The Removal of CO2 and N2 From Natural Gas Review (Rufford Et Al 2012)

8/18/2019 The Removal of CO2 and N2 From Natural Gas Review (Rufford Et Al 2012)

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The removal of CO2 and N2 from natural gas: A review of conventional andemerging process technologies

T.E. Rufford a, S. Smart b, G.C.Y. Watson a, B.F. Graham a, J. Boxall a, J.C. Diniz da Costa b, E.F. May a,n

a The University of Western Australia, Centre for Energy, 35 Stirling Highway, Crawley, WA 6009, Australiab The University of Queensland, FIMLab—Films and Inorganic Membrane Laboratory, School of Chemical Engineering, College Road, St Lucia, Qld. 4072, Australia

a r t i c l e i n f o

Article history:

Received 5 September 2011Accepted 2 June 2012Available online 19 June 2012

Keywords:

CO2 capture

nitrogen rejection

contaminated gas

membrane

absorption

pressure swing adsorption

hydrates for gas separation

a b s t r a c t

This article provides an overview of conventional and developing gas processing technologies for CO2

and N2 removal from natural gas. We consider process technologies based on absorption, distillation,adsorption, membrane separation and hydrates. For each technology, we describe the fundamental

separation mechanisms involved and the commonly applied process flow schemes designed to produce

pipeline quality gas (typically 2% CO2, o3% N2) and gas to feed a cryogenic gas plant (typically 50 ppmv

CO2, 1% N2). Amine absorption technologies for CO2 and H2S removal (acid gas treating) are well-

established in the natural gas industry. The advantages and disadvantages of the conventional amine-

and physical-solvent-based processes for acid gas treating are discussed. The use of CO2 selective

membrane technologies for bulk separation of CO2 is increasing in the natural gas industry. Novel low-

temperature CO2 removal technologies such as ExxonMobil’s Controlled Freeze ZoneTM process and

rapid cycle pressure swing adsorption processes are also emerging as alternatives to amine scrubbers in

certain applications such as for processing high CO2 concentration gases and for developing remote gas

fields. Cryogenic distillation remains the leading N2 rejection technology for large scale (feed rates

greater than 15 MMscfd) natural gas and liquefied natural gas plants. However, technologies based on

CH4 selective absorption and adsorption, as well as N2 selective pressure swing adsorption technol-

ogies, are commercially available for smaller scale gas processing facilities. The review discusses the

scope for the development of better performing CO2 selective membranes, N2 selective solvents and N2selective adsorbents to both improve separation power and the durability of the materials used in novel

gas processing technologies.

& 2012 Elsevier B.V. All rights reserved.

1. Introduction

In 2010 natural gas (NG) supplied 23.81% of the world’s energy

demand and the volume of natural gas consumed increased by 7.4%

over 2009 levels driven by economic recovery the USA and the

continuing economic development of China, India and Russia (BP

Statistical Review of World Energy June 2011, 2011). The growth in

the global demand for natural gas has led to a re-evaluation of thedevelopment potential of unconventional, stranded and contami-

nated gas reserves that were previously considered economically

unviable. Furthermore, many of the significant natural gas reserves

are located far from the large established gas markets in Western

Europe, Japan and South Korea. Therefore, significant volumes of

natural gas must be transported long distances from exporting

countries either by pipeline or by tanker as liquefied natural gas

(LNG); the economics of this choice have been discussed by many

authors such as Rojey et al. (1997). Table 1 contains a list of the

major natural gas importing and exporting countries in 2010

including a breakdown of the amounts transported by pipeline

and LNG. The production of LNG is clearly already essential to the

international trade of natural gas and its importance is set to

increase further over the next two decades, particularly in the

Asia-Pacific region.

The estimated world volumes of sub-quality natural gasreserves, including sour natural gas reserves, are significant.

Sub-quality natural gas reserves are defined as gas fields contain-

ing more than 2% CO2, 4 % N2 and 4 parts per million (ppm)

hydrogen sulphide (H2S) (Kidnay and Parrish, 2006). Burgers et al.

(2011), for example, estimate that 50% of the volume of known

gas resources contain more than 2% CO2. The development of

sub-quality, unconventional and remote natural gas reserves,

including development via LNG production, can present new

challenges to gas processing that require more efficient

approaches to the conventional absorption and cryogenic con-

densation technologies that are most commonly used for the

removal of carbon dioxide and nitrogen, respectively, from

Contents lists available at SciVerse ScienceDirect

journal homepage: w ww.elsevier.com/locate/petrol

Journal of Petroleum Science and Engineering

0920-4105/$- see front matter & 2012 Elsevier B.V. All rights reserved.

http://dx.doi.org/10.1016/j.petrol.2012.06.016

n Corresponding author.

E-mail address: [email protected] (E.F. May).

Journal of Petroleum Science and Engineering 94 –95 (2012) 123–154

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natural gas. Additional gas processing challenges arise from new

environmental regulations that may call for capture and seques-

tration of CO2 from gas fields and stricter regulation of CH4emissions in N2 vent streams from natural gas production

facilities.

Carbon dioxide, as well as H2S and other acid gases, must be

removed from natural gas because in the presence of water theseimpurities can form acids that corrode pipelines and other

equipment. Although this paper is not focussed on the removal

of H2S, it should be noted that health and safety is a key driver for

removal of this highly toxic gas from sour natural gas streams.

Several of the CO2 capture technologies described are more

selective for H2S than CO2. Furthermore, CO2 provides no heating

value and must be removed to meet gas quality specifications

before distribution to gas users. The maximum level of CO2permitted in natural gas transmitted to customers by pipeline is

typically less than 3% (Hubbard, 2010) although local contracts

may stipulate quality specifications different to these values. The

specifications of CO2 removal from natural gas to be processed in

a cryogenic plant to produce LNG, recover liquefied petroleum gas

(LPG) or natural gas liquids (NGL) are more stringent than thosefor typical gas pipelines. For example, in addition to more

extensive dehydration, the CO2 concentration of the natural gas

should be less than 50 ppmv (Hubbard, 2010) before it enters the

cryogenic processes within the plant to avoid the formation of

dry ice.

A typical gas pipeline specification for N2 is 3% (Kidnay and

Parrish, 2006); because N2 is inert the main driver for its removal

from pipeline sales gas is to increase the heating value of the gas.

Nitrogen will not freeze or lead to corrosion in a cryogenic gas

plant, but a maximum concentration of 1% N2 is often specified for

LNG, for example to avoid stratification and rollover of the liquid

product during shipping. To ensure this specification is met in

plants liquefying NG with a high N2 content, cryogenic distillation

columns known as nitrogen rejection units (NRU) are integrated

into the process; these columns are both expensive and energy

intensive. In addition, high N2 concentrations in natural gas

processed by an LNG plant are energetically parasitic because a

significant amount of energy is wasted cooling the N2 from

ambient temperatures to those of LNG (about 161 1C at atmo-

spheric pressure). Furthermore, even if the NG does not have a

high N2 concentration, an NRU may be necessary because the

liquefaction process produces two product streams: the LNG and

an ‘end-flash’ gas which is a mixture of N2 and CH4 vapour. This‘end-flash’ gas can be used as fuel gas for the plant and/or blended

into sales gas destined from a conveniently located pipeline.

However, if production of end-flash gas exceeds the fuel gas

requirements and if no pipeline is available for its disposal an

NRU may be necessary to produce a stream pure enough that it

can be vented to atmosphere. Given the increasing regulation of,

and costs associated with mitigation of, greenhouse gas emissions

this relatively expensive NRU option could become increasingly

common.

This review provides an overview of conventional, developing

and novel gas processing technologies for CO2 and N2 removal

from natural gas, with particular attention paid to large scale LNG

production. The introduction sections provide an overview of a

typical process flow sheet and the fundamental unit operations

involved in natural gas processing. The subsequent sections of

this review describe the core concepts and industrial application

of absorption, distillation, adsorption, membrane and hydrate gas

separation technologies.

Many of the process technologies we describe for CO2 capture

from CH4 can also be applied to CO2 capture from combustion flue

gases, for example in coal-fired power stations. The recent

advances and future trends in technologies for capturing CO2 at

the power plant are reviewed elsewhere by many other authors,

including Figueroa et al. (2008), MacDowell et al. (2010) and

Ebner and Ritter (2009). However, the process conditions avail-

able in the natural gas processing facility can be very different to

those conditions available for post-combustion flue gas (predo-

minantly a mixture of CO2, N2 and H2O) treatment. The two most

significant differences between these applications are the CO2partial pressure, and the level of CO2 removal required. In the first

case, the natural gas feed to the CO2 removal unit is typically

available at high pressures (more than 3000 kPa) while flue gases

are typically at close to atmospheric pressure, so the driving force

for CO2 capture from flue gas (partial pressure CO2 typically less

than 15 kPa) is much lower than from natural gas. In the second

case the level of CO2 removal required for natural gas production

is greater, especially for LNG production plants, than bulk separa-

tion of CO2 from flue gas. Thus, some promising strategies for

CO2 capture in thermal power plants such as oxy-combustion

(a modified combustion process that can produce a high CO2concentration flue gas, Plasynski et al., 2009) are not relevant to

CO2 removal from natural gas.

1.1. Overview of conventional gas processing flow schemes

The major operations that can be used in natural gas proces-

sing and LNG production are shown in Fig. 1. Common process

operations include inlet gas compression, acid gas removal,

dehydration, LPG/NGL recovery and hydrocarbon dewpoint con-

trol, nitrogen rejection, outlet compression and liquefaction.

Depending on the available markets, feed gas properties, product

specifications and the gas flow rate, the units identified in Fig. 1

may not all be required. For example a nitrogen rejection unit

(NRU) may not be required for the production of pipeline gas from

a feed gas containing only low N2 concentrations, or if a high N2content gas can be blended with richer natural gas streams to

meet pipeline gas specifications.

Table 1

Volumes of natural gas traded as pipeline and LNG (billions of cubic metres) by top

natural gas exporting and importing countries in 2010 (BP Statistical Review of

World Energy June 2011, 2011).

Top natural gas exporters

Pipeline LNG Total

1 Russian Federation 186.5 13.4 199.9

2 Norway 95.9 4.7 100.63 Qatar 19.2 75.8 94.9

4 Canada 92.4 0.0 92.4

5 Algeria 36.5 19.3 55.8

6 Netherlands 53.3 0.0 53.3

7 Indonesia 9.9 31.4 41.3

8 Malaysia 1.5 30.5 32.0

9 U.S. 30.3 1.6 32.0

10 Australia 0.00 25.4 25.4

Top natural gas importers

Pipeline LNG Total

1 United States 93.3 12.2 105.5

2 Japan – 93.5 93.5

3 Germany 92.8 – 92.8

4 Italy 66.3 9.1 75.35 United Kingdom 35.0 18.7 53.6

6 France 35.0 13.9 48.9

7 South Korea – 44.4 44.4

8 Turkey 28.8 7.9 36.7

9 Spain 8.9 27.5 36.4

10 Ukraine 33.0 – 33.0

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In conventional gas processing H2S and CO2 are removed in an

acid gas removal unit (AGRU) using aqueous amine absorptionprocesses. The sweetened gas leaving the amine process is saturated

with water, so typically the AGRU is located upstream of the

dehydration facilities. The final sections of the gas liquefaction plant

(main cryogenic heat exchange (MCHE) in Fig. 1) can operate at

temperatures as low as 161 1C, and therefore, it is essential that

components that could freeze, and cause blockages in the cryogenic

equipment, at low temperatures are removed from feed to the

cryogenic plant. In addition to CO2, components that could freeze in

the cryogenic plant include water (typically removed to less than

0.1 ppmv), heavier paraffinic hydrocarbons and aromatics such as

benzene. To protect the aluminium plate-fin heat exchangers used

in the NRU the feed gas to the cryogenic plant must be free of

mercury (below 0.01 mg/Nm3) (Kidnay and Parrish, 2006). The gas

feed entering the (Pre-Cool section of the) liquefaction plant istypically delivered, or compressed to, a pressure of more than

3400 kPa and at a temperature close or slightly above ambient.

The conventional method of removing N2 from natural gas is by

cryogenic distillation; hence NRUs are usually closely integrated

within the liquefaction process. There are several variations of

cryogenic liquefaction processes including, for example, Air Products

and Chemicals (APCI) C3MR process and the ConocoPhillips Opti-

mised Cascade process. Commonly a propane refrigeration loop is

used to pre-cool the gas before it enters a main cryogenic heat

exchanger (MCHE) which might use a mixed refrigerant or a cascade

of several pure fluid refrigeration cycles. If an NRU is incorporated

with the liquefaction process then it would commonly be located

downstream of the MCHE and before the final depressurisation

stage of the LNG production process.

1.2. Overview of gas separation mechanisms

Separation processes can be designed to exploit differences in

the molecular properties or the thermodynamic and transport

properties of the components in the mixture. Molecular proper-

ties that could be exploited to achieve a separation of CO2, N2 and

CH4 include the differences in kinetic diameter, polarizability,

quadrupole and dipole moments of the molecules (Table 2).

Thermodynamic and (interphase) transport properties that could

be exploited include vapour pressure and boiling points, solubi-

lity, adsorption capacity and diffusivity. Based on these properties

of the components to be separated, the primary operations for the

separation and purification of gases apply one of the following

inherent separation mechanisms: (1) phase creation by heat trans-

fer and/or shaft work to or from the mixture, (2) absorption in a

liquid or solid sorbent, (3) adsorption on a solid, (4) permeationthrough a membrane and (5) chemical conversion to another

compound (Kohl and Nielsen, 1997; Seader and Henley, 2006).

The first four of these operations are discussed in this review;

direct chemical conversion of CO2 from natural gas to a useful

product is beyond the current scope. An example of such direct

conversion, which is attracting significant current research inter-

est, is the dry reforming process where CO2 reacts with CH4 to

produce syngas (a mixture of H2 and CO2) which can then be used

for the production of liquid fuels via Fischer–Tropsch reactions.

Recent reviews on the conversion of CO2 include articles by

Havran et al. (2011), Song (2006) and Zangeneh et al. (2011).

Separations involving the creation of a phase include the partial

condensation or partial vaporisation of species with very different

volatility from a feed mixture; and this category can also include

Fig. 1. Conventional gas processing operations in a typical natural gas plant with a cryogenic process for liquefied natural gas production. The required operations and

process flow arrangements vary depending on the local feed gas properties, size of the plant, and product specifications; and may not all be required (MCHE¼main

cryogenic heat exchanger).

Table 2

Physical properties of methane, carbon dioxide and nitrogen.

Property CH4 CO2 N2

Kinetic diameter ( Å) (Tagliabue et al., 2009) 3.80 3.30 3.64

Normal boiling point (NBP) (K) (Lemmon et al., 2010) 111.7 – 77.3

Critical temperature (K) (Lemmon et al., 2010) 3.80 304.1 126.2

Critical pressure (kPa) (Lemmon et al., 2010) 4600 7380 3400

DH vap at NBP (kJ/mol) (Linstrom and Mallard, 2011) 8 .17 26.1 5.58

Polarisability ( Å3) (Tagliabue et al., 2009) 2.448 2.507 1.710

Quadrapole moment (D Å) (Tagliabue et al., 2009) 0.02 4.3 1.54

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desublimation of CO2. If the volatility differences among species

to be separated are not sufficiently large, as is the case for

N2–CH4, to achieve the desired separation in a single partial vapor-

isation or partial condensation contact stage, then a distillation

process involving multiple vapour-liquid contact stages may be

required.

One of the most fundamental properties of an inherent separa-

tion mechanism is its selectivity with respect to the components i

and j being separated. As will be discussed, some separationmechanisms have more than one type of selectivity; however, the

most common and applicable type is the equilibrium selectivity, aij,which is defined in terms of phase compositions:

aij xi x j

y j yi

ð1Þ

Here xi and x j are the mole fractions (or convenient concentra-

tion units) of components in one equilibrium phase, and y i and y jare the mole fraction of components in a second equilibrium

phase. This definition of equilibrium selectivity derives from its

central role in distillation theory, where it is also known as the

relative volatility (McCabe et al., 2005). However, the equilibrium

selectivity can be applied more broadly than to vapour–liquid

equilibrium in distillation; it is useful in the analysis of adsorp-tion- and absorption-based processes. The equilibrium selectivity

can also be applied, through the use of appropriate analogies and

use of concentrations of the permeate ( x) and feed ( y) streams, to

the analysis of membrane processes which are not strictly

equilibrium-based. It is important to note that selectivity is often

linked to the characteristic energy of the separation mechanism,

and therefore the regeneration energy needed if the mechanism is

to be used in a cyclical process.

The performance of a separation process is governed by two

factors: the inherent selectivity of the separation mechanism being

utilised and the degree to which that mechanism can be exploited

through appropriate engineering design. This gives rise to another

metric, known as the separation power, SP ij (Seader and Henley,

2006), which can be used to quantify the performance of an entireseparation process in terms of the product stream compositions. For

a single stage equilibrium operation aij ¼ SP ij; however, for a processutilizing multiple separation stages SP ij values much larger than the

single stage equilibrium selectivity can be achieved.

SP ij C ð1Þ

i

C ð1Þ j

0@

1A C ð2Þ j

C ð2Þi

! ð2Þ

Here C ð1Þi and C ð2Þi are the concentrations of component i in each of

the product streams, and C ð1Þ j , C ð2Þ

j the concentrations of component j

in each of the product streams. The symbol C for concentration has

been used here to distinguish Eq. (2) as the separation power of the

process from the equilibrium selectivity of the mechanism, although

in many cases the composition of the product streams used in the

calculation of SP ij may be in units of mole fraction. In principle any

selective mechanism with aij41 could be engineered to achieve a

given separation power if no constraints exist on capital, operationaland energy costs. In practice a separation process will be selected if it

can achieve the separation power necessary to meet the desired

product specifications at a cost that is (i) lower than other processing

alternatives, and (ii) economically viable in terms of the

product’s value.

In Table 3, indicative values of aij and SP ij are given for the mostcommon industrial processes used to separate CO2 or N2 from natural

gas for the purpose of evaluating the prospects of new and emerging

technologies. To generalise this measure for any of the four primary

separation operations the product streams must be defined. For CO2removal from natural gas, a convenient approach may be to compare

the aCO2,CH4 values of the various separation technologies by definingthe equilibrium phases to be used with Eq. (1) as the CO2 selective

phase: that is the amine solution or physical solvent for absorption,

the high-purity liquid CO2 product for distillation, the adsorbed phase

( xi or qi) for gas–solid adsorption systems, and the permeate stream

from a CO2 selective membrane stage. Using available data for CO2removal processes, such as the amine absorption data included in

Kohl and Nielsen (1997), the typical separation power of processes to

produce pipeline quality gas (2% CO2) from a feed mixture containing

5% CO2 can be estimated with Eq. (2) as shown in Table 3. The

inherent equilibrium selectivity can be exploited for each separation

mechanism to achieve SP ijbaij through the arrangement of multipleseparation stages into an engineering process system, for example:

trayed absorption and distillation columns, multiple adsorption beds

operating in adsorption–desorption cycles, and membrane stages

with interstage recompression and recycle loops. This simplified

analysis shows clearly that the typical inherent process separation

powers for chemical absorption technologies are much larger thanthe best alternatives currently available to treat large gas flows; hence

amine absorption is the most commonly applied type of AGRU

despite the energy required for regeneration and corrosive nature of

the amine solutions.

For N2 rejection technologies the conventional cryogenic distilla-

tion technology has a separation power more than 8 times that of N2selective adsorption and membrane processes. However, the

Table 3

Inherent equilibrium selectivity, ai,j, for the separation of CO2 from CH4 and N2 from CH4 by different operations, with the typical process separation power ( SP i,j) achievedin example technologies implementing these separation operations. The separation powers are calculated for typical processes to (a) remove CO 2 from a feed gas

containing 5% CO2 to produce pipeline quality gas with 2% CO2, and (b) to reject N2 from a gas containing 4% N2 to produce a stream with 1% N2 for LNG production.

Process Separating agent Typical inherent

equilibrium selectivity

Typical process

separation power

(a) CO2/CH4 separation aCO2 ,CH4 SPCO2 ,CH4Amine absorption (MDEA) Liquid absorbent 860 3300

Physical solvent (chilled methanol) Liquid absorbent 318 1900

Adsorption—CO2 selective Solid adsorbent 2–8.5 6–22

Membrane—CO2 selective Membrane 15–20 20–40

(b) N2/CH4 separation aN2 ,CH4 SPN2 ,CH4Cryogenic distillation Heat transfer 5–8 320

Adsorption—N2 selective Solid adsorbent 1.3–2a 8–40a

Adsorption—CH4 selective Solid adsorbent 0.25–0.5 1.05–2

Membrane—N2 selective Membrane 2–3 2–10

Membrane—CH4 selective Membrane 0.25–0.3 2–10

a Inherent kinetic selectivities of narrow pore adsorbents are reported from 2 to 10, which could allow a N2 selective process with much higher separation power to be

engineered.



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separation of N2 from CH4 in adsorption-based processes is

enhanced by the differences in rates of diffusion of N2 and CH4 that

have been reported for small pore titanosilicate ETS-4 materials

(Marathe et al., 2004a). Guild Associates’ Molecular GateTM pressure

swing adsorption (PSA) process (Guild Associates, 2007) i s a

commercial example of a N2 rejection process that relies on the

kinetic selectivity of N2 over CH4.

In practice, the achievable separation power may be much

lower than the ideal separation powers estimated in Table 3.Furthermore, there are many other factors beside a separation

power or inherent selectivity that a process engineer designing or

selecting a gas separation process will need to consider. Other

factors that influence the selection of process technology for CO2/

CH4 and N2/CH4 separations include the level of contaminants in

the feed gas, the required level of contaminant removal or

product purity, the flow rate and condition of the feed gas

(temperature, pressure, water content), and for AGRUs the need

for simultaneous or selective removal of H2S. Process selection is

also influenced by the available disposal routes for the removed

contaminants, which may include reinjection of CO2 for enhanced

oil recovery (EOR) or enhanced gas recovery (EGR), and venting of

N2 to atmosphere. The process plant layout and available plant

space for the separation process must be considered. If the

separation process is to be installed on an offshore platform, a

floating LNG (FLNG) plant or as a retrofit in an existing production

plant, then the process footprint – the plan area and/or height

occupied by the process equipment – may influence the choice of

process technology. Each of these process selection criteria may

have material impacts on the feasibility, energy requirements and

costs of CO2 and N2 removal processes.

There are many factors that influence the cost of a separation

process including the extent to which it has been used

successfully in the past. Consequently, a process engineer design-

ing or selecting a gas separation process is likely to be more

interested in the ratio of its separation power to its cost rather

than in the inherent selectivity upon which the process is based.

However, once a separation process is sufficiently mature, its

separation power to cost ratio will generally only improveasymptotically, unless a significant improvement in its inherent

selectivity can be achieved. Thus, the starting point for scientists

and engineers aiming to develop a new separation technology

should be an analysis of the inherent selectivities of both the

new technology and the conventional one with which it will

compete.

2. Absorption

This section focuses primarily on CO2 absorption processes,

but also introduces technologies for N2 rejection by the selective

absorption of CH4 in hydrocarbon solvents (Mehra and Gaskin,1997) and the potential for the development of N2 selective

solvents. Although this review is concerned primarily with CO2removal, the selection of the AGRU process is more often

determined by the H2S removal requirements. Thus, most of the

technical literature concerning acid gas treating focuses on the

absorption of H2S. Commonly, the capacity of a sorbent is

reported as an acid gas loading capacity which includes the

capacity for CO2 and H2S. We have noted in Tables 5 and 10

which of the process technologies are suitable for the simulta-

neous or the selective absorption of H2S.

A large number of commercial processes are available for CO2absorption in chemical and physical solvents, including the

technologies listed in Table 4. Chemical absorption processes

with amine solutions are the most commonly used acid gasremoval technologies in the natural gas industry (GPSA

Engineering Data Book, 2004). The chemical absorption processes

rely on reactions of the CO2 with the sorbent to form weakly

bonded intermediate compounds, and these reactions can be

reversed by the application of heat to release the CO2 and

regenerate the sorbent (Olajire, 2010). Physical solvents, such as

the mixture of polyethylene glycol–dimethyl ethers used in the

Selexols process, selectively absorb CO2 from the natural gas feed

according to Henry’s law so that absorption capacity increases at

high pressure and low temperature.

The two major cost factors in gas–liquid absorption processes

are (1) the required sorbent circulation rate, which is determined

by the amount of CO2 that must be removed from the feed gas

and the CO2 loading capacity of the sorbent, and (2) the energy

required to regenerate the sorbent (Kidnay and Parrish, 2006).

The acid gas loading capacity of physical solvents at low to

moderate CO2 partial pressures is generally lower than that of

chemical absorbents. However, physical solvent processes have

lower energy requirements for regeneration because the heat of

Table 4

Examples of commercial absorption technologies for CO2 capture and gas sweetening.

Vendor/licensor Sorbent Reference

Chemical absorption

EconamineSM Fluor Digylcolamine, mono-ethanolamine http://www.fluor.com

ADIP-X Shell MDEAþaccelerator www.shell.com

aMDEAs BASF MDEA www.basf.de

GAS/SPEC Ineos MDEA www.gasspec.com

UCARSOL DOW MDEA http://www.oilandgas.dow.com

KM CDR Mitsubishi Heavy Industries KS-1 hindered amine Mimura et al., (1995)

Benfield UOP Potassium carbonate UOP Overview of Gas Processing

Technologies and Applications, (2010)

Catacarb Eickmeter and Associates Potassium carbonate (þorgani c addi ti ve) www.catacarb.com

Flexsorb HP Exxon Mobil Potassium carbonate (þsteric amine) www.exxonmobil.com

Physical absorption

Fluor SolventSM Fluor Dry propylene carbonate www.fluor.com

Selexol UOP/DOW Mixed polyethylene glycol dimethyl ethers www.uop.com

Purisol Lurgi n-Methyl-2-pyrrrolidone www.lurgi.com

Rectisol Lurgi Chilled methanol www.lurgi.com

Ifpexol IFP Chilled methanol Larue and Lebas, (1996)

Mixed-solvent processes

Sulfinol-D Shell SulfolaneþDIPAþwater www.shell.com (Rajani, 2004)

Amisol Lurgi Methanolþ secondary amineþwater Kohl and Nielsen, (1997)


http://www.fluor.com/http://www.shell.com/http://www.basf.de/http://www.gasspec.com/http://www.oilandgas.dow.com/http://www.exxonmobil.com/http://www.fluor.com/http://www.uop.com/http://www.lurgi.com/http://www.lurgi.com/http://www.shell.com/http://www.shell.com/http://www.lurgi.com/http://www.lurgi.com/http://www.uop.com/http://www.fluor.com/http://www.exxonmobil.com/http://www.oilandgas.dow.com/http://www.gasspec.com/http://www.basf.de/http://www.shell.com/http://www.fluor.com/


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absorption for physical solvents is much lower than the heat of

absorption for chemical solvents.

2.1. Chemical absorption processes for acid gas treating

2.1.1. Aqueous amine processes

Amines are organic compounds derived from ammonia (NH3)

where one or more hydrogen atoms have been substituted with

an alkyl or aromatic group. It is the (–NH2) functional group of the

amine molecule that provides a weak base that can react with the

acid gases. The absorption of CO2 occurs via a two-step mechan-

ism: (1) the dissolution of the gas in the aqueous solution,

followed by (2) the reaction of the weak acid solution with the

weakly basic amine. The first physical absorption step is governed

by the partial pressure of the CO2 in the gas feed. The reactions

involved in the second step of CO2 absorption in aqueous amines

have been widely studied, with a large number of reference

materials on the reaction mechanisms (Bindwal et al., 2011;

Kohl and Nielsen, 1997; Penny and Ritter, 1983; Vaidya and

Kenig, 2007; Versteeg et al., 1996) and guidelines for process

operation (GPSA Engineering Data Book, 2004) available in the

literature. The fundamental reactions involved in CO2 absorption

in amine treating are (Kohl and Nielsen, 1997):

Dissociation of water:

H2O"HþþOH (3)

Hydrolysis and dissociation of dissolved CO2:CO2þH2O"HCO3

þHþ (4)

Protonation of the amine:

RNH2þHþ"RNH3

þ (5)

Carbamate formation:

RNH2þCO2"RNHCOOþHþ (6)

The dissociation reactions are shown here to highlight that the

pH of the amine solution is an important process parameter

because the concentrations of the ionic species Hþ , OH- and HCO3-

in the amine solution affect the other reactions involving

the amine.

Amines can be classified according to the number of hydrogen

atoms that have been substituted, as primary (R–NH2, where R is

a hydrocarbon chain), secondary (R–NH–R 0) or ternary (R 0–NR–

R 00). For primary and secondary amines, such as monoethanola-

mine (MEA) and diethanolamine (DEA), the carbamate formation

reaction (Eq. (6)) predominates; this reaction is much faster than

the CO2 hydrolysis reaction (Eq. (4)). The stoichiometry of the

carbamate reaction suggests that the capacity of primary and

secondary amines is limited to approximately 0.5 mol of CO2 per

mole of amine. However, DEA-based amine processes can achieve

loadings of more than 0.5 mol of CO2 per mole of amine through

the partial hydrolysis of carbamate (RNHCOO-) to bicarbonate

(HCO3), which regenerates some free amine (Kidnay and Parrish,

2006).Tertiary amines such as MDEA, which do not have a free

hydrogen atom around the central nitrogen, do not react directly

with CO2 to form carbamate. Instead, CO2 reactions with tertiary

amines proceed via equivalent reactions to those shown in Eqs.

(4) and (5), which are much slower than the reaction in Eq. (6), to

give the overall reaction:

RR 0R 00NþCO2þH2O"RR 0R 00NHþþHCO3

(7)

The stoichiometry in Eq. (7) shows that theoretically tertiary

amines can achieve a loading of 1 mol of CO2 per mole of amine,

which is double the CO2 loading capacity of primary amines. Also

the required heat of regeneration is lower for tertiary amines. A

disadvantage of tertiary amines is that the absorption kinetics areslower than for primary and secondary amines. For some natural

gas treating processes the slow kinetics of CO2 absorption in

tertiary amines can be utilised to achieve selective H2S removal

by optimisation of the contact time in the absorber to minimise

CO2 uptake (GPSA Engineering Data Book, 2004). Alternatively, to

enhance CO2 the absorption kinetics of tertiary amines an

activator (usually a primary or secondary amine) may be added

to increase the rate of hydrolysis of carbamate and dissolved CO2(GPSA Engineering Data Book, 2004).

Since the first application of tertiary amines in the mid-1970s,

significant research has been directed into the further develop-

ment of novel amine solvents. To accelerate the reaction rate of

tertiary amines with CO2, a primary or secondary amine can

be included as an activator (GPSA Engineering Data Book, 2004).

Fig. 2. Process flow diagram of a typical amine-solvent (MDEA)-based chemical adsorption system for the separation of CO 2 and other acid gases from natural gas

(Hubbard, 2010; Kohl and Nielsen, 1997).



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For example, the cyclic diamine piperazine has been studied as a

promoter to improve the CO2 mass transfer rates of MDEA and

MEA (Bishnoi and Rochelle, 2002); a commercial example of this

technology is the aMDEA solvent from BASF. The piperazine

reacts rapidly with CO2 in the vapour phase, which accelerates

the dissolution of CO2 into carbonic acid, which can then react

quickly with MDEA. Sterically hindered amines are either primary

or secondary amines with large bulky alkyl or alkanol groups

attached to the nitrogen (Seagraves and Weiland, 2011), whichreduces the carbamate stability. The molecular configuration

dictates the amount of CO2 removal: severely hindered amines,

such as ExxonMobil’s Flexsorb SE, have very low rates of CO2absorption and allows selective H2S removal. In contrast, moder-

ately hindered amines, such as 2-amino, 2-methyl, 1-propanol

(AMP), are characterised by high rates of CO2 absorption and high

capacities for CO2. Weiland et al. (2010) evaluated the use of AMP

with MEA for CO2 capture from flue gases and found it had several

advantages including at least a 15% reduction in the required

regeneration energy.

A typical process flow diagram for the removal of acid gas from

a sour gas feed using methyldiethanolamine (MDEA) is shown in

Fig. 2. Properties and typical operating conditions for commonly

used aqueous amine solutions are shown in Table 5. The basic

process flow for other amine absorption systems is similar to that

shown for MDEA, although some commercial process designs

often feature multiple column feeds and contactor sections. Any

liquids or solids in the sour feed gas are removed in an inlet

separator before the gas enters at the bottom of amine contactor.

Typical operating pressures for amine contactors are in the range

of 50–70 bar (Kidnay and Parrish, 2006). The lean amine solution,

typically an aqueous solution containing 10–65%wt amine, is fed

at the top of the column. As the amine solution falls down the

contactor and mixes with the gas, the acid gases dissolve and

react with the amine to form soluble carbonate salts. The

sweetened natural gas leaves the top of the contactor saturated

with water and so dehydration is normally required before the

gas is sold or fed to a cryogenic gas plant. Process temperatures

inside the contactor rise above ambient temperature due to theexothermic heat of absorption and reaction, with a maximum

temperature observed near the bottom of the column.

The rich amine leaves the bottom of the contactor at a tempera-

ture of approximately 60 1C and containing 0.20–0.81 mol of acid

gas per mole of amine (GPSA Engineering Data Book, 2004). The

pressure of the rich amine stream is reduced to around 6 bar in a

flash tank, to separate any dissolved hydrocarbons from the rich

amine, and preheated to 80–105 1C before entering the stripping or

regenerator column. In the stripping column, heat supplied by a

steam reboiler generates vapour, which removes the CO2 from the

rich amine as the vapour travels up the column. A stream of lean

amine is removed from the bottom of the stripper, cooled to

approximately 40 1C and recycled to the amine contactor. Thevapour stream from the top of the stripping column is cooled to

condense and recover water vapour, and the acid gas may be vented,

incinerated, sent to a sulphur recovery plant (for H2S rich feed gas)

or compressed for reinjection into a suitable reservoir for enhanced

oil/gas recovery (Hughes et al., 2012).

The process disadvantages with conventional amine treating

processes include: (1) the large amounts of energy required for

regeneration of the amine, (2) the relatively low CO2 loading

capacity of amines requires high solvent circulation rates and

large diameter, high-pressure absorber columns, (3) the corrosive

amine solutions induce high equipment corrosion rates, (4) degra-

dation of amines to organic acids, and (5) co-absorption of

hydrocarbon compounds such as benzene, toluene, ethylbenzene

and xylene (BTEX) which subsequently are emitted with the acid

gas stream (Collie et al., 1998; Morrow and Lunsford, 1997).

Operational issues also include solution foaming, emulsions,

excessive solution losses, heat stable salts and high-filter change

out frequency (Seagraves and Weiland, 2011). Aqueous ammonia

and hot carbonate systems are among the alternative chemical

absorption processes to amines. However, many of the disadvan-

tages of amine treating are also associated with aqueous ammo-

nia and hot carbonate processes.

Future innovations in conventional absorption column tech-

nology (e.g., tray and packing designs) could be expected to

achieve only incremental improvements in process efficiencies

(MacDowell et al., 2010). However, one promising strategy to

intensify the CO2 absorption process is the use of a hollow fibre

membrane as a gas–liquid contactor device (Cai et al., 2012;

Ebner and Ritter, 2009; Favre and Svendsen, 2012; Zhou et al.,2010). In this concept, the membrane does not show any

selectivity for CO2 over CH4; instead the membrane provides a

physical barrier between the gas and liquid phases, and a large

interfacial surface area for mass transfer of CO2. The selective

Table 5

Properties of common aqueous amine solvents for acid gas treating.

Solvent Monoethanolamine Diethanolamine Digylcolamine Methyldiethanolamine

Acronym MEA DEA DGA MDEA

Normally capable of meeting H2S specification

(Kidnay and Parrish, 2006)

yes yes yes yes

Removes COS, CS2

, mercaptans (Kidnay and Parrish,

2006)

partial partial partial partial

50 ppm CO2 for cryogenic plant feed (Kidnay and

Parrish, 2006)

no, 100 ppm possible yes, 50 ppmv in

SNEA-DEA

no, 100 ppm

possible

no, pipeline

quality only

Solvent degradation concerns (components) (Kidnay

and Parrish, 2006)

yes - COS, CO2, CS2, SO2,

SO3, mercaptans

some - COS, CO2,

CS2, HCN, mercaptans

yes - COS, CO2,CS2 no

Solution concentrations, normal range wt% (GPSA

Engineering Data Book, 2004)

15-25 30-40 50-60 40-50

Acid gas pickup, mole acid gas / mole amine ( GPSA

Engineering Data Book, 2004)

0.33-0.40 0.20-0.80 0.25-0.38 0.20-0.80

Rich solution acid gas loading, mol/mol amine

normal range (GPSA Engineering Data Book, 2004)

0.45-0.52 0.21-0.81 0.35-0.44 0.20-0.81

Lean solution acid gas loading, mol/mol normal

range (GPSA Engineering Data Book, 2004)

0.12 0.01 0.06 0.005-0.01

Stripper reboiler normal range, 1C (GPSA Engineering

Data Book, 2004)

107-127 110-127 121-132 110-132

Approximate integral heats of absorption of CO2, kJ/

mol Kohl and Nielsen, 1997

84.4 71.6 83.9 58.8



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absorption of CO2 in the liquid solvent occurs at the liquid

interface of the membrane. For the hollow fibre membrane to

operate effectively as a gas–liquid contact device the pores must

remain gas filled and preventing liquid penetration into the

membrane pores is one of the practical challenges hindering the

commercialisation of this technology (Favre and Svendsen, 2012).

A gas–liquid membrane contactor pilot plant is reported to have

been tested in Scotland to remove CO2 from a 5000 Nm3/h natural

gas feed (Mansourizadeh and Ismail, 2009). Other improvementsin the performance of CO2 absorption processes are likely to come

through the development of new solvent materials. Some acid gas

treating process licensors are also working to develop chemical

additives to inhibit corrosion and solvent degradation, so that

amine-based solvents can be operated at higher amine concen-

trations and regenerated at higher temperatures (Normand et al.,

2012). Improved understanding of the thermodynamics and

kinetics of CO2-amine systems, including the development of

mass transfer rate-based modelling approaches, is also allowing

optimised design of absorption processes.

2.1.2. Hot carbonate (alkali salt) systems

Technologies using hot solutions of potassium carbonate

(K2CO3) or sodium carbonate (NaCO3) have been employed sincethe 1950s to remove CO2 from high pressure gas streams (Kohl

and Nielsen, 1997). The overall reactions for CO2 with potassium

carbonate can be represented by (GPSA Engineering Data Book,

2004):

CO2(g)þK2CO3þH2O(l)"2KHCO3(s) (8)

The basic potassium carbonate process was developed by the

US Bureau of Mines and commercialised as the Benfield Process in

1954, and now licensed by UOP with over 700 units constructed

(UOP Overview of Gas Processing Technologies and Applications,

2010). Other commercial potassium carbonate technologies com-

peting with the Benfield Process include the Catacarb Process

(Eickmeter and Associates), which is mainly used in the ammonia

industry, and the Flexsorb HP Process (Exxon Research andEngineering).

The process flow diagram for a potassium carbonate absorp-

tion system shares many features with the general amine process

flow diagram shown in Fig. 2. In a typical hot carbonate process

design the absorber and stripping columns operate in a tempera-

ture range of 100–116 1C (GPSA Engineering Data Book, 2004). If

H2S removal is required or if low CO2 concentrations are required

in the product gas, then alternative designs with a two-stage

contactor and a lean-solution pumped to the middle of the

absorber may be used. For gas treatment requiring CO2 removal

to low levels for cryogenic gas processing, or GTL plant feed, the

UOP process design can be modified to a Hi-Pure design that

combines the potassium carbonate process and an amine process

(UOP Overview of Gas Processing Technologies and Applications,2010; Miller et al., 1999). Amines such as DEA and MEA are also

used as activators to increase the rate of absorption of CO2 in the

potassium carbonate solution (Kohl and Nielsen, 1997). The

Catacarb Process is characterised by the use of a proprietary

organic additives to improve mass transfer rates (Kohl and

Nielsen, 1997). The distinguishing feature of the Flexsorb HP

Process is the use of a sterically hindered amine as the activator

(Kohl and Nielsen, 1997) which is claimed to improve CO2 loading

capacity and mass transfer rates.

2.1.3. Aqueous ammonia solvents

Similar to the acid gas absorption processes using aqueous

amine solutions, processes based on the reaction of CO2 with

ammonia (NH3) in solution have been developed for the capture

of CO2 from natural gas, coal seam gas, and post combustion flue

gases (Darde et al., 2010; Gonzalez-Garza et al., 2009). There are

two variants of the aqueous ammonia process (AAP) reported:

(i) chilled AAP designs operating with absorber temperatures in

the range 0–20 1C and (ii) processes operating with absorber at

ambient temperatures (25–40 1C). Both variants are based on the

same reactions of CO2 and ammonia (NH3) described by Bai and

Yeh (1997), but at low temperatures the chilled AAP allows the

precipitation of ammonium bicarbonate shown in Eq. (9):CO2(g)þNH3(l)þH2O(l)"NH4HCO3(s) (9)

A further advantage of the chilled AAP design is that absorber

operation at low temperatures reduces ammonia slip into the

sweetened gas.

The process flow scheme of an AAP plant is very similar to the

flow scheme for amine absorption cycles (Fig. 2) with an absorp-

tion column and a solvent regeneration system. The absorption

column in the AAP operates at low pressures, usually close to

ambient, and the CO2-rich slurry leaving the bottom of the

absorber column must be pumped to a high-pressure, high

temperature regeneration column (Gal, 2006). In the regeneration

column the ammonium bicarbonate solid can be decomposed to

NH3 and CO2 at temperatures greater than about 50 1C (Dardeet al., 2010; Olajire, 2010), although temperatures of 100–150 1C

are preferred in some designs (Gal, 2006). Typical AAP solvent

concentrations are in the range 13–30% wt NH3 (Kim et al., 2008;

Olajire, 2010).

Although the aqueous ammonia process has potentially lower

energy requirements than amine absorption processes (one study

on AAP for postcombustion CO2 capture suggests 2100–3100 kJ/

kg CO2 compared to 3700 kJ/kg for CO2 capture by MEA, Darde

et al., 2010), the energy savings are not sufficiently large to offset

the additional costs associated with complexity of the ammonia

process (Kohl and Nielsen, 1997) and the need to recompress the

sweetened gas in AAPs. Also, the removal efficiency of chilled

AAPs is only 90% (Gal, 2006), hence ammonia processes may not

be capable of achieving very low CO2 concentrations in productgas required for cryogenic gas processing.

2.2. Physical solvent and hybrid solvent processes

Physical solvent processes may be competitive with amine

absorption when the feed gas is available at high pressure

(generally greater than about 20 bar) or when the acid gas partial

pressure is 10 bar or greater (Nichols et al., 2009). For onshore

natural gas processing facilities all the commercial physical

solvents listed in Table 4 could be used for bulk removal of CO2.

Due to their large plant footprints, physical solvent technologies

are generally not suitable for AGRUs on offshore facilities ( Nichols

et al., 2009). To treat feed gas with very high CO2 concentrations,

the leading physical absorption technologies include the Selexols

and Rectisols processes (Burr and Lyddon, 2008).

The regeneration of physical solvents can be achieved by

reducing the pressure of the rich solvent stream in a series of

multi-stage flash vessels, as shown in Fig. 3, or by stripping the

absorbed gas species in a regeneration column. Importantly, the

heat inputs required for regeneration of a physical solvent are

generally much lower than the heat required for the regeneration

of amine or potassium carbonate sorbents. A potential short-

coming of low pressure regeneration cycles is the cost of recom-

pressing the acid gas if it is to be further processed for CO2sequestration, EOR or sulphur recovery.

The Selexols process, based on a mixture of polyethylene

glycol-dimethyl ethers, is able to remove CO2 simultaneously

with H2S and water (GPSA Engineering Data Book, 2004). In fact,



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H2

S has a greater solubility in most organic solvents than CO2

, a

property that can be used to design H2S selective processes (Kohl

and Nielsen, 1997). As many physical solvents also absorb water,

in contrast to the aqueous amine, AAP and potassium carbonate

technologies which saturate the sweet gas with water, the

required capacity of any dehydration units downstream of the

AGRU may be smaller if a physical absorption process is used.

To produce a sweet gas containing less than 50 ppmv CO2 for

feed to a LNG plant, the Rectisols process using a methanol

solvent operating at temperatures as low as 35 to 75 1C has

been applied successfully (GPSA Engineering Data Book, 2004). A

reported example of a chilled methanol plant is at Riley Ridge,

Wyoming, where 200 MMscfd of ultra-rich CO2 gas (70%) is

treated (CO2 Extraction & Sequestration Project Riley Ridge, WY,

2011).

The main weakness of physical solvent technologies relative toamine AGRUs remains the issue of relatively low acid gas

adsorption capacities of the commercially available physical

solvents. Consequently physical solvent circulation rates are high;

thus large diameter absorption columns and solvent circulation

equipment are required in physical solvent processes. For onshore

gas processing facilities, the capital costs associated with the high

solvent circulation rates may be at least partially offset by the

lower costs of the carbon steel materials required when using

non-corrosive physical solvents compared to the more expensive

materials required to handle the highly corrosive aqueous amine

solutions (GPSA Engineering Data Book, 2004).

Several mixed-solvent (also known as hybrid solvent) gas

treating processes combine the effects of physical and chemical

absorption processes in a single operation. The most well-knownmixed-solvent processes for CO2 absorption are the Sulfinol-D

s

process licensed by Shell Global Solutions (Rajani, 2004) and the

Amisol process licensed by Lurgi (Kohl and Nielsen, 1997). The

Sulfinol-Ds process based on a mixture of Sulfolane (tetrahy-

drothiophene dioxide), DIPA (diisopropanolamine) and water is

capable of deep removal of CO2 to less than 50 ppm from feed gas

containing very high concentrations of CO2 (Kohl and Nielsen,

1997). Another variation of the Shell Global Solutions technology

is the Sulfinol-Ms process, which is also based on Sulfolane but

uses MDEA instead of DIPA, used mainly for selective removal of

H2S. The process flow scheme of the Sulfinol-Ds process is

essentially the same as that for an amine absorption process

(Fig. 2) with the addition of a flash tank to remove the bulk of the

acid gas from the rich solvent upstream of the stripper column.

The Amisol process is based on a mixture of methanol, water and

either diethylamine (DETA) or diisopropylamine (DIPAM). This

process has most commonly been used for purification of synth-

esis gas derived from coal, peat, or heavy oils (Kohl and Nielsen,

1997). The advantages of the hybrid solvent technologies over

conventional amine absorption technologies include low energy

consumption for regeneration of the solvent, high acid gas loading

capacities, low foaming tendency, and reduced corrosion. Hybrid

solvent processes are usually only suited for treatment of natural

gas with an acid gas partial pressure of more than 100 kPa. The

main drawback of the mixed-solvent processes is that hydrocar-

bon losses (to the solvent) are slightly higher than the typical

losses in conventional amine processes.

2.3. Ionic liquids and switchable solvents

Among the materials investigated as new solvents for CO2absorption processes, ionic liquids (ILs) are one of the solvents

that may in the future offer an alternative to amines and the low

capacity physical solvents. Ionic liquids are commonly defined as

organic salts with melting temperatures of less than 373 K. They

have a range of properties that may make them useful replace-

ments for volatile organic solvents such as extremely low vapour

pressure, nonflammability, and in many cases low toxicity (Zhao,

2006). Furthermore, ILs are often considered to be designer

solvents because of the many different possible combinations of

cations and anions which can be used to tune their chemical and

physical properties (Brennecke and Gurkan, 2010). These proper-

ties of ILs have generated great scientific interest in their devel-

opment and investigation into their use in chemical engineeringapplications, including gas separations. Reviews by Zhao (2006),

Plechkova and Seddon (2008) and Werner et al. (2010) describe

the present industrial processes which use ionic liquids and

discuss many other possible industrial applications of ionic

liquids. The application of ILs to CO2 capture and natural gas

sweetening has been discussed by many authors, including Bara

et al. (2010b), Karadas et al. (2010), Brennecke and Gurkan (2010)

and MacDowell et al. (2010).

The development of ILs for CO2 capture and natural gas sweet-

ening has been driven by the desire to develop a solvent with a CO2capacity comparable to that of amine-based solvents but with

greatly reduced energy requirements for regeneration. Initially,

much of the research focussed on the potential of imidazolium-

based ILs as alternative physical solvents. Early studies focussed on

Fig. 3. Process flow diagram of a typical physical solvent process for absorption of CO2 and other acid gases from natural gas.



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the most convenient ILs to synthesise, such as 1-n-alkyl-3-methy-

limidazolium [Cnmim] cations paired with various anions including

bis(trifluoromethylsulfonyl)amide [Tf 2N] (Bara et al., 2010b; Hughes

et al., 2011). However, the alkyl chains (in the imidazolium cation)

are not the ideal functional group for separating CO2 from CH4 or N2and, thus, these alkyl chains were substituted with various groups

containing either ethylene glycol or nitrile units to form

[Rmim][Tf2N] (Bara et al., 2010b). Karadas et al. (2010) summarised

the results of using anions other than [Tf 2N] on the solubility of CO2in various ILs, with anions containing fluorinated derivatives show-

ing a modest increase.

Despite these efforts to optimise the molecule, the solubilities

achieved with ILs on a volumetric basis (dissolved moles of CO2 per

liquid volume) have remained comparable with most common

organic solvents (Bara et al., 2010b; Karadas et al., 2010). Much

research has thus focussed on the development of task specific ionic

liquids (TSIL) which incorporate an amine functional group into the

IL, enabling it to serve as a reactive, chemical solvent for CO2. Bates

et al. (2002) reported an imidazolium-based TSIL containing an amine

functional group attached to one of the alkyl chains. The stoichio-

metry of the absorption reaction achieved was 0.5 mol of CO2 per

mole of TSIL; regeneration of the IL solvent was also achieved by

heating the carbamate product to 80–100 1C under vacuum. Gurkan

et al. (2010) and Zhang et al. (2009) used TSILs with amino-acid

functional groups that improve the stoichiometry to nearly 1 mol of

CO2 per mole of IL, which is important if the IL solvent is to be

competitive with the volumetric capacities of aqueous amine solu-

tions (Brennecke and Gurkan, 2010). However, TSILs can be difficult

to synthesise and have large viscosities at ambient temperature,

which increase further upon complexation with CO2 (Brennecke and

Gurkan, 2010; Karadas et al., 2010). Brennecke and co-workers

suggest that the viscosity increase of TSILs upon complexation with

CO2 can be eliminated through the use of aprotic heterocyclic anions

and have filed a provisional patent (Brennecke and Gurkan, 2010).

Nevertheless, the synthesis and viscosity challenges associated with

TSILs currently limit their commercial viability.

Currently, the most viable method of applying ILs to CO2 capture

or natural gas sweetening is the use of IL þamine mixtures, in whichMEA or DEA is dissolved in an [Rmim][Tf2N] solvent; such solutions

can have up to 116 times the CO2 solubility on a volumetric basis

than the IL alone (Bara et al., 2010a). Camper et al. (2008) found that

MEA-IL and DEA-IL solutions could rapidly and reversibly capture

1 mol of CO2 per mole of amine and thereby reduce feed gas CO2concentrations to the ppm level, even at CO2 partial pressures below

0.133 kPa. Such IL þamine solvents offer significantly reduced

energy requirements relative to conventional aqueous amine sol-

vents: Bara et al. (2010a) compared a IL-amine process with flash

regeneration (similar to the process flow scheme shown in Fig. 3),

against a conventional amine-based gas sweetening plant (similar to

that shown in Fig. 2). They considered the sweetening of 100

MMSCFD to a sales gas specification of 2% CO2 for an NG feed

containing either 15% or 5% CO2 and concluded that over a 20 yrplant life cycle, the amineþIL process had combined CAPEX and

OPEX savings of about 25%. This was due primarily to the removal of

a regeneration column, the higher amine loadings and lower solvent

circulation rates of the amineþIL solvent, and the reduced duties

required to cool, heat and regenerate the amineþIL solvent. The

difference in the calorific properties of water and the IL, and in

particular the duty reduction associated with not vaporizing any

solvent during regeneration, is probably the most significant advan-

tage of the amineþIL solvent approach to gas sweetening. A pilot

scale unit is under construction for planned NG sweetening field

tests in 2011–2012 (Bara et al., 2010a).

Brennecke and Gurkan (2010) point out that all ILs are

hygroscopic (that is, a material that can adsorb or absorb water

molecules), even those that are usually designated hydrophobic

because they are insoluble in water. This fact represents a serious

challenge for the application of ILs to CCS but it also has

implications for NG sweetening. If the NG contains some water,

then the absorption of that water by the IL could decrease the

solvent’s capacity, and will also degrade the reduction in regen-

eration duty associated with an amineþIL solution. Thus in

contrast with conventional gas processing practice, a sweetening

process utilising an amineþIL solvent should probably be situ-

ated downstream of the dehydration process. Jessop and co-workers reported the development of ‘switch-

able’ solvents ( Jessop et al., 2005; Phan et al., 2008) where a basic,

non-polar liquid mixture converts into a polar IL upon the

addition of CO2. Several groups are researching the optimisation

of these switchable solvents for improved CO2 capture or NG

sweetening. For example, Heldebrant et al. (2011) reported the

development of 2nd generation switchable solvents that could

capture of nearly 1.3 mol of CO2 per mole of solvent. In these

cases, the viscosity of resulting IL is appreciable and its reduction

is the focus of ongoing research.

2.4. Absorption processes for N 2 rejection

Nitrogen rejection using absorption-based technologies is not

a common practice in the natural gas industry. There are several

commercial N2 rejection processes that operate by the physical

absorption of CH4 in a hydrocarbon oil which have been built to

process gas feed rates of 2–30 MMscfd (AET-Technology, 2007;

TGPE, 2009). The costs of large solvent rates and CH4 recompres-

sion are prohibitively high for the use of CH4 selective absorption

NRUs in large scale LNG production plants; thus there is a

growing need for the development of a N2 selective absorbent.

An example of a CH4 selective NRU process is that designed

and constructed by Advanced Extraction Technologies, Inc. (AET,

Houston USA). The flow scheme of the AET process is similar to

the physical solvent process shown in Fig. 3. In a typical CH4selective absorption process the feed gas is cooled before entering

the absorption column where CH4 is dissolved into a lean oil

solvent (TGPE, 2009). The CH4 is then recovered from the richsolvent through a series of flash vessels to produce sales gas

containing less than 4% N2. Th e N2 from the feed gas flows

through the top of the column and is used for precooling the unit

feed, and is then vented. The AET process operates at approxi-

mately 32 1C to optimise the CH4 absorption in the solvent.

However, the capacity of most solvents for CH4 is relatively low

and large solvent circulation rates are required to achieve eco-

nomic recovery of the CH4. Furthermore, because the CH4 is

recovered from the solvent by flashing at low pressure the gas

must be recompressed for pipeline or cryogenic gas plant feed.

Although most solvents have a greater solubility for CH4 than

N2, some organo-metallic complexes (OMCs) have been reported

to preferentially bind N2 from natural gas mixtures. Reversible

transition metal complexes forming with N2 were first reported inthe scientific literature by Allen and Senoff (1965) who described

the reversible reaction of hydrazine (a N2 source) with ruthenium

chloride. Several scholarly articles describing the preparation and

N2 absorption capacity of these types of transition metals forming

complexes were also published in the 1970s (Allen et al., 1973;

Chatt et al., 1978; Sellmann, 1974). However, there have been few

studies reported on the application of these N2 binding complexes

to natural gas processing. The most extensive applied studies of

N2 selective solvents are reported by Stanford Research Institute

International (SRI), on contract for the US Department of Energy

(Alvarado et al., 1996), and Bend Research Inc. (Friesen et al.,

2000; Friesen et al., 1993). Both the SRI process and the Bend

process exploit the reversible chemical complexing abilities of

multi-dentate transition metal complexes as shown in Fig. 4.



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The absorption of N2 in OMCs is achieved through using metal

ions with six coordination sites, the four equatorial sites complexed

with bi, tri, or tetra dentate ligands; with an electron withdrawing

ion on one axial site leaves the other axial site ready to complex

with a gaseous N2 molecule in an end-on configuration. The strength

and therefore reversibility of the complex formation is dictated

through the use of different ligands and transition metals. The most

promising OMC reported by SRI was a (bis)tricyclohexylphosphine

molybdenum tricarbonyl, in a toluene solution, which formed a

yellow precipitate when bound with a N2

molecule (Bomberger

et al., 1999). This phosphine complex absorbed up to 0.12 mmol of

N2/mL of solution at 1979 kPa (Alvarado et al., 1996) in equilibrium

measurements. However, during absorption–regeneration cycles

measured on a 0.02 MMscfd bench-scale apparatus the SRI research-

ers encountered issues with the degradation of the phosphine

complex and regeneration of the complex to release the N2. The

problems with the regeneration of the phosphine complex included

difficulties in controlling the size of the N2-bound precipitates,

which would bypass the regeneration system if the size was less

than that of the phosphine solids (412mm) (Bomberger et al.,1999). Although SRIs economic analysis suggests that this OMC

process could become cost competitive with cryogenic NRUs, the

challenges to reduce the cost of OMC synthesis, improve the stability

of the phosphine complex in the presence of water and oxygen, and

to overcome the problems encountered with solids handling in theregeneration process are all significant.

The patents held by Bend Research Inc. (Friesen et al., 2000,

1993) describe OMCs based on several transition metal–ligand

combinations including complexes with iron, as well as the perfor-

mance of these complexes in N2 absorption–regeneration cycles. The

patents report N2 uptakes of 0.5 mol N2/mol OMC at a N2 partial

pressure of approximately 1130 kPa and a selectivity for N2 over CH4close to 6. These complexes are reported to exhibit a high degree of

stability, showing no decrease in capacity through repeated use over

100 days, and exposure to an atmosphere of 3% CO2 and 100 ppm

O2. Although the publically available information on the Bend

process indicates the concept of N2 absorption in these OMCs may

be sound, no commercial process has been developed for this

technology. The main barriers to the use of OMCs in large scalegas processing operations remain the high cost of synthesis of the

OMCs and improving the chemical resistance of the complexes to

common gas contaminants such as water and H2S. Furthermore, the

Bend Research Inc. iron-based OMCs may present safety and

materials handling issues because these are pyrophoric compounds,

which can ignite spontaneously on contact with air.

3. Condensation, desublimation and distillation

3.1. N 2 rejection by cryogenic distillation

The normal boiling point (NBP) of CH4 is 161.5 1C and at a

typical pressure of 3150 kPa in the intermediate stages of the

cryogenic LNG plant the boiling point (BP) of CH4 is 94.7 1C

(Lemmon et al., 2010), which provides a sufficient difference in

relative volatility with N2 (NBP 195.8 1C, BP of 148.7 1C at3150 kPa) for separation of CH4þN2 mixtures by cryogenic

distillation. Relevant to such separations are vapour–liquid equi-

librium curves, such as the one shown in Fig. 5 for N2–CH4 at

conditions accessible in a cryogenic LNG plant. The equilibrium

curve in Fig. 5 represents a typical relative volatility aN2,CH4 valueof 5–8. To illustrate the separation of N2 and CH4 in the cryogenic

distillation column, we have applied the McCabe–Thiele method

(McCabe and Thiele, 1925) to estimate the number of ideal

equilibrium stages required to produce a liquid CH4 bottoms

product of 95% from a feed containing 50% N2 and 50% CH4. These

feed and production compositions were selected here to allow

illustration of the separation process; in a real process to produce

a CH4 containing less than 1% N2 and with a high rate of CH4

recovery, the number of actual equilibrium stages required in the

Fig. 4. General chemical scheme for the reversible binding of nitrogen with an organo-metallic complex adsorption (Miller et al., 2002). R ¼generic organic functionality.

Fig. 5. Illustration of the binary CH4–N2 vapour–liquid equilibrium relationship

and the construction of a McCabe–Thiele diagram to calculate the number of ideal

equilibrium stages for separation by distillation. The vapour–liquid equilibrium

data shown here represents a relative volatility, aN2CH4 , of 7 (calculated using

REFPROP (Lemmon et al., 2010) for an operating pressure of 2757 kPa. In this

hypothetical example, the feed contains 50% N2 with product specifications of 5%

N2 in the CH4-rich bottom product and 5% CH4 in the N2-rich overheads product.

These compositions were selected as an example which could be shown clearly on

this figure and do not represent a typical set of operating conditions. To meet more

realistic processing objectives, such as a higher purity CH4-rich bottom product

and 98% recovery of the methane from the feed gas, a much larger number of

equilibrium stages would be required.



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distillation column would be higher than the eight ideal trays

shown in the McCabe–Thiele construction on Fig. 5. (The corre-

sponding graph would be more difficult to read.) In practice,

modern cryogenic NRUs (with SP N2,CH4 320) can produce very

high purity CH4 at high recovery rates, which also reduces the CH4content in the N2 overheads vapour to less than 3%. This is

achieved through the use of columns with large numbers of

stages and the design of systems with multiple columns.

Currently, cryogenic distillation is the only N2 rejection methodthat has been demonstrated at gas flows above 25 MMscfd to

achieve very high methane recovery (typically above 98%) and high

purity N2 (approximately 1% CH4). The selection, and optimised

design, of a cryogenic NRU is principally determined by the

concentration of N2 in the feed gas (Wilkinson and Johnson, 2010).

The feed gas pressure, feed gas flow rate, concentration of con-

taminants and product specifications (sales gas or LNG) also

influence process selections. The optimum design of cryogenic

NRUs, like for most cryogenic processes, is an exercise in balancing

the energy efficiency and process flow sheet integration to reduce

the power consumption required for compression of the CH4refrigerant loops, which are used to provide the reboiler and

condenser duties (Finn, 2007; Wilkinson and Johnson, 2010). The

compression requirements of this distillation-based separation are

the largest contributor to capital and operating costs of the NRU

process. Cryogenic NRU processes have been constructed by most of

the major process designers such as Linde, Costain, Praxair, Con-

ocoPhillips and APCI. The main variants of cryogenic NRU designs

are (1) the single-column heat-pumped process, (2) the double-

column process (Agrawal et al., 2003) and (3) three or two column

designs featuring a prefractionation column (MacKenzie et al., 2002;

Wilkinson and Johnson, 2010).

A typical single-column heat-pumped NRU process is illu-

strated in Fig. 6. Upstream of the cryogenic NRU, contaminants

that could freeze at cryogenic temperature such as water, CO2 and

heavy hydrocarbons have been removed from the gas. The feed

gas to the NRU is cooled, throttled and fed to an intermediate

stage of the distillation column operating at pressures from 1300–

2800 kPa (Agrawal et al., 2003). Rejected N2 vapour (typicallyo1% CH4) is drawn from the column overheads and the CH4-rich

liquid product is drawn from the bottom of the column. The

bottoms product then can be reheated against the NRU feed gas.

A closed-loop CH4 heat-pump cycle driven by an external com-

pressor provides the reboiler and condenser duties, with the

closed-loop CH4 condensed at a high pressure in the reboiler

and revaporised at low pressure in the condenser.

As the N2 content in the feed gas increases, the CH4 in the upper

stages of the column becomes more difficult to condense. The

operating flexibility of a single-column NRU process is limited by

(1) the critical pressures of nitrogen–methane mixtures, which

limits the maximum pressure of the distillation column to approxi-

mately 2800 kPa, and (2) the minimum practical temperature of CH4

after the throttling valve of the heat-pump cycle (MacKenzie et al.,2002). These limitations mean that the single-column NRU process

is, generally, used for feed gases containing less than 20% N2.

A double-column NRU can provide additional process flexibility

compared to the single-column process to allow the separation of

gases containing higher N2 concentrations or gases in which the feed

gas quality varies. In the double-column N2 rejection process the

NRU feed gas is cooled, throttled and fed to a high pressure (HP)

column operating typically at 1000–2500 kPa (Agrawal et al., 2003).

Having had some of the N2 removed, the crude natural gas liquid

stream from the bottoms of the high pressure column is sub-cooled,

throttled and fed to the low pressure (LP) column (operating at

approximately 150 kPa). In practice both the HP and LP columns are

usually integrated into a single tower to improve process heat

integration and minimise heat transfer to the atmosphere. The

N2-rich vapour from the HP column is condensed to provide refluxfor both the high pressure column and the low pressure column. The

low pressure column produces a high purity N2 stream (o1% CH4);

which is used to cool N2 reflux fed to the LP column in the LP

condenser, sub-cool the bottoms of the HP column and to pre-cool

the feed gas. The CH4-rich bottom product of the LP column is

pumped through the crude sub-cooler, which provides the LP

column’s reboiler duty by condensing the overheads of the HP

column, vaporised and reheated against the feed gas. Alternatively,

for LNG production the CH4-rich bottoms product of the LP column

may remain liquid.

The three-column NRU process consists of the double-column

process described and a prefractionation column. The prefractio-

nation column recovers some of the hydrocarbons at a higher

temperature than the double-column system and increases the N2concentration of the feed gas. Importantly, the prefractionation

column reduces the volume of N2-rich gas that must be processed

at low temperatures in the main NRU column(s) and this reduc-

tion in gas volume can significantly reduce the power require-

ments of the NRU (MacKenzie et al., 2002; Wilkinson and

Johnson, 2010). A further variant of an NRU with a prefractionator

described by MacKenzie et al. (2002) is a two column process

with a high pressure prefractionation column, an intermediate

liquid–vapour separator and a low pressure column.

3.2. Low-temperature CO 2 removal processes

The separation of CO2 from natural gas by low-temperature

processes (operating at temperatures below 0 1C) can be

Fig. 6. A process flow schematic of a typical single-column N2 rejection unit

(schematic adapted from GPSA Engineering Data Book, 2004; Agrawal et al., 2003;

MacKenzie et al., 2002).



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categorised as: (1) gas–liquid phase separations operating at

temperatures above the CO2 triple point temperature of

56.6 1C (Lemmon et al., 2010) and (2) gas–solid phase separa-

tions where desublimation of CO2 occurs at temperatures below

the triple point. Although the term ‘‘cryogenic’’ is often used by

the vendors and in the literature to describe these types of CO2capture technologies, most of the processes operate at tempera-

tures above the scientific definition of cryogenic as 153 1C

(Agrawal et al., 2003; Radebaugh, 2007). To overcome theproblems associated with the formation of CO2 solids during

cryogenic distillation two technological approaches have been

pursued: (1) extractive distillation by the addition of a heavier

hydrocarbon to alter the solubility of components in the column

(Ryan/Holmes process) and (2) controlled freezing and re-melting

of the solids (Controlled Freeze ZoneTM and CryoCells processes).

Other low-temperature CO2 removal technologies under devel-

opment include systems in which mechanical methods are used

to separate the CO2 rich phase from the natural gas. For example,

Willems et al. (2010) report the C3sep (condensed contaminant

centrifugal separation) process in which condensed CO2 droplets

are separated from the natural gas using rotational separators,

and Clodic et al. (2005) describe the ALSTOM process which

features a multi-stage thermal swing process that freezes then

melts CO2 on mechanical fins. We focus our discussions in this

review on the commercialised Ryan/Homes process and the pilot-

plant demonstrated Controlled Freeze ZoneTM (along with the

similar CryoCells process).

The extractive distillation approach to solving the problem of

CO2 freezing in CH4–CO2 distillation is most well known through

the Ryan/Holmes process described in the 1982 US Patent

4,318,723 (Holmes and Ryan, 1982). The Ryan/Holmes process

is representative of several similar technologies patented by

various other inventors. The addition of a heavier hydrocarbon

stream (typically a C2–C5 alkane) to the condenser of the distilla-

tion column shifts the operation away from conditions that favour

solids formation, because the solubility of CO2 in the liquid phase

can be increased, the overheads temperature can be raised, and

the column can be operated at a higher pressure since themixture’s critical pressure increases. A typical four column

Ryan/Holmes process configuration incorporates a de-ethaniser

column, a CO2 recovery column, a demethaniser column, and a

column for recovery of the hydrocarbon additive. Further details

on Ryan/Holmes configurations and operating issues for separa-

tions of methane–CO2, ethane–CO2, and CO2–H2S are discussed in

the GPA Engineering Data Book.

The Controlled Freeze ZoneTM (CFZTM) process was first patented

by ExxonMobil in 1985 (Valencia and Denton, 1983) and tested in a

Texas pilot plant during 1986–1987 (Nichols et al., 2009). More

recently a commercial demonstration project designed to treat a

feed gas of 14 MMscfd has been constructed in LaBarge, Wyoming

(Controlled Freeze ZoneTM—increasing the supply of clean burning

natural gas, 2010). The separation tower of the CFZTM process is splitinto three sections with an upper rectification section and a lower

stripping section (both conventional distillation sections) separated

by the CFZTM section, as shown in Fig. 7 (Fieler et al., 2008). In the

CFZTM section, the liquid falling from the rectification section is

contacted with a cold methane stream (90 to 85 1C), which

causes the CO2 to freeze out of the methane mixture. The CO2 solids

(62 to 45 1C) drop to a liquid layer on a melt tray in the lower

stripping section; the solids melt before falling as liquid through the

downcomers of the melt tray. The standard CFZTM process can

produce pipeline quality gas, and when implemented with a

modified rectification section is claimed to be capable of producing

a sweet gas of less than 50 ppm CO2 (Nichols et al., 2009).

Cool Energy’s CryoCells was developed by researchers at

Curtin University in Western Australia with industrial partners

Woodside Petroleum and Shell Global Solutions (Hart and

Gnanendran, 2009). Like the CFZTM process, the CryoCells process

operates by the controlled freezing and subsequent remelting of

CO2. The basic thermodynamic path for the CryoCells operation

involves cooling a dry, feed gas (at 5600–6600 kPa) to just above

the CO2 freezing point (for example to 60 1C) to condense some

or all of the vapour, followed by an isenthalpic flash to furthercool the mixture to obtain solids, liquid CO2 and a CH4-rich

vapour. Pilot plant trials of the CryoCells process demonstrated

the production of

The Removal of CO2 and N2 From Natural Gas Review (Rufford Et Al 2012)

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