical to minimize operating 

problems as well as operating costs. 

The analysis results indicated that for the NREL thermochemical ethanol synthesis application, 

in which the syngas to be treated in the MDEA unit has appreciably higher concentration of CO2

relative to H2S, the use of a piperazine promoter in the MDEA solution to enhance the rate of 

CO2 absorption is very favorable. Due to the enhanced CO2 absorption capability of the amine 

solution, the MDEA unit requires less amine circulation rate and fewer absorber stages than one 

using a generic 50 wt% solution (Case 1). Consequently, due to the smaller plant size, the total 

installed equipment cost is approximately 16% lower compared to Case1. Additionally, the 

operating costs of the MDEA unit in Case 2 are 30-45% lower compared to Case 1, primarily 

due to lower DMDS requirement for controlling the H2S content of the clean gas at appropriate 

levels in order to maintain the alcohol synthesis catalyst activity. Thus, for the NREL design, the 

MDEA system utilizing an activated MDEA solution has more economics advantages over the 

one utilizing a generic MDEA solution. 
 

Task 1: Acid Gas Removal Technology Survey and Screening 

United States Department of Energy/National Renewable Energy Laboratory 

28 

 References

1. Nexant Inc. Task 1: Acid Gas Removal Technology Survey and Screening for 

Thermochemical Ethanol Synthesis. Subcontract Report NREL/KAFT-8-882786-01. 

March 2009. 

2. Phillips, S., Aden, A., Jechura, J., Dayton, D., and Eggeman, T., “Thermochemical 

Ethanol via Indirect Gasification and Mixed Alcohol Synthesis of Lignocellulosic 

Biomass,” NREL Report No. TP-510-41168, April 2007. 

3. Kohl, A. L., and Nielsen, R. B., Gas Purification, 5th ed., Gulf Publishing, Houston, TX, 

1997. 

4. Bishnoi, S. and G. T. Rochelle, "Absorption of Carbon Dioxide in Aqueous 

Piperazine/Methyldiethanolamine," AIChE Journal, Vol. 48, No. 12, December 2002, pp. 

2788 - 2799. 
 

Task 1: Acid Gas Removal Technology Survey and Screening 

United States Department of Energy/National Renewable Energy Laboratory 

29 

Appendix A Case 1
Task 1: Acid Gas Removal Technology Survey and Screening 

United States Department of Energy/National Renewable Energy Laboratory 

Table A-1 Case 1 – Detailed Utilities Requirements 

Load BHP Elect. Power 

Steam 

1000 lbs/hr 

Water Requirement, 

1000 lbs/hr 

Item No Item Name Norm. Max. KW 50 psig / 298 F 

Return Cond. 

50 Psig / 298 F Make-u

C-301 Acid Gas Absorber 2E-301A/B Lean Amine Cooler E-303 A/B Stripper Condenser E-304 Stripper Reboiler 119.0 (119.0)G-301 A/B Lean Amine Pump 663.6 549.8G-302 A/B Reflux Pump 6.7 5.2G-303 A/B Makeup Water Pump 16.9 14.2G-304 Amine Sump Pump (Note 2) 0.5 0.4G-305 A/B DMDS Injection Pump 0.1 0.1

TOTAL 687 569 119.0 (119.0) 2

NOTES:

(1) Negative numbers (xxx) are productions while positive numbers are consumptions. (2) Numbers shown are intermittent consumptions and are NOT included in NORMAL total.
Table A-2 Case 1 - Equipment List

Columns Item Name Quantity Type I.D (FT) Temp (°F) Shell C-301 Acid Gas Absorber 1 Vertical 10.0 220 CS & SR 

 Valve Trays 32 Internal Devices 10.0 

 Liquid Distributors 2 Internal Devices 10.0  Demister Pads 1 Internal Devices 10.0 

C-302 Acid Gas Stripper 1 Vertical 10.5 310 CS/304 SS Clad

 Valve Trays 20 Internal Devices 10.5 

 Chimney Trays 1 Internal Devices 10.5  Liquid Distributors 2 Internal Devices 10.5  Demister Pads 1 Internal Devices 10.5 

Vessels & Tanks Item Name Quantity Type I.D (FT) Temp (°F) M.O.C 

D-301 Treated Gas Knockout Drum 1 Vertical 7.0 180 CS De

D-302 Flash Drum 1 Horizontal 9.0 210 CS & SR

 Flash Drum Top Column 1 Vertical 3.0 210 CS & SR De

D-303 Reflux Drum 1 Vertical 5.5 170 CS & SR De

T-301 MDEA Surge Tank 1 Cone Roof 15.0 310 CS

T-302 MDEA Sump 1 In Ground Concrete 

Total Duty Tot. Bare Tube Area

Exchangers Item Name Quantity Type MMBTU/HR Temp (°F) M.O.C

S 80 240 CS & SR

T 480 240 CS & SR

S 30 310 304 SS

T 50 310 304 SS

S 30 280 304 SS

T 80 280 304 SS

S 30 350 304 SS

T 70 350 304 SS

Pumps Item Name Quantity Type Flow (GPM) Inlet PSIG Delta PSI Driver HP Casing 

G-301 A/B Lean Amine Pump 2 Cent 1,715 3.6 432 737 CS 

G-302 A/B Reflux Pump 2 Cent 130 5.6 14 7 CS 

G-303 A/B Makeup Water Pump 2 Cent ATM 435 44 19 CS 

G-304 Amine Sump Pump 1 Vert. Cent ATM 10 17 1 CS 

G-305 A/B DMDS Injection Pump 2 Meter 0.08 43 942 0.07 CS 

Miscellaneous Item Name Quantity Temp (°F) M.O.C

Y-301 Filters (Carbon + Pre & Post Filters) 1 Package 153.40 50 310 CS & SR 

Mat'l of Construction 

M.O.C

50

Design

Press. (PSIG)

17

7

E-304 Stripper Reboiler 1 Kettle 109.67

E-303 A/B Stripper Condenser 2 S & T 63.48

E-302 A/B/C/D Lean/Rich Exchangers 4 S & T 51.83

Design

Press. (PSIG)

E-301A/B Lean Amine Cooler 2 S & T 41.59

SQ FT

7,660

20

ATM

Press. (PSIG)

440

50

Design

50

DimensionsDimensions

Design

Press. (PSIG)

450

T-T (FT)

80.0

N/AN/AN/A

66.0

N/AN/AN/AN/A

T-T (FT)

8.0

22.0

6.0

10.046.0

30,400

8,288

14,260

Design Capacity Power

Pump BHP

664

GPM of lean MDEA

Design Capacity

0.5

0.07

Task 1: Acid Gas Removal Technology Survey and Screening 

United States Department of Energy/National Renewable Energy Laboratory 
Task 1: Acid Gas Removal Technology Survey and Screening 

United States Department of Energy/National Renewable Energy Laboratory 

Table A-3 Case 1 - Equipment Cost Estimates

 COSTS IN US$ (2Q09) 

DESCRIPTION QTY UNIT EQUIPMENT BULK LABOR 

SUBC

OTH

PROCESS EQUIPMENT & DUCTWORKPUMPS & DRIVERS 10 EA 503,200 37,000 HEAT EXCHANGERS 9 EA 2,017,591 23,000 TANKS & STORAGE FACILITIES 1 EA 53,500 2,000 MATL HANDLING & PROCESSING - FILTERS 1 EA 134,500 5,000 COLUMNS & VESSELS (INCL HEAVY/OVERSIZE TRANSPORT) 5 EA 1,401,770 66,500 FREIGHT (EQUIPMENT EXCEPT HEAVY/OVERSIZE TRANSPORT) 5 % 205,500 TOTAL PROCESS EQUIPMENT & DUCTWORK 26 EA 4,316,061 133,500 INSTRUMENTS 647,400 252,500 PIPING 1,294,800 1,431,000 STEELWORK 259,000 185,000 INSULATION 4ELECTRICAL 647,400 799,500 CONCRETE 215,800 308,500 BUILDING SITEWORK 215,800 449,000 PAINTING TOTAL DIRECT COSTS 3,280,200 3,425,500 4TOTAL INSTALLED EQUIPMENT COSTS 4,316,061 3,280,200 3,559,000 4
 

Task 1: Acid Gas Removal Technology Survey and Screening 

United States Department of Energy/National Renewable Energy Laboratory 

33 

Appendix B Case 2
Task 1: Acid Gas Removal Technology Survey and Screening 

United States Department of Energy/National Renewable Energy Laboratory 

Table B-1 Case 2 – Detailed Utilities Requirements 

Load BHP 

Elect. 

Power 

Steam 

1000 lbs/hr 

Water Requirement, 

1000 lbs/hr 

Item No Item Name Norm. Max. KW 

50 psig / 

298 F 

Return 

Cond. 50 

Psig / 298 F Make-u

C-301 Acid Gas Absorber 2E-301A/B Lean Amine Cooler E-303 A/B Stripper Condenser E-304 Stripper Reboiler 119.3 (119.3) G-301 A/B Lean Amine Pump 574.1 475.9 G-302 A/B Reflux Pump 7.0 6.0 G-303 A/B Makeup Water Pump 18.3 14.9 G-304 Amine Sump Pump (Note 2) 0.8 0.7 G-305 A/B DMDA Injection Pump 0.04 0.03 

TOTAL 600 497 119.3 (119.3) 2

NOTES:

(1) Negative numbers (xxx) are productions while positive numbers are consumptions. (2) Numbers shown are intermittent consumptions and are NOT included in NORMAL total. 
Table B-2 Case 2 - Equipment List

Columns Item Name Quantity Type I.D (FT) Temp (°F) Shell C-301 Acid Gas Absorber 1 Vertical 10.0 220 CS & SR 

 Valve Trays 12 Internal Devices 10.0  Liquid Distributors 2 Internal Devices 10.0 

 Demister Pads 1 Internal Devices 10.0 

C-302 Acid Gas Stripper 1 Vertical 10.5 310 CS/304 SS Clad

 Valve Trays 20 Internal Devices 10.5  Chimney Trays 1 Internal Devices 10.5 

 Liquid Distributors 2 Internal Devices 10.5 

 Demister Pads 1 Internal Devices 10.5 

Vessels & Tanks Item Name Quantity Type I.D (FT) Temp (°F) M.O.C 

D-301 Treated Gas Knockout Drum 1 Vertical 7.0 180 CS Dem

D-302 Flash Drum 1 Horizontal 9.0 210 CS & SR

 Flash Drum Top Column 1 Vertical 3.0 210 CS & SR Dem

D-303 Reflux Drum 1 Vertical 5.5 170 CS & SR Dem

T-301 MDEA Surge Tank 1 Cone Roof 14.0 310 CST-302 MDEA Sump 1 In Ground Concrete 

Total Duty Tot. Bare Tube Area

Exchangers Item Name Quantity Type MMBTU/HR Temp (°F) M.O.C

S 80 240 CS & SR

T 480 240 CS & SR

S 30 310 304 SS

T 50 310 304 SS

S 30 280 304 SS

T 80 280 304 SS

S 30 350 304 SS

T 70 350 304 SS

Pumps Item Name Quantity Type Flow (GPM) Inlet PSIG Delta PSI HP Casing 

G-301 A/B Lean Amine Pump 2 Cent 1,535 2.60 433 638 CS 

G-302 A/B Reflux Pump 2 Cent 5.60 138 14 8 CS G-303 A/B Makeup Water Pump 2 Cent ATM 435 49 20 CS 

G-304 Amine Sump Pump 1 Vert. Cent ATM 10 29 1 CS 

G-305 A/B DMDA Injection Pump 2 Meter 0.05 43 942 0.04 CS 

Miscellaneous Item Name Quantity Temp (°F) M.O.C

Y-301 Filters (Carbon + Pre & Post Filters) 1 Package 140 50 310 CS & SR 

Mat'l of Construction 

M.O.C

50

Design

Press. (PSIG)

18

7

E-304 Stripper Reboiler 1 Kettle 109.9

0.04

E-303 A/B Stripper Condenser 2 S & T 63.3E-302 A/B Lean/Rich Exchangers 2 S & T 44.8

Design

Press. (PSIG)

E-301A/B Lean Amine Cooler 2 S & T 36.8

SQ FT

7,290

20

ATM

Press. (PSIG)

440

50

Design

50

DimensionsDimensions

Design

Press. (PSIG)

450

T-T (FT)

39.5

N/AN/AN/A

65.5

N/AN/AN/AN/A

T-T (FT)

8.0

19.0

6.0

10.0

45.0

26,500

9,100

15,325

Design Capacity Driver

Pump BHP

574

GPM of lean MDEA

Design Capacity

1

Task 1: Acid Gas Removal Technology Survey and Screening 

United States Department of Energy/National Renewable Energy Laboratory 
Task 1: Acid Gas Removal Technology Survey and Screening 

United States Department of Energy/National Renewable Energy Laboratory 

Table B-3 Case 2 - Equipment Cost Estimates

 COSTS IN US$ (2Q09) 

DESCRIPTION QTY UNIT EQUIPMENT BULK LABOR 

SUBC

OTH

PROCESS EQUIPMENT & DUCTWORKPUMPS & DRIVERS 10 EA 476,600 36,000 HEAT EXCHANGERS 7 EA 1,778,400 23,000 TANKS & STORAGE FACILITIES 1 EA 48,533 2,000 MATL HANDLING & PROCESSING - FILTERS 1 EA 136,500 5,000 COLUMNS & VESSELS (INCL HEAVY/OVERSIZE TRANSPORT) 5 EA 996,590 36,500 FREIGHT (EQUIPMENT EXCEPT HEAVY/OVERSIZE TRANSPORT) 5 % 171,800 TOTAL PROCESS EQUIPMENT & DUCTWORK 24 EA 3,608,423 102,500 INSTRUMENTS 541,300 211,000 PIPING 1,082,500 1,196,000 STEELWORK 216,500 155,000 INSULATION 3ELECTRICAL 541,300 668,500 CONCRETE 180,400 258,000 BUILDING SITEWORK 180,400 375,000 PAINTING TOTAL DIRECT COSTS 2,742,400 2,863,500 3TOTAL INSTALLED EQUIPMENT COSTS 3,608,423 2,742,400 2,966,000 3
F1146-E(10/2008) 

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1. REPORT DATE (DD-MM-YYYY)

May 2011 

2. REPORT TYPE

Subcontract Report

3. DATES COVERED (From - To)

 

4. TITLE AND SUBTITLE

Survey and Down-Selection of Acid Gas Removal Systems for the 

Thermochemical Conversion of Biomass to Ethanol with a Detailed 

Analysis of an MDEA System

Task 1: Acid Gas Removal Technology Survey and Screening for 

Thermochemical Ethanol Synthesis; March 2009

Task 2: Detailed MDEA Process Analysis; September 2009

5a. CONTRACT NUMBER

DE-AC36-08GO28308

5b. GRANT NUMBER

5c. PROGRAM ELEMENT NUMBER

6. AUTHOR(S)

Nexant Inc.

5d. PROJECT NUMBER

NREL/SR-5100-50482

5e. TASK NUMBER

BB07.3710

5f. WORK UNIT NUMBER

7. PERFORMING ORGANIZATION NAME(S) AND ADDRESS(ES)

Nexant Inc.

101 2nd St # 1000

San Francisco, CA 94105-3651

8. PERFORMING ORGANIZATION

REPORT NUMBER

KAFT-8-882786-01

9. SPONSORING/MONITORING AGENCY NAME(S) AND ADDRESS(ES)

National Renewable Energy Laboratory

1617 Cole Blvd.

Golden, CO 80401-3393

10. SPONSOR/MONITOR'S ACRONYM(S)

NREL

11. SPONSORING/MONITORING

AGENCY REPORT NUMBER

NREL/SR-5100-50482

12. DISTRIBUTION AVAILABILITY STATEMENT

National Technical Information Service

U.S. Department of Commerce

5285 Port Royal Road

Springfield, VA 22161

13. SUPPLEMENTARY NOTES

NREL Technical Monitor: Andy Aden

14. ABSTRACT (Maximum 200 Words)

The first section (Task 1) of this report by Nexant includes a survey and screening of various acid gas removal 

processes in order to evaluate their capability to meet the specific design requirements for thermochemical ethanol 

synthesis in NREL’s thermochemical ethanol design report (Phillips et al. 2007, NREL/TP-510-41168). MDEA and 

selexol were short-listed as the most promising acid-gas removal agents based on work described in Task 1. The 

second report section (Task 2) describes a detailed design of an MDEA (methyl diethanol amine) based acid gas 

removal system for removing CO2 and H2S from biomass-derived syngas. Only MDEA was chosen for detailed study 

because of the available resources.

15. SUBJECT TERMS

acid gas removal; AGR; thermochemical; biomass; biofuels; ethanol; MDEA; syngas; nexant.

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