H2S REMOVAL FROM CO2 BY DISTILLATION

Galip H. Guvelioglu, Ph.D. Paul Higginbotham, Ph.D.

Air Products and Chemicals, Inc. Air Products and Chemicals, Inc.
7201 Hamilton Blvd. Molesey Road, Hersham
Allentown, PA 18062 USA Walton-on-Thames
+1 610-481 -5926 Surrey KT12 4RZ Great Britain
guvelig@airproducts.com +44 1932-249-506
higginp@airproducts.com

John E. Palamara, Ph.D. Gaurav Arora, Ph.D.

Air Products and Chemicals, Inc. Air Products and Chemicals, Inc.
7201 Hamilton Blvd. 7201 Hamilton Blvd.
Allentown, PA 18062 USA Allentown, PA 18062 USA
+1 610-481 -1108 +1 610-481 -1543
palamaje@airproducts.com arorag@airproducts.com

Darryl L. Mamrosh, P.E. Kevin S. Fisher, P.E.

Trimeric Corporation Trimeric Corporation
100 South Main Street 100 South Main Street
Buda, TX 78610 USA Buda, TX 78610 USA
+1 512-434 -9912 +1 512-431 –6323
darryl.mamrosh@trimeric.com kevin.fisher@trimeric.com

Abstract
This paper introduces a new process configuration developed by Air Products and Chemicals,
Inc. to efficiently separate carbon dioxide (CO2) and hydrogen sulfide (H2S) using traditional
distillation methods without needing an additive to increase the relative volatility of CO2 and
H2S. The removal of H2S from CO2 is of high importance in industries such as Enhanced Oil
Recovery where purified CO2 is often preferred; however the separation of H2S from CO2 by
distillation is technically challenging and energy-intensive due to the low relative volatility at
low H2S concentrations. Traditionally, distillation has been considered technically feasible but
deemed to have high capital and energy requirements. This paper describes the application of
advances in distillation technology, heat exchanger design, and process optimization to lower the
cost of this challenging separation even in cases where nitrogen, methane, natural gas liquids
(NGL) or other components are present. Many of these advances are currently applied at large
scale in other challenging distillation separations, such as O2/Ar in the air separation industry.
The advantages of this new distillation process for H2S and CO2 will be demonstrated through
comparison with conventional solvent-based processes that have been applied to the separation
of H2S from CO2.

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 H2S REMOVAL FROM CO2 BY DISTILLATION

Galip H. Guvelioglu, Ph.D. Air Products and Chemicals, Inc. Allentown, PA USA
Paul Higginbotham, Ph.D. Air Products and Chemicals, Inc. Hersham Great Britain
John E. Palamara, Ph.D. Air Products and Chemicals, Inc. Allentown, PA USA
Gaurav Arora, Ph.D. Air Products and Chemicals, Inc. Allentown, PA USA
Darryl L. Mamrosh, P.E. Trimeric Corporation. Buda, TX USA
Kevin S. Fischer, P.E. Trimeric Corporation. Buda, TX USA

I. Introduction
A growing fraction of today’s oil is produced via Enhanced Oil Recovery (EOR) using carbon
dioxide (CO2) [1]. CO2 EOR requires large quantities of CO2, and sources capable of supplying
these requirements can be grouped into a few categories:

• Natural sources that are mostly CO2

• Recycle from EOR
• Byproduct from natural gas processes
• Byproduct from industrial processes
• Byproduct from power generation

Some of these CO2 sources can contain hydrogen sulfide (H2S) in amounts that can be
considered higher than desired, e.g., greater than 0.5 mol%. Hydrogen sulfide is not particularly
detrimental to the efficiency of CO2 in EOR service, however it is often removed prior to CO2
transmission pipelines due to various reasons, such as its toxicity [2]. The need for the H2S
removal process is situational, depending upon location, regulations, public perception, safety,
concerns for H2S buildup in the reservoir, the quality of oil to be extracted or pipeline
requirements.

One of the most significant issues that may force the need for removing H2S from EOR CO2 is
pipeline requirements. Many (but not all) CO2 supply pipelines that serve EOR areas have a
specification for H2S; a range of 10 – 100 ppmv is common. Therefore, in situations where CO2
that contains significant amounts of H2S needs to be transported via pipeline for a significant
distance, H2S removal is often required.

A. Overview of the Conventional H2S Removal Technologies

The choice of process technology for removal of H2S from CO2 streams in EOR applications
depends on several factors and varies greatly with CO2 to H2S ratio, size of the plant, presence of
other components in the gas, feed pressure and constraints on the disposition of the H2S. In turn,
these factors determine the capital and operating costs [3]. Similar to a common rule-of-thumb
for the removal of H2S from natural gas, the total amount of elemental sulfur that needs to be
removed is a good indicator for screening the suitability of alternative technologies.

Typically, for natural gas streams with low H2S levels, e.g., where the total amount of H2S to be
removed is significantly less than 1 LTPD of elemental sulfur, H2S scavengers are usually found

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to be a good techno-economic fit. As the amount of H2S to be removed from the gas increases to
a medium scale of roughly 1 to 20 LTPD, liquid redox type processes and some other
technologies have been applied. For the larger scale of more than about 20 LTPD, amine units
with Claus units are frequently selected for removal and disposal of H2S [3].

For the removal of H2S from gas streams that are primarily CO2, the application of conventional
selective amine technologies tend to be more difficult due to the fact that amines cannot achieve
perfect selectivity for H2S removal in the presence of CO2. Therefore treatment of streams that
are primarily CO2 tends to be more difficult and the processes applied more complex compared
to the treatment of natural gas, particularly at large scale. The concentration of H2S in the CO2
feed stream, and the concentration to which it must be removed to in the treated gas are also very
important in the selection of process technology.

For larger scale applications requiring the removal of more than about 20 LTPD of H2S from a
gas stream that is primarily CO2, amines, physical solvents (e.g., DEPG), conventional
distillation, extractive distillation (e.g., Ryan-Holmes), or combinations of these technologies
may be preferred. All of these technologies effect the separation of H2S from the CO2, resulting
in the production of a hydrogen sulfide-rich acid gas byproduct stream. The hydrogen
sulfiderich byproduct stream may then be converted to sulfur via the Claus process, or may be
injected. Acid gas injection may be practiced preferentially in conjunction with CO2 treatment
for EOR.

For large CO2 sources that contain relatively high concentrations of H2S (greater than 0.5 mol%)
the equivalent amount of elemental sulfur that needs to be removed can be an order of magnitude
larger than 20 LTPD, which is already considered large scale. For example, for a CO2 source at
a flow rate 100 MMSCFD that contains 1% H2S, the amount of elemental sulfur that needs to be
removed is approximately 226 LTPD. As the H2S concentration in the CO2 source increases
above 5 mol%, the process technologies that can economically remove H2S to meet the CO2
pipeline requirements are limited. For high H2S concentration CO2 sources, solvent-based Acid
Gas Removal (AGR) processes like amines [3] and physical solvents have typically been used to
remove H2S from gas streams [4][5]. When amines are applied for the absorption of H2S from
CO2, an Acid Gas Enrichment (AGE) unit may also be required because the H2S concentration of
the AGR regeneration gas may not be sufficient (due to selectivity limitations) to allow for the
economical operation of a Claus unit or acid gas injection.

Purification of EOR-grade carbon dioxide, where the hydrogen sulfide-rich byproduct stream is
prepared for acid gas injection with greater than 50 mol% hydrogen sulfide content, has been
accomplished commercially by solvent-based Acid Gas Removal and extractive distillation
processes such as Ryan-Holmes [6].

Solvent-based AGR systems, such as Dimethyl Ethers of Polyethylene Glycol (DEPG) have
been commercially proven in removal of acid gases such as carbon dioxide and hydrogen sulfide
from natural gas sources or syngas. While most of the physical solvents have an order of
magnitude higher affinity for H2S than for CO2[7], they still co-absorb significant quantities of
CO2. The CO2 that is co-absorbed during H2S removal requires a series of lower pressure flashes
and/or a hydrogen sulfide enrichment column in order to produce a hydrogen sulfide stream from
the solvent regenerator that does not contain excessive amounts of CO2. The flashed CO2 then
has to be compressed and recycled back into the feed of the absorber in order to achieve high

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carbon dioxide recovery and/or for environmental and safety considerations. This leads to higher
energy consumption and capital requirement compared to an equivalent system without the low
pressure flashes.

B. Overview of CO2/H2S Distillation Processes

The vapor liquid equilibrium of the CO2 and H2S binary system has been studied extensively.
Bierlein et al. [8] concluded that while no azeotrope existed for CO2 and H2S, evidence of strong
intermolecular forces existed at high CO2 concentrations. While the traditional distillation of
CO2 and H2S is technically feasible, it has been regarded as a difficult separation due to the low
relative volatility of CO2 and H2S at high CO2 concentrations. Traditionally, this would dictate
large columns with high energy requirements to meet the purity requirements typically dictated
by EOR use [9].

Ryan-Holmes processes employ distillation to effect the CO2-H2S separation, however they
differ from traditional distillation in that a hydrocarbon-based additive like n-butane or a mixture
containing n-butane is used to increase the relative volatility of CO2 with respect to H2S. Figure
1 shows the CO2/H2S relative volatility at 600 psia for a CO2 and H2S binary system and
CO2/H2S/n-butane ternary system.

Figure 1 - CO2 - H2S - nC4 System at 600 psia[6]

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For the CO2 and H2S binary system, previously reported CO2/H2S relative volatilities at CO2 rich
concentrations are approximately 1.42-1.45 at 600 psia [8]. With the addition of n-butane, the
relative volatility of CO2 and H2S can be increased to 1.8-2.0 thus reducing the number of stages
required, reflux requirements and energy intensity of the H2S/CO2 separation. The Ryan-Holmes
process is more suited to applications where the additives used for improving CO2/H2S relative
volatility are already present within the feed gas. These additives are typically an n-C4+ cut of
Natural Gas Liquids (NGL). In the Ryan-Holmes process, the increase in relative volatility can
be controlled by recycling more or less of the additive. The Ryan-Holmes process is employed
commercially in high CO2 gas fields or EOR facilities that recover NGL from produced gas and
recycle carbon dioxide [9]. While the Ryan-Holmes process reduces the energy requirements for
the separation of CO2 and H2S by distillation, it requires additional fractionation steps to recover
the additives from the H2S stream. This additive recovery requirement and the plant
infrastructure needed to store and recycle these additives increases the capital and energy
intensity of the Ryan-Holmes process. In applications where the feed gas does not contain
enough heavy hydrocarbons to balance the process additive losses, the Ryan-Holmes process
requires purchase of these additives, thus increasing the operating cost.

While processes for the removal of H2S from CO2 by distillation have been developed and
studied, no known commercial plants have been built. Pryor [10] describes a process for
removal of CO2 and H2S from a hydrogen-rich gas by condensation, where the stream of
condensed CO2 and H2S then feeds into a distillation column and is separated into a CO2 rich
overhead and a bottoms stream containing at least 10 vol% H2S. Overhead vapor from the
column is condensed using an external closed-loop propane refrigerant and bottoms liquid is
reboiled using process cooling water. The distillation column has 100 trays and operates at about
590 psia.

C. Distillation of Mixtures with Low Relative Volatility

While a relative volatility of approximately 1.4 for CO2 at rich CO2 concentrations indicates a
difficult separation requiring a large number of theoretical stages and high reflux ratios, more
difficult separations have been performed commercially when the economics support them, e.g.,
in cryogenic air separation units to separate oxygen and argon. The argon/oxygen system has a
relative volatility ranging from 1.05 to 1.1 at dilute oxygen concentration. A typical commercial
argon product specification is less than 1 ppmv oxygen content. 180 to 220 theoretical stages
are required to achieve the separation. The oxygen/argon separation energy intensity and capital
requirements have been reduced significantly since the 1990s with advances in distillation
technology using advanced heat integration techniques and structured packing distillation
column internals[11][12].

Although the high energy intensity of the CO2 and H2S separation by distillation cannot be
eliminated, the efficiency of the separation process can be maximized with advanced process
configurations tailored for systems containing CO2. Similar configurations have been studied for
the purification of CO2 following capture from power plants. The quantities of CO2 used in EOR
operations are similar to CO2 flow rates planned in the Carbon dioxide Purification Units (CPU)
of oxyfuel power plants in the 200-300 MW power plant size range.

The technical advances in auto-refrigerated CO2 purification processes are the foundation for a
first-of-a-kind commercial helium recovery plant where CO2 makes up more than 95% of the

335
 feed gas. It is in the final stages of construction at the time of writing this paper [13]. The plant
will be treating 170 MMSCFD of CO2, the scale associated with EOR operations.

II. The New Process for H2S Removal from CO2 by Distillation
Building on these advances, Air Products has developed a novel process that separates CO2 and
H2S efficiently using traditional distillation methods without needing to use an additive to
increase the relative volatility of CO2 and H2S.

The process flow diagram of this new configuration is shown in Figure 1 [14]. The feed in this
example is primarily a CO2/H2S mixture in the liquid phase at approximately 650 psia. These
feed conditions could be obtained as a result of processes that rely on the condensation or
freezing of CO2 and/or H2S [15] [16] [4] from low quality natural gas resources. Liquid feed is
not a requirement; alternatively the feed can be gaseous CO2 and H2S as typically seen in
recycled CO2 from sour EOR fields.

The process includes a multi-stream Brazed Aluminum Heat Exchanger (BAHX), a CO2 recycle
compressor and aftercooler, a CO2/H2S distillation column and product pressurization pumps.

The liquid feed is first subcooled in the BAHX then reduced in pressure via a feed Joule
Thompson (JT) valve. The resulting two-phase feed stream is fed to the BAHX to be warmed up
and vaporized and fed to the distillation column. The subcooling of the feed before let down in
pressure increases the utilizable refrigeration that is provided when the feed stream evaporated,
improving the efficiency of the process. The distillation column operates at approximately 300
psia. The main light component in this distillation process is CO2 and the purified CO2 product
is obtained from the top of the distillation column. The top section of the column rectifies CO2
and is designed to meet the specification of H2S in the CO2 for EOR applications, typically less
than 100 ppmv of H2S. The heavier component in this separation is H2S and this is obtained at
the bottom of the column. The bottom of the column strips CO2 from the H2S to produce a H2S
stream that contains greater than 80% H2S at the column pressure of 300 psia; this H2S by-
product would usually be sent to a Claus unit or to acid gas injection. In contrast to solvent-
based acid gas separation processes, both the CO2 and the H2S are kept at elevated pressure
throughout the process and this significantly reduces the product pumping or compression
requirements.

The overhead of the column is warmed up in the BAHX, compressed in the recycle compressor
to approximately 450 psia and then sent to the aftercooler. The higher pressure gaseous recycle
CO2 stream is then condensed and subcooled in the BAHX against the side and bottom reboilers
of the distillation column. The now subcooled CO2 recycle stream is reduced in pressure to
approximately 350 psia in the condenser separator. The condenser separator contains a liquid
CO2 stream that contains less than 100 ppmv of H2S which is then split into two streams; a
portion of it is sent to the distillation column as reflux and the remaining portion is pumped
above critical pressure via a CO2 booster pump. By keeping the condenser separator sufficiently
above the column pressure, reflux can be sent to the top of the column without the need for
reflux pumps.

The CO2 after the CO2 booster pump is warmed up in the BAHX and pumped to the final carbon
dioxide pipeline pressure of 1,500 - 2,500 psia via the CO2 product pump.

336
This type of process is typically referred to as an auto-refrigerated configuration since it does not
require any additional external refrigeration. The condenser and reboilers of the CO2/H2S
distillation column are configured as a heat pump circuit; vaporizing CO2 in the reboilers
provides the refrigeration for the recycle CO2 condensing at higher pressure.

In this configuration, high CO2 recoveries in excess of 99% can be achieved with a 5 mol% H2S
feed and 85 mol% H2S in the waste H2S stream. The CO2 losses will be primarily due to CO2
leaving with the waste H2S stream and, to a lesser extent, losses that typically occur in
compressor and pump seals.

RECYCLE
COMPRESSOR

Feed
Stream

DISTILLATION
COLUMN
BAHX
NETWORK
Purified
Carbon Dioxide

PRODUCT
PUMP

Concentrated
Hydrogen Sulfide

INJECTION
PUMP

CO2 (<100 ppm H2S) CONDENSER

SEPARATOR
CO2 + H2S

H2S (>80% H2S) BOOSTER

PUMP

Figure 2 - Process Flow Diagram

III. Case Study

The following case study has been developed to demonstrate the performance of this process. In
this example, a feed stream needs to be purified in order to meet the typical pipeline
requirements of carbon dioxide for EOR. The feed in this example could be from another
process that condenses CO2 and H2S from a gas reservoir [15][4].

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 Table 1 - Feed and Product Requirements

Feed CO2 Product Waste H2S

Pressure, psia 650 2,500 3,500
Temperature, F 45 < 120 < 120
Flow Rate, MMSCFD* 100
Molar Composition
CO2 92.725% > 95% < 15%
H2S 6.0% < 25 ppmv > 85%
COS 250 ppmv
CH4 1.0% < 5%
N2 0.25% < 5%
Notes:
* MMSCF is 1,000,000 Standard Cubic Feet (SCF) at 14.696 psia and 60°F.

The separation of the liquid phase feed stream into a CO2 product and waste H2S was
characterized using two different distillation process configurations: external refrigeration and
auto-refrigerated processes. The power requirements and energy intensity of these two process
configurations are compared in this section.

Both processes were simulated with Aspen Technology Inc.’s Aspen Plus® simulation software
using a customized equation of state model. The CO2 and H2S parameters for the equation of
state model were obtained through Vapor Liquid Equilibrium (VLE) laboratory measurements
performed for the pressure, temperature and composition ranges of interest.

The CO2 and H2S distillation column in each process was simulated with 80 theoretical stages
plus a condenser and a bottom reboiler. Air coolers were used for the process cooling with a
design ambient air temperature of 90 °F and a minimum temperature approach of 20 °F. In both
cases, a polytropic efficiency of 80% was used for the compressors.

A. External Refrigeration Distillation Process Configuration

Figure 3. The heat exchangers used in this process configuration are shown as a single heat
exchanger network for the example case. This heat exchanger network can be split into multiple
heat exchangers depending on the physical equipment requirements and economic consideration.
The feed is subcooled in the heat exchanger network and then let down to approximately 300
psia, evaporated against subcooling refrigerant, and fed to the CO2/H2S distillation column
towards the bottom.

The distillation column features a total condenser and bottom reboiler that are integrated into the
heat exchanger network. The condenser separator contains a liquid CO2 stream that contains less
than 25 ppmv of H2S, which is then split into two streams; a portion of it is sent to the distillation
column as reflux, and the remaining portion is pumped above critical pressure via a CO2 booster
pump. The CO2 after the CO2 booster pump is warmed up in the heat exchanger network and
pumped to the final CO2 pipeline pressure of 2,500 psia via the CO2 product pump. The

338
refrigeration for the process is provided with a single stage propane refrigeration system with
subcooling. The propane compressor is a two stage compressor with air cooled intercooler and
aftercooler. The compressed and cooled propane is then condensed against the column reboiler
and subcooled against evaporating feed in the heat exchanger network.

The waste H2S product stream with greater than 85% H2S is obtained from the bottom of the
column as liquid and pumped to a final pressure of 3,500 psia using the acid gas injection pump.
REFRIGERANT
AFTERCOOLER
COMPRESSOR

Feed
Stream

DISTILLATION
COLUMN
HX
NETWORK
Purified
Carbon Dioxide

PRODUCT
PUMP

Concentrated
Hydrogen Sulfide

INJECTION
PUMP

CO2 (<25 ppm H2S) CONDENSER

SEPARATOR
CO2 + H2S

H2S (>80% H2S) BOOSTER REFLUX

PUMP PUMP
Refrigerant

Figure 3 – External Refrigeration Distillation Process Configuration

A summary of process power consumption is shown in Table 2.

Table 2 - Externally Refrigerated Process Power Summary

Power Consumption, hp
Refrigeration Compressors 27,000
Waste H2S Injection Pump 190
CO2 Booster Pump 560
CO2 Pipeline Pump 1,520
Total Power 29,270

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The propane condenser integrated with the reboiler of the column in the heat exchanger network
reduces the condensing pressure of the propane to approximately 90 psia compared to
condensing against air which would have required 220 psia, thereby reducing the propane
compressor power.

B. Auto Refrigerated Distillation Process Configuration

In this example, a similar configuration to the auto-refrigerated distillation process that was
introduced in Figure 2 is used. The new process configuration is shown in Figure 4. The change
to the process in Figure 2 is the addition of a high pressure CO2 recycle compressor.
Approximately 10% of the recycle flow goes through the high pressure recycle compressor and
this further improves the efficiency of the process.

RECYCLE
COMPRESSORS

Feed
Stream

DISTILLATION
COLUMN
BAHX
NETWORK

PRODUCT
PUMP

Concentrated
Hydrogen Sulfide

INJECTION
PUMP

CO2 (<25 ppm H2S) CONDENSER

SEPARATOR
CO2 + H2S

H2S (>80% H2S) BOOSTER

PUMP

Figure 4 - Auto-Refrigerated Distillation Process Flow Diagram

In this two-stage auto-refrigerated distillation process, the reboiler duty required is split into side
and bottom reboilers and provided by the condensing carbon dioxide at two pressures in a heat
pump arrangement.
The auto-refrigerated process power consumptions are given in Table 3.

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 Table 3 - Auto Refrigerated Distillation Process Power Summary
Power Consumption, hp
Recycle Compressor 8,600
HP Recycle Compressor 590
Waste H2S Injection Pump 190
CO2 Booster Pump 610
CO2 Pipeline Pump 1360
Total Power 11,350

As can be seen from the comparison of the results presented in Table 2 and Table 3, the auto-
refrigerated process offers significantly lower power consumption than the externally
refrigerated process. These benefits are primarily due to the higher thermodynamic efficiency of
the process. In addition, it offers potential capital cost savings due to the substitution of the large
external refrigeration system with a much smaller recycle compressor.

IV. Process Sensitivities

In this section. additional considerations that affect the power consumption and capital
requirements of the new auto-refrigerated distillation process are summarized.

A. Feed Conditions; Temperature and Pressure

If obtained from another process upstream, the feed to the distillation process can be a liquid
CO2/H2S mixture at an elevated pressure as shown in the example case. Alternatively the feed
can be at a lower pressure and the CO2/H2S mixture can be gaseous as typically seen in the CO2
recycle streams of EOR operations. The feed to the distillation column is vapor phase at
approximately 300 psia, so for lower pressure feed sources, compression of the feed will be
required.

In the case of a liquid feed, some of the refrigeration required for the condensation of the recycle
CO2 is obtained by vaporization of the subcooled feed. If the feed is in the vapor phase, the
recycle compressor power requirements increase to replace the refrigeration that is otherwise
obtained from vaporization of the liquid feed stream. At the same temperature and pressure, a
vapor feed requires less than 10% additional power to substitute the refrigeration obtained from
the feed stream.

B. Effect of CO2 Product H2S specification

The impact of changing the H2S content of the CO2 product on the power consumed by the
process is shown in Figure 5. As the CO2 product H2S specification is tightened, the reflux
requirement for the carbon dioxide/hydrogen sulfide distillation increases for a fixed total
number of distillation column stages. For this example, the reduction of H2S content of the CO2
product leads to a nearly exponential increase in the power consumption. To mitigate this effect,

341
the total number of stages of the distillation column can be optimized based on power and capital
requirement considerations on a case by case basis.

14,000

13,000

12,000
Total Power Consumption, hp

11,000

10,000

9,000

8,000

7,000
1 10 100 1,000
H2S content of the CO2 Product, ppmv

Figure 5 - CO2 Product H2S Specification Sensitivity for a fixed total number of distillation column stages

C. Effect of H2S concentration in the Feed

The relative volatility of CO2 in a CO2/H2S mixture increases to reach about 2.6 at high H2S
concentrations at 600 psia as shown in Figure 1. This indicates that stripping of CO2 from H2S is
relatively easier at high H2S concentrations. As a result, the total power consumption of the
process tends to increase linearly with increasing H2S concentration in the feed gas, as shown in
Figure 6.

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 14,000

13,000

12,000
Total Power Consumption, hp

11,000

10,000

9,000

8,000

7,000
0% 5% 10% 15% 20%
H2S content of the Feed, mol %

Figure 6 - Feed H2S Content Impact for a fixed total number of distillation column stages

D. Effect of Light Components in the Feed

The process configuration in Figure 2 can be used for feed streams that have less than about 2
mol% lights (like methane and nitrogen) without any modification. As the lights concentration
increases above 2%, some of the lights are preferably rejected from the heat pump circuit in
order to prevent build-up of lights causing an increase in total power consumption as the
condensing pressure increases. As the subcooled recycled CO2 stream is let down in pressure to
the condenser separator, the light components dissolved in the CO2 flash off and can be purged
from the overhead of condenser separator. Purging light components also results in a small
amount of CO2 loss from the process. Depending on the value of the CO2, the amount of CO2
lost can be minimized or it can be recovered with alternative process configurations. The best
approach can be evaluated on a case by case basis.

E. Effect of NGLs in the Feed

Depending on the source of the feed to the plant, the feed can contain NGLs (C3+). Typically
CO2 that is recycled at EOR fields contains NGLs, or the gas source could contain NGLs. The
NGLs do not have to be removed upstream of the process and most of the NGLs condense with

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H2S. In this case, an additional small column can be used to separate the NGLs from the
condensed H2S where H2S can be stripped out of the NGLs to low ppm levels. With this type of
configuration approximately 90% recovery of C3+ and near 100% recovery of C4+ can be
achieved. The energy and capital requirements of recovering NGLs are comparatively small.
The C3 recovery is limited to approximately 90% as some C3 is lost with the H2S at the C3/H2S
azeotropic concentration.

F. Feed Carbonyl Sulfide (COS) Impact

Sour gas fields with high H2S concentrations can also have a few hundred parts per million of
Carbonyl Sulfide (COS). While there are no commonly accepted COS specifications for EOR
pipelines, there may be total sulfur specifications that need to be met.

The removal of COS from gas streams is challenging for conventional solvent-based AGR
systems, as COS removal efficiencies are typically limited because the COS affinity is close to
CO2 [7]. In cases where COS is a concern for the pipeline, additional COS removal processes
can be required.

In the distillation process, based on phase equilibrium considerations, nearly all of the COS is
expected to be removed from the CO2 along with the H2S. Thus the COS may be conveniently
disposed of using acid gas injection.

V. Conventional Solvent-Based AGR Process

Acid gas removal processes have been demonstrated commercially for high H2S content natural
gas fields at Exxon Mobil’s Shute Creek Facility [4] and the Lost Cabin Gas Processing Plant [5]
where the H2S concentrations are greater than 4% mole. Both of these plants use DEPG as a
solvent to remove H2S and CO2 from the gas.

A typical AGR process flow diagram that uses DEPG as solvent is shown in Figure 7. The
liquid CO2/H2S feed is vaporized against the condensing treated CO2 in the economizer heat
exchanger. The feed is then contacted with the cold lean solvent in the absorber column for
removal of H2S to meet the CO2 product requirements. The treated gas leaving the absorber
column is dehydrated and compressed to approximately 800 psia which enables condensing
against the incoming liquid feed that is at 45 °F. The condensed product CO2 is then pumped to
the final pipeline pressure of 2,500 psia.

While H2S is removed in the absorber, some CO2 is also co-absorbed with the lean cold DEPG
solvent. This co-absorbed CO2 is recovered from the rich solvent with High Pressure (HP),
Medium Pressure (MP) and Low Pressure (LP) flashes: 290 psia, 90 psia and 19 psia
respectively. These flash stages concentrate H2S in the semi-lean solvent that is sent to the
stripping column.

The flashed CO2 and H2S are compressed in a three stage recycle compressor to the absorber
feed pressure and recycled to the absorber column. The semi-lean solvent after LP flash is
warmed up in the lean-rich heat exchanger against the hot lean solvent leaving the stripper
column bottom. The warmed-up semi-lean solvent then goes to the stripping column where the
solvent is regenerated using heat supplied from an external source such as steam or hot oil. Lean

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solvent leaving the stripper column is pumped and cooled, first in the lean-rich heat exchanger,
then in the trim air cooler and in the lean solvent cooler that uses external refrigeration. The
overhead of the stripper column has greater than 85 mol% H2S and the remainder is CO2. The
low pressure gaseous concentrated H2S is then compressed in a four stage acid gas compressor
that uses air cooled intercoolers and an aftercooler. The acid gas is compressed above 700 psia to
enable condensing against air in the aftercooler. The condensed acid gas is then pumped to final
injection pressure of 3,500 psia with the injection pump.

The liquid knock-outs after the intercoolers (which are not shown Figure 7) are recycled with
make-up water to the stripping column.

DEHYDRATION

Feed REFRIGERANT
Stream LEAN SOLVENT
COOLER

STRIPPER

RECYCLE
COMPRESSORS

PRODUCT
PUMP
ABSORBER
Concentrated
Hydrogen Sulfide

INJECTION SOLVENT
PUMP PUMP

CO2 (<25 ppm H2S)

H2S COMPRESSION
CO2 + H2S

H2S (>80% H2S)

Solvent

Figure 7 - Solvent Based AGR Process Flow Diagram

Alternative configurations of the solvent based AGR system to that shown in Figure 7 can be
used to optimize the capital and operating cost trade-offs. The process in Figure 7 is modeled
using the Aspen Plus® simulation software package with customized DEPG properties for the
example case feed and product composition summarized in Table 1.

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 Table 4 - Solvent-Based AGR and Auto Refrigerated Process Comparison

Economic Parameters Value

Power $/kWh 0.06
Fuel $/MMBtu 3.5
On Stream day/yr 356
Process Configuration DEPG Auto-Refrigerated
Distillation
Total Power hp 14,800 11,350
Fuel MMBtu/hr 48 -
Power Cost $/yr 5,650,000 4,340,000
Fuel Cost $/yr 1,435,000 -
Total Utility $/yr 7,085,000 4,340,000

In Table 4, the DEPG solvent-based AGR system utility costs are compared against the auto-
refrigerated distillation process. The auto-refrigerated distillation process shows yearly utility
cost savings of nearly 40% over the DEPG solvent based AGR system. In addition, the DEPG
AGR system may require dehydration of the treated gas that will also require additional fuel.
The DEPG solvent removes approximately 20% of the COS in the feed thus leaving
approximately 220ppmv of COS in the treated gas. The COS in the CO2 product might not be
desirable for the EOR operators and so might require additional COS removal either upstream or
downstream of the DEPG process.

VI. Summary and Conclusions

The novel technical features of the auto-refrigerated distillation process configuration have been
described and a significant energy intensity advantage over using external refrigeration is
demonstrated for the separation of CO2 and H2S using conventional distillation. The energy
intensity is reduced primarily due to the efficiency of the heat pump arrangement; a significant
reduction in the heat exchanger duty by using the same fluid in a reboiler/condenser arrangement
rather than a separate reboiler and condenser, in addition the improved distillation efficiency due
to having the side reboiler. In addition, it offers potential capital cost savings due to the
substitution of the large external refrigeration system with a much smaller recycle compressor.
Since the refrigeration is provided by recycled CO2, which is non-flammable, this type of process
offers additional safety benefits for the facility. Additional refrigerant storage and routine make-
up may be reduced or eliminated. The auto-refrigerated process configuration is also compared
against the conventional DEPG solvent based AGR process, and yearly utility cost savings
nearly 40% are shown in the example case.

For the example case, the power cost of the CO2 purification process is $0.13 per MSCF of CO2
assuming a unit cost for power of $0.06/kWh. However, this includes product compression and,
if the next best alternative is acid gas injection of the feed stream, the power required to increase
the pressure of the feed to 3,500 psia would be approximately 2,600 hp. Therefore the additional
power cost of removing H2S from the CO2 is $0.10 per MSCF compared to pumping of the feed
mixture for acid gas injection.

Detailed operating and capital cost evaluation of this process configuration leads to a total
processing cost in the range of $0.70-1.20 per MSCF of CO2 in 2014 for a 100 MMSCFD plant

346
located in the continental US. While the estimate of CO2 processing cost is provided in this
paper for the benefit of the reader, it should be used with caution as the actual cost of any plant
will depend greatly on the size, location and other critical project considerations and will require
detailed engineering and cost estimation to determine accurately.

Given that historically the price of CO2 for EOR is $0.50 per MSCF plus 1.5% of the oil price,
approximately $2/MSCF of CO2 for $100 per barrel of oil [17], the cost of CO2 cleanup using
this technology in the range of $0.70 to $1.20 could bring additional CO2 to the market for
beneficial EOR use.

Acknowledgements
The authors would like to acknowledge Dr. Steve Auvil, Mr. Vipul Parekh, Mr. Ken Kovak, and
Mr. Steve Carney from Air Products and Chemicals, Inc. for their contributions in the technical
and commercial development of this new technology.

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