Integration of Cryogenic Air Separation and IGCC with Carbon Capture

8/11/2019 Integration of Cryogenic Air Separation and IGCC with Carbon Capture

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Integration of Cryogenic Air Separation and IGCC with Carbon

Capture

2.60/2.62, Spring 2010

Richard Bates and J. Fernando Rodrguez


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Introduction

Society stands at a crossroads where action to reduce carbon emissions must be taken to

avoid irreversible damage to the environment. Unfortunately the energy sources the world dependsthe most on also happen to be the most polluting. When asked to think of potential solutions, some

might suggest the increased implementation renewable energy sources like wind and solar. Few

might propose that improved methods of oxygen production will be critical to reducing the carbon

impact of future energy conversion systems. In fact, oxygen production (also known as air

separation) is a major energetic and capital cost associated with proposed carbon-free coal power

technologies. Two routes exist to improving the viability of such power plants. The first is research

and analysis to directly improve either air separation technologies or clean coal plants separately.

The other route involves identifying specific ways how the integration of the two processes can

increase the overall efficiency. By using simplified but robust thermodynamic modeling of the two

technologies we show how integration of cryogenic air separation (CAS) with integrated

gasification combined cycle (IGCC) coal plants can increase overall plant efficiency by 3 percentage

points.

The first section of the paper will survey and review existing and developing air separation

technologies as well as proposed carbon-free coal power technologies which utilize oxygen. For

each technology we provide a basic explanation of operation as well as address its associated

technical challenges. In the second section of the paper we focus in on the benefits of integrating

CAS and IGCC by comparing the modeled efficiency for an un-integrated base case versus an

integrated case.

1. Summary of Current & Developing Air Separation & Clean

Coal Technologies

Air Separation

Mechanical Separation Methods

Absorption

Absorption methods make use of materials with preferential absorption. These materials

are usually preferential nitrogen absorbents. The performance of these materials relies on their

molecular structure, attracting the nitrogen molecules more strongly than oxygen or argon

molecules.


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After passing through the absorption material, the oxygen-rich flow is passed through a

carbon-based molecular sieve. These types of sieves contain molecular pores with pore sizes

preferential to the quick diffusion of oxygen molecules, ultimately producing an oxygen output of

acceptable purity.

The materials used to preferentially select nitrogen are called zeolites. These materials are

usually of igneous origin, with highly regular molecular structure. Zeolites are mined in open pit

mines in areas of New Mexico, Arkansas, and Idaho, and are mostly exported to China for concrete

production. Even with a strong national supply, the leaders in zeolite production come from

Europe, Asia, and Australia.

As the absorption method is still in development, the importance of zeolites in North

America is beginning to be realized, and its applications in the air separation industry are being

more thoroughly researched and appraised.

Figure 1: Schematic of a general absorption plant. Filtered air is passed through zeolite beds,where nitrogen is removed, until the bed is saturated with nitrogen. Then the airflow is diverted to anew bed. Each saturated bed must then be regenerated, which can be done in two ways: Byincreasing the bed temperature or by reducing the pressure in the bed. Both processes reduce theequilibrium absorption capacity of the zeolite bed, which releases the trapped nitrogen and readiesthe bed for reuse. The heat addition method is commonly known as temperature swing absorption(TSA) and the pressure reduction is known as pressure or vacuum swing absorption (PSA or VSA).Usually, the PSA method has a faster turnaround time and its operation relatively more simple thantemperature change. Thus, PSA is the industry preferred method for air separation. [Smith, et. al,

2001]

Several variations in the process can be introduced that can affect the operating efficiency.

These include the inclusion of air pretreatment units to remove water and carbon dioxide. Multiple

zeolite beds can optimize the energy recovery during bed switching and vacuum operations.


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The process optimization for air separation is based on the output purity and pressure,

energy costs, and expected operating life. Oxygen purity can be expected to be at 93-95 vol%. Due

to the reliance of the absorption method on the zeolite beds, their size is a controlling factor in the

capital cost of each facility. Production is proportional to the zeolite bed volume, so cost increases

more rapidly with increasing production capacity than it would for cryogenic plants.

Polymeric membrane separation

Membrane separation methods take advantage of polymeric materials with different

diffusion rates for oxygen and nitrogen which separates high and low pressure streams. There are

two factors that affect the cost of membrane systems, flux and selectivity. Flux determines the

membrane surface area, and it is related to the pressure difference across the membrane material.

Selectivity is defined as the ratio of the permeabilities of the gases to be separated, which in turn

depends on the molecule size. Since oxygen is a smaller molecule than nitrogen, most membrane

materials have a higher permeability for oxygen. Membrane methods are however limited to the

production of oxygen-rich air, which can range from 25 vol.% to 50 vol.% pure.The most important component of the membrane separation process are the membranes

themselves. They are usually based on polymeric materials, and might be supplemented with

oxygen-complexing materials that aid in the absorption and separation of oxygen from the air

stream. However, this type of supplemented membrane materials is still in development, and

depend on the chemical compatibility between the complexing material and the membrane

material.

The cost of membrane separation plants is mostly linearly related to the production rate,

which can reach up to 20 tons/day in applications where water and carbon dioxide contaminants

can be tolerated. Improvements in membrane materials are crucial for the improvement of the

process quality. However, this is an attractive option for its simple, continuous operation with a

fast startup time.

Ion transport membrane (ITM) separation

ITM separation makes use of solid inorganic oxide ceramic materials that produce oxygen

by separating oxygen ions from air as it passes through the ceramic microstructure. These systems

operate at high temperatures, usually 1100F. Oxygen molecules are ionized at the membrane

surface and transported across it by an applied electric voltage or oxygen partial pressure

difference. The oxygen ions are reformed after passing through the membrane material.

For large processes, the pressure difference method is preferred over the electric potentialmethod. The membranes operate through a pressure difference, transporting oxygen ions and

electrons, and thus they are referred as mixed transport conducting membranes. The ITM

processes require mixed conducting membranes that are able to operate at the high temperatures

involved in this process.


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Figure 2: Schematic of a general membrane separation plant. A major benefit of membraneseparation processes is in their simplicity, and their operation near atmospheric pressures. The feed

air needs only to be pressurized to overcome the head losses through the filters, piping andmembrane materials. A vacuum pump is used to maintain the pressure difference across themembrane material. In the separation process, the oxygen rich air is also loaded with carbon dioxideand water impurities, since they are usually more permeable than the oxygen for most membranematerials. [Smith, et. al, 2001]

This method is capable of producing very large production rates and separate oxygen as a

nearly pure product. The oxygen gas can then be mixed with a sweep the permeate side of the

membrane to create a lower purity product, or reacted with other gases for use in other processes,

such as a natural gas methane sweep to make synthesis gas for gas-to-liquid applications.

This method is thus well suited for applications that can be integrated with power

generation and energy conversion processes that require pure oxygen as a feedstock forcombustion or even gasification, such as the example given for GTL processes. It is very efficient

and produces a nearly pure oxygen stream.


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Figure 3: General schematic of a ITM separation plant. This process takes in air, then filters itto remove contaminants such as water and carbon dioxide. The purified air is then heatedthrough heat exchangers and with supplemental heat. The hot air is then passed through theITM units to separate it into oxygen and nitrogen streams. The hot product gases are then cooledoff by exchanging heat with the incoming purified air flow. The flows are then processed for use,oxygen is compressed for commercial use and nitrogen can be expanded in a turbine to generate

power. [Smith, et. al, 2001]

Chemical Separation Methods

Molten salt separation

Chemical separation methods take advantage of materials that are able to absorb oxygen at

one thermodynamic state and release it at a different state. For example, a molten salt process uses

chemical oxidation of the molten salt at one temperature and pressure and releases it by pressure

and temperature reduction.

This process makes use of molten salt mixtures which need to be reactive with oxygen and

also release the oxygen at another thermodynamic state. The MOLTEX process is known to be

used in this type of chemical air separation. However, the use of salts leads to equipment corrosion,

which is one of the main concerns of this process

A major process advantage comes from the low pressurization requirements of the process.

The air only needs to be pressurized to overcome the pressure losses through the air pretreatment

and heat exchanger stages. This requires relatively less compression power when compared to

cryogenic processes. However, a source of thermal energy must be available for releasing the

oxygen from the molten salt via heating. Process efficiency in terms of oxygen production has been

shown to produce 99.9 vol.% pure oxygen. However, the molten salt-oxygen reaction product is

highly corrosive and a major problem in the development of the molten salt absorption process.

Figure 4: Schematic of a general chemical separation plant. The air is compressed andtreated to remove water and carbon dioxide. This is accomplished by flowing the air past


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molecular sieves. The sieve beds are regenerated using dry nitrogen, a byproduct of theseparation process. The purified air is then heated to temperatures ranging from 900F and1200F. This hot air is then passed through to the molten salt absorbers. The oxygen in theair then reacts chemically with the molten salt. The salt is processed by heating it to highertemperatures before it is reduced in pressure. This pressure reduction releases gaseousoxygen, and the salt is regenerated and ready for use. The hot products and byproducts are

cooled off by the incoming purified air flow, and then the oxygen is compressed tocommercial delivery pressures, while part of the nitrogen is used to regenerate thesaturated molecular sieves used to remove water and carbon dioxide. [Smith, et. al, 2001]

Cryogenic Separation Methods

Cryogenic separation methods work through six major operations to separate air into useful

products. First, air is compressed from atmospheric pressure to approximately 6 bar. The air is

then filtered to remove suspended matter in the feed air, such as dirt, with the use of mechanical

filters. The second step in cryogenic separation is air cooling and purification. In this process

carbon dioxide and water, and any other dissolved contaminants in the air are removed. The

filtering processes make use of molecular sieves to separate the carbon dioxide, water and

hydrocarbon contaminants.

The feed air is then cooled against incoming the nitrogen byproduct. The third step in the

cryogenic separation process is a low temperature heat exchange to lower the air temperature to

near-liquefaction states. Waste nitrogen byproduct is again used to cool the incoming purified feed

air stream. The fourth step in the process takes the feed air through further compression cycles,

and then through expansion and liquefaction cycles. The cryogenically cooled air is then passed

through rectification columns to separate the gases because of their differing partial pressures.

Figure 5: General cryogenic air separation procedure. There are various types ofcryogenic separation processes. These are dependent on whether only oxygen is necessary,or whether pure nitrogen and argon will also be commercially produced. [Smith, et. al, 2001]


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Technical Comparison

Below is a summary of the existing and upcoming technologies, along with some of the most

important design considerations. Economic range is given in short tons per day available for each

type of plant. The byproduct capability describes the ability of the process to produce nitrogen and

argon if they are desired. The purity limit refers to the purity of the oxygen output, while the start-uptime is the time that the plant needs to reach steady state production at the given purity limits.

Comparison of existing and emerging technologies [Smith, et. al, 2001]

Process StatusEconomic

range (sTPD)

Byproduct

capability

Purity limit

(vol.%)Start-up time

Absorption semi-mature 20 excellent 99+ hours

Membrane semi-mature


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(c)

Oxy-fuel combustion CO2capture-Also can be applied to existing plants. The coal is

combusted in a mixture of pure oxygen diluted with an external recycle stream (CO2,

H20).Figure 6 shows a brief schematic of the three methods of carbon capture

Though post-combustion capture seems like a simple solution which could be applied to

both new and existing plants the main disadvantage to post-combustion CO2captureis that CO2is

present at dilute concentrations (13-15 percent by volume in coal-fired systems and 3-4 volume

percent in gas-fired turbines) requiring the treatment of large gas volumes and energy intensive

separation processes. Moreover, nitrogen oxides, particulate matter, and sulfur dioxide in the flue

gas can potentially degrade certain sorbents and reduce the effectiveness of certain CO2 capture

processes3.

Figure 6- Schematics of Carbon Capture Technologies4

3SECA, 164SECA,13


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Pre-combustion and oxy-fuel combustion technologies also have their challenges. But since

they both utilize pure oxygen it is more relevant to discuss these two options in depth. In the

following section we'll examine the advantages and disadvantages of each.

Oxy-Combustion CO2capture

Technology Summary

Coal (or natural gas) is combusted in a mixture of oxygen (typically of greater than 95%

purity) and recycled flue gas. The resulting flue gas, consisting mainly of CO2and water, does not

require additional separation of the CO2. After the H2O is condensed, the remaining CO2is ready for

sequestration. The recycled flue gas is used to control flame temperature and make up the volume

of the missing nitrogen.5

An important factor in designing the oxy-combustion systems is the proportion of oxygen to

recycle gas required. In order to maintain the same heat transfer and temperature characteristics as

air-combustion furnace this mixture must be optimized. Depending on whether the water is

recycled the fraction of oxygen injected into the combustor proportion is between 24-27% oxygen6.

Technology Advantages

There are several advantages to oxy-combustion besides the fact that it does not require

additional CO2 separation equipment. One is the ability to retrofit to existing plants. The oxygen-

flue gas mixture proportions can be optimized to produce similar in-furnace gas temperature

profile as the normal air-fired combustion7. This allows the same furnace to be used.

Oxy-combustion also helps decrease the emissions of other pollutants. A 60-70 percent

reduction in NOx emissions compared to air-fired combustion can be achieved. This occursprimarily from flue gas recycling, but also from the reduced thermal NOx levels due to lower

available nitrogen8. Furthermore, SOx emissions are decreased somewhat due to the condensation

of sulphates in the ducts and the absorption of sulphur in the ash. Unburnt carbon in fly ash is

reduced, resulting in improvements in boiler efficiency.9 The increased mercury oxidation

facilitates removal. 10

Cost is also important to consider. Oxy-fuel combustion has been shown to be less

expensive than post-combustion as a retrofit option. For post combustion scrubbing, CO2 capture

costs are 40.4-55 US$/tonne CO2. This results in an increase in the levelized cost of electricity

5Buhre, 2856Buhre, 2897Buhre, 289

8SECA 199Buhre, 29110SECA,19


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(LCOE) of 2.9-3.3 US cents/kWh. For oxy-fuel combustion costs are 30-35 US$/tonne CO 2 resulting

in an LCOE increase of 2.2-2.4 US cents/kWh.1112

Technology Limitations

Oxy-combustion faces several challenges. The potential for corrosion of the furnace and CO2

transportation systems due to the high SO2 concentrations in the flue gas may require

desulphurization of the recycled flue gas for oxy-fuel combustion13 Even though no work is

required for CO2 separation, the energy required for oxygen production and CO 2

compression/liquefaction reduces the overall thermal efficiency by approximately 9%14. This is

hampered by the fact that the oxy-combustion requires significant quantities of oxygen. In fact, oxy

combustion plants consumes roughly three times more oxygen per kilowatt-hour (kWh) than

compared to IGCC plants.15The oxygen requirement for an industrial sized 300 MWe is more than

5000 tons/day16. Moreover, the ASU incurs a large capital cost. Half of the capital cost in retro-

fitting a coal plant goes to purchasing the air separation unit.17 This is demonstrated inFigure 7

below.

Figure 7- Investment Breakdown for Oxy-Combustion Retrofit18

11Buhre, 299

12Katzer et. al, 30

13Buhre, 29414Buhre, 301

15NETL(2006), 1316Hong, 21

17IEA, 5818IEA, 59


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Pre-Combustion CO2capture

Technology Summary

In pre-combustion capture, the coal is partially oxidized in a gasifier with either air or pure

oxygen to produce syngas (a mixture of hydrogen and carbon monoxide). The syngas is thenprocessed in a water-gas-shift (WGS) reactor, which converts the CO to CO2 and increasing the

molar concentrations of CO2and H2 to 40% and 55%, respectively19. Due to the exothermic nature

of the gas-shift reaction two shift reactors are required to fully convert the CO to CO 2. It is at this

point (before any combustion has occurred) that the CO2is separated. Because the CO2 is at a much

higher concentration than in post-combustion flue gas, CO2capture is less expensive. A variety of

gas cleanup technologies can be utilized. Some proposed systems simultaneously remove CO2and

hydrogen sulfide20 before combustion therefore reducing the need for a desulphurization unit.

Others remove the acid gas components in separate steps. After condensation of water vapor, the

nearly pure hydrogen stream can be utilized in a gas turbine with a combined cycle for increased

efficiency. Lastly, the CO2 must be compressed to 2000-2215 psi (136-150 bar) for transport via

pipeline.

Technology Advantages

IGCC has gas cleanup benefits. In fact this is one of the reasons IGCC plants were designed

for in the first place. In addition to being able to achieve NOX, SOX emissions more than an order of

magnitude lower than U.S permit levels21, a sulfur recovery process can be integrated producing

valuable by-products. Moreover ,a >95% mercury removal is possible with minimal cost increase.

By cleaning syngas as opposed to flue gas (which has low density), the gas cleanup equipment is

much smaller. Oxygen driven gasifiers are utilized to avoid the unnecessary introduction ofnitrogen. As a result oxy-gasifiers have high compactness and efficiency and the amount of NOX

generated during oxy-gasification is very low. Hydrogen in the syngas reduces fuel NOX to

elemental nitrogen22.

The high thermal efficiency of IGCC plants makes them an ideal candidate for carbon

capture technology. This is because increased thermal efficiency results in lower fuel and CO 2

scrubbing/compression costs. Excluding the energy penalty from carbon capture, thermal

efficiencies as high as 53% with typical values around 38%.23 Even including the efficiency penalty

for CO2capture, IGCC plants are still more or equally efficient than conventional coal fired power

plants while producing few to no carbon emissions.

19SECA, 1820Known as acid gas components21Katzer et. al, 37

22Tzimas, 399623Guillermo Ordorica-Garcia et al, 2255


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The economics of IGCC over its lifetime are favorable. The IGCC capture option provides for

the lowest LCOE of any clean coal technology. It is estimated that the capture costs are between

19.3-24 $/tonne resulting in a LCOE increase of 1.75 cents/kWhr.

Technology Disadvantages

The primary limitation of pre-combustion carbon capture is that it can't be retrofitted to

existing plants and therefore only applies to newly built plants. However, the IGCC plants built

without carbon capture capability can be somewhat easily retrofitted in the future when regulation

(i.e carbon taxes) begin to take effect.

Though the oxygen requirement is not as high as for oxy-combustion plants it is still

significant. Roughly 10% of electricity production goes towards air separation24. A typical 250MW

IGCC requires 2000 t/day of oxygen. Current cryogenic separation requires 210 kWh/t oxygen. 8%

of electricity production goes towards CO2pressurization.

Currently IGCC plants have high initial capital costs and availability issues. IGCC plants with

carbon capture have an installed capital cost between 1890-2900 $/kWe compared to 600 $/kWefor NGCC plants.25 In the case where high capacity factor is a target, a spare gasifier is needed

which increases investment costs further (150-200$/kWe installed). 26 Lignite fired IGCC plants

cost $400/kW (installed) more than a hard-coal fired IGCC. Because IGCC utilizes emerging

technologies, these investment costs are expected to go down in the future.

24IEA,5325Guillermo Ordorica-Garcia et al, 225626IEA, 54


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Proposed Integration

Given that the several patents already show the benefits of nitrogen injected into the

combustor, we've focused on a novel use for the high pressure nitrogen. In researching IGCC with

carbon capture it is apparent that large amounts of energy are required for compression of CO2up

to 150 bar for insertion into pipeline. Realizing that pumping a liquid to high pressure requires lessenergy than pumping a gas to high pressure our design adds a liquefaction component to the IGCC

plant. The integration comes from the fact that the cooling load for liquefaction is provided by

turbo-expanding and throttling a certain fraction of the high pressure, low temperature nitrogen.

The following section will address our modeling/simulation tools and describe the schematics for a

base case and integrated case. For each we'll describe the process schematics while justifying and

explaining simplifying assumptions when they are made. To organize our discussion, the plant will

be divided into three main sections. Those sections are the gasification and power cycle, the air

separation unit, and the CO2 compression unit.

Plant Design and Model

Modeling and Simulation Tools

We decided to make an Excel based model to calculate data and connect input/ouput data

between components . This required input thermo-physical property data which was acquired from

NIST31. Because excel was used most calculations required manual lookup of property data,

however the Excel framework allowed different values and scenarios to be tested with some ease.

Chemical equilibrium data was used from in-class lectures.32

Gasification and Power Cycle

Because our proposed integration only affects the CO2compression and air separation ofthe plant, the gasification and power cycle analysis remains the same for both the base case and

integrated case. Our schematic is shown below:

31http://webbook.nist.gov/chemistry/fluid/

32Ghoniem [2], Slide 10


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Gasifier

T=1200 C

P=40 bar

HTSR

T=600 C

P=40 bar

QHTSR

HRSG

QHRSG

LTSR

T=200 C

P=40 bar

QLTSR

CO2SeparationT=200 C

P=40 bar

QCO2 sep.WCO2 sep

1 mol H2O

1 mol CH0.8

0.5 mol H2O

0.5 mol H2O

Condenser

QCond

.931 mol O2(From ASU)

Gas Turbine

(Combined Cycle)

eff= 0.55

WccQreject

0.537 mol H2

1.86 mol H2O


Schematic

1 2

4

3

1 mol CO2

56

(To steam cycle) (To steam cycle)

(To steam cycle)

Figure 9 - Gasification and Power Cycle Schematic. Note: Molar quantities shown per mole of coalinput. Note that the temperature, pressures, and molar concentration of states 1-6 are reported inFigure 21 of Appendix A.

Gasification

Pressurized oxygen (from ASU), coal, and water are fed into the pressurized entrained flow

gasifier with an output temperature of 1200 C and operating pressure of 40 bar. For the

gasification reaction we assumed the products to be at chemical equilibrium at the designated

output temperature. Therefore kinetics are ignored and the reactions approach their steady state

equilibrium value. For entrained flow gasifiers, this assumption is justified because of their high

operating temperatures between 10501400C which result in fast kinetics. From studying of a

commercial Shell coal gasifier, the equilibrium predictions were within 0.7% absolute of the

measured values, and equilibrium temperatures were very close to the gasifier exit temperatures33.

We also assumed that the gasifier is completely adiabatic and that the carbon is completely

converted to CO. The composition of the output stream (state 1) and the amount of oxygen required

was determined solved for using the chemical equilibrium equation for water gas shift reaction

(CO+H2O CO2+H2) in addition to mass and energy balance equations.

Water Gas Shift Reactor and Fuel Quenching

33Prins, 1252


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Next, the high temperature shift reactor (HTSR) converts more CO to CO2 with the addition

of water. The output temperature of the reactor was set as 600C to push the reaction forward. As

Figure 20 in the Appendix A demonstrates, this reaction only goes forward significantly for

temperatures below lower temperatures. This is a result of the exothermic nature of the water gas

shift reaction As expected, large quantities of heat must be removed from this HTSR and given the

high quality of this energy, some of it is recovered by inputting it to the steam cycle in the combined

cycle. More heat is recovered when the fuel stream temperature is lowered to 200C via a heat

recovery steam generator (HRSG).

After this quenching (state 3) the fuel syngas is sent into the low temperature shift reactor

where the remaining CO is converted to CO2. Again, since the shift reaction is exothermic, more low

grade heat (200C) is recovered and sent to the steam bottoming cycle. For the shift reactors it was

also assumed that kinetics were not an issue and that products were in chemical equilibrium.

In all cases where heat is recovered, the efficiency of conversion of this heat to energy in

the steam cycle was estimated using an adjusted Carnot efficiency discussed in class:

*car=1- ln(Th/To)/(Th/To-1) [1]

Where Th is the temperature of the source heat and To is the environmental temperature

assumed to be 300K.

CO2Separation, H2O Condensation

At the output of the LTSR, virtually all of the input carbon is in the form of CO2. Therefore

no CO remains, and at this point the CO2 is separated. We estimated the work required for CO2

separation by calculating the ideal separation work34and adjusting this by an expected efficiency.

Wsep,ideal= nmoles *R*Tsep*[ln(XCO2 /(1- XCO2) + ln(XCO2)] [2]

Where nmoles is the number of moles of syngas being purified, R is the ideal gas constant in

J/mol*K, and XCO2 is the concentration of CO2in the syngas.

Typical efficiencies for separation processes are between 15-30% so we assumed a 25%

second law efficiency. Additionally, it was assumed that the separation process was isothermal and

isobaric so that work put into the stream would also be rejected as heat (first law). After the CO 2

has been separated (state 5), the water vapor must be condensed via heat exchange with a cooling

tower. Finally at this point the hydrogen can be sent into the combined cycle.

Combined Cycle

We based the power cycle on an advanced GE 7H combined cycle configuration35. According to GE a

net plant efficiency of 60% is possible. Currently no H-class turbine is designed to run on pure

34Ghoniem [3], Slide 18


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hydrogen36. Therefore, we are making the assumption that H-class turbines will- in the near future-

be ready for pure hydrogen operation. However, due to the fact that the fuel stream is already at

elevated pressure (40 bar) the efficiency of the combined cycle would actually be higher. Therefore

60% efficiency is a reasonable assumption.

It should be clarified that the significant amount of heat recovered during the water gas shift

components and gas quenching is assumed to be sent to the a separate steam cycle from the

combined cycle.

Cryogenic Air Separation Unit Design

The integration between IGCC plants and cryogenic air separation is not uncommon.

However, as mentioned before, the only integration that occurs is the integration of the oxygen

output from the ASU. This neglects the potential use of the byproduct, whose composition is

dominated by nitrogen. In considering the integration of the nitrogen byproduct, we consider two

cases: full and partial integration of the nitrogen byproduct in the air separation process.

For the base case of full integration, all the nitrogen byproduct is used to intercool a hot,compressed input air stream. The schematic of the proposed plant is shown below:

35Existing IGCC plants run on E or F class turbines with net plant efficiencies of 49% and 57% respectively.

From a cost perspective, these small differences in power cycle efficiency are significant especially for IGCC

plants where all the components are sized based on the required flow rate of syngas.

36GE, 15


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Figure 10. Proposed air separation plant with full nitrogen use. The nitrogen

byproduct is expanded and used to in a heat exchanger to reduce the refrigeration work

needed to cool the compressed hot air stream.

The second case examines the possibility of integrating part of the nitrogen byproduct into

the coal gasification in the CO2 liquefaction process and use the rest as before, in a heat exchanger to

cool the incoming hot, compressed air stream. This process is illustrated by Figure (11) below:

Figure 11. Alternative case, with nitrogen flow split after State 11. The nitrogen gas is

split into two streams after it is expanded and integrated with the CO2 compression

process in the IGCC plant.

Description of the Integrated Cryogenic ASU Plant Process

The processes described by Figures (6) and (7) differ only slightly, and thus we will analyze

them them as one. The incoming air stream, composed of oxygen and nitrogen gases, comes in at

atmospheric pressure and temperature (T1=300K and P1=1 bar). The air mixture is then

compressed isothermally until it reaches a pressure of P2=20 bar at State 2. The compressed air

mixture is then cooled in counter-current heat exchanger with a pinch point T pinch=17K. The thencooled and pressurized air stream is further cooled by a refrigeration process, until it reaches a

critical temperature for the liquefaction of oxygen (T4=132.74K, P4=20 bar).

The critically-cooled air then enters a pressure column to be separated into its components.

Liquid oxygen forms at the bottom of the pressure column while gaseous nitrogen forms at the top.

The oxygen product and the nitrogen product are then extracted from the pressure column. The

oxygen is sent directly to the counter-current heat exchanger, while the nitrogen is expanded in an


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turbine, in which it reaches State 8 (T8=98K, P8=7 bar), and expanded in a Joule-Thomson valve,

reaching State 11 (T11=89.6K P11=1 bar).

At this point, the nitrogen stream can be split to go to the heat exchanger and the CO 2

liquefaction process or be fully utilized in the heat exchanger to refrigerate the incoming air. After

passing through the heat exchanger, the nitrogen stream exits at the environmental state

(T10=273K, P10=1 bar).

The oxygen stream is also directed towards the counter-current heat exchanger, where it

enters at the State 4 temperature and pressure. In the heat exchanger, the oxygen exits at State 7

(T7=273K, P7=20 bar). To integrate the oxygen stream into the IGCC plant, it must first be

isothermally compressed to its final state in the cryogenic ASU (T9=273K, P9=40 bar).

Assumptions for the Cryogenic ASU Analysis

In order to properly analyze this system without a chemical process simulation program,

several simplifying assumptions were necessary. Amongst the most important assumptions made

in the analysis of the ASU plant is that the air separation occurs in only one pressure column, whilereal processes have several stages to produce an oxygen product with high purity. The pressure

column as designed by Linde is presented below:

Figure 12: Lindes third, and succesful, attempt for an air separation column.The incoming compressed air is cooled by the liquid oxygen-rich mix at the bottomof the column. It is then pumped up the pressure column. There, through heattransfer interactions, the compressed air begins liquefy into an oxygen-richmixture. The oxygen rich mixture (~40% O2) is raised further in the separationcolumn, where oxygen condenses and descends into a catch tray where is extractedat high purity.


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In our analysis, we assumed a compressor efficiency for the ASU plant of 0.75. This is a

conservative estimate, yet it is still within the ranges available for current compressor technology.

However, current compressor technologies have higher efficiencies, so our proposed design could

be improved further by assuming higher compressor efficiencies.

Additionally, the COP of the refrigeration cycle was assumed to have a value of COP=3,

which is an average value for the available refrigeration technology.

CO2compression unit

We assumed that CO2 would be transported via pipeline and so research into pipeline

specifications was required. Depending on the source, these specifications vary from 136 bar37all

the way to 18538bar so we assumed a pressure of 150 bar. Another requirement is that the CO 2

must be injected in at temperatures below 49C39.

Base Case:

To model the compression in the base case we assumed a single stage isothermalcompression from 40 bar to 150 bar. In reality compression would be staged (four or more stages)

with intercooling in order to approximate isothermal compression. The schematic below shows the

compression unit for the base case.

Isothermal CO2 Compression

Schematic (Non-integrated)

1 mol CO2(From CO2 Separation)

T1=300K

P1=40 Bar

Isothermal

Compressor

T2=300 K

P2=150 Bar

(To Pipeline)

eff= 0.8

WCO2 comp

QCO2 comp

Figure 13- CO2 Compression Schematic for base case.

The isothermal compressor has an 80% efficiency. Although the CO2 was originally 200C at

the output of the separation component we have assumed it would be cooled to 27C (300K) via

cooling tower.

Integrated Case:

37Santos, 23

38In Salah Gas

39Santos, 23


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In order to reduce the compression energy, we designed a system to refrigerate the CO2to a

saturated liquid prior to pumping up 150 bar. The high pressure, low temperature nitrogen from

the ASU is turboexpanded and throttled40to 1 bar resulting in a temperature of 89.6K.

Cryogenic CO2

Compression

Schematic (Integrated)

1 mol CO2(From CO2 Separation)

T1=300K

P1=40 Bar Counter-Current HX Cryogenic Pump

CO2 (Saturated Liquid)Tsat=278.45K

P2=40 Bar

1.867 mol N2(From ASU)

T4=89.6 K

P4=1 Bar

T5=290K

P5=1 Bar

T3=294.34 K

P3=150 Bar

(To Pipeline)

eff= 0.8

(Vented to atmosphere)

WCO2 cryo

Figure 14- Cryogenic CO2 Compression. Note: Molar quantities are per mole of coal input

At the moderate pressure (40 bar) that the CO2 exits the separation, the saturation

temperature is 278.45K. Despite this high saturation temperature, significant cooling is required by

the nitrogen in order to absorb the latent energy of condensation from the CO 2. To take fulladvantage of nitrogen refrigeration, it we assumed a counter-current heat exchanger which brings

the nitrogen to within 10 degrees41of the CO2 inlet temperature. Using a control volume around the

HX, we solved for the number of moles of nitrogen required to completely condense the CO2.

Through this analysis only 1.87 (mole N2)/(mol Coal) were required. This is less than the 3.5 (mole

N2)/(mol Coal)42 created by the air separation unit. As a result, the remaining 1.637 (moles

N2)/(mol Coal) are sent back to the ASU to pre-cool the inlet stream. In the base case, allN2is sent

40Throttling is required because turboexpansion alone cools the N2 so much that it liquefies - a situation not

handled well by the expansion turbine. As a result, turbo expansion occurs from 20 down to 7 bar while the

remaining pressure drop to 1 bar occurs through the throttle. The result is low temperature nitrogen vapor at

89.6 K and 1 bar.

41Also known as pinch point or approach.

42 As shown in the gasification schematic .931 mol O2 are required for 1 mol of coal. Therefore .931*3.76

moles of nitrogen are produced in the ASU.


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back to the ASU. The amount of N2 determines the amount of auxiliary cooling refrigeration

required. The pump is assumed to be adiabatic with an isentropic efficiency of 80%.

Results and Discussion


In order to compare the base case versus the integrated case, we assume that an equal mass

flow rate of coal is input into both plants and compare the work production and requirement for

each component. Since the gasification and power cycles for both the base case and integrated case

are equal, the power producing components will provide equivalent amounts of power. The

difference in the two will come from analyzing the difference in parasitic components (O2

production and CO2compression.)

One important gasification performance parameter is the cold gas efficiency of the

gasification, shift, and quenching. Since 0.537 (molH2/molCH0.8) are produced, the cold gas

efficiency can be calculated by:

cold gas = (nH2LHVH2)/(nCH0.8LHVCH0.8) [3]

cold gas = 0. 402 [3.1]

It is relatively low because it does not include the heat recovered from the gas stream. If this

recovered work is included in the numerator, the cold gas efficiency increases to 76.7%. The figure

below shows the work contribution for all power producing components.

Figure 13- Power output for each component in megawatts. Note that total work output

is 400MWe


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In both cases the coal mass flow rate required to attain a total work output43of 400 MWe

was 2289 metric tons of coal per day and 5327 tons of O2 per day. Of note is how significant the

contribution of the heat recovery steam generator (179 MWe) is to total work output. The

combined cycle contributes 153.5 MWe or 38% of total work output.

CO2Separation

CO2 separation requires a significant portion (~17%) of total plant power

(67MWe/400MWe). This is very sensitive to the assumed second law efficiency of separation.

CO2Compression

As expected the power required for CO2compression is much lower (13 times lower). The

integrated case (which refrigerates the CO2) only requires 1.7 MW for compression while the base

case needs 28.3 MW. Both are relatively low because the high initial pressure of the CO2(40 bar).

Air Separation

The net work requirement for the base and integrated cases were calculated and comparedto the industry average. The results from these calculations are tabulated in Figure 14 below:

Figure 14: Comparison of the energy requirement of the base and integratedcase against the industry average.

The increased net work requirement for our proposed systems stems from the initial high

compression of the air flow into the ASU. The average compression ratio for industry cryogenic

processes is about 6, while our proposed method compresses the incoming air flow to 20 bar.

43Total work output is defined as the sum power output of all power producing components. Net work will be

discussed later.


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Additional work is required for compressing the oxygen to delivery pressure for integration into

the coal gasification process, but it is relatively smaller when compared to the large energy

requirement for compressing the air feed to 20 bar.

Second Law Efficiency Analysis for the Cryogenic ASU plant

The minimum separation work can be determined using the inlet and output conditions for

the cryogenic ASU plant. The minimum work was calculated using an availability analysis,

Assuming the environmental conditions to be Tenv=300K and Penv=1 bar, the minimum separation

work for our cryogenic plant design was Wmin=14.925 kJ/mol air, which translates to an energy

requirement of 129.55 kWh/ton O2. The Second Law efficiency was then calculated using the

minimum separation work and the net work requirement in each case.

Using the value for the minimum separation work, we calculated the following Second Law

Efficiencies for the base and integrated cases:

Figure 15. Comparison of Second Law efficiencies for the base case and theintegrated case.

We were expecting that the integrated case has a lower efficiency since a fraction of the

nitrogen is sent to refrigerate the CO2 before compression. Thus, additional refrigeration is required

to compensate the loss of cooling due to the absence of the full quantity of nitrogen in the base case.


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Overall Plant Performance

The following chart summarizes power requirements of each component for both

integrated and base case.

Figure 16- Power requirement (MWe) for each component. (Base Case vs. Integrated Case)

Since the integrated case loses a certain amount of refrigeration to the CO 2 compression

component, we would expect that the auxiliary refrigeration requirement be higher. The base case

requires almost no auxiliary refrigeration (0.8 MWe) whereas the integrated case needs (7.6 MW).

Overall, the major power requirement for the air separation components is the initial air

compression from 1-20 bar and requires 92.8 MWe out of a total plant output of 400 MWe. The

additional oxygen pressurization from 20-40 bar requires 4.8MWe. In both the integrated and base

case the same quantity of oxygen is required so this energy requirement is the same in both cases.

Overall Plant Performance

Overall the integrated case produces higher net work at 226 MWe compared to the base

case of 206.2 MWe. The first law efficiencies44for each are 34.1% and 31.11% for integrated and

base case, respectively. This is summarized in the figure below:

44Input thermal energy is taken as the coal molar flow rate multiplied by its LHV of 320,000 J/mol.


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Figure 18- Plant efficiency vs. Auxillary Refrigeration COP- (Dotted Line Indicated Assumed

Value)

As the COP of refrigeration goes down, the refrigeration requirement increases rapidly for

the integrated case. Realistically, the COP will not be below 1 so this sensitivity analysis

demonstrates that regardless of what COP we assume, the integrated case will always be more

efficient.

One parameter which was found to strongly affect the performance of the base case was the

assumption of isothermal compression. For compression of an ideal gas, isothermal compression

requires less work than adiabatic compression. Interestingly, it was found that for compression of

CO2 in the pressure range we are dealing with (40-150 bar) the opposite is true. Adiabatic

compression requires far less work than isothermal compression. Instead of requiring 28.3 MWe in

the base case, only 7.7 MWe are required.


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Figure 19- Plant efficiency vs. CO2 compression assumption. Note: 80% compression efficiencywas assumed in all cases

As a result, plant efficiency for the adiabatic case is 34.24% which is slightly higher than

integrated case. More research should go into determining whether commercial adiabatic CO2

compressors with a compression ratio of (150bar/40bar) = 3.75 are available, as this analysis

shows significant energy savings. For adiabatic compression, the CO2 output stream was calculated

to exit at 423K (150C); therefore the CO2would have to be cooled prior to pipeline injection.

Since neither perfect adiabatic compression nor perfect isothermal compression is achievable in

practice, we can safely assume that the efficiency of the base case will be at some point between

31.13% and 34.24%.

Conclusion:

Base case and integrated models of an IGCC plant with carbon capture and cryogenic air

separation were developed in Excel based on thermodynamic and chemical equilibrium

calculations. It was found that integration increases net work and efficiency by drastically reducing

the CO2 compression work required. Adiabatic CO2 compression also resulted in an equally

significant energy reduction compared to isothermal CO2 compression. Initial air compression into

the air separation unit from 1- 20 bar was the single largest energy requirement (93 MWe) with

CO2 separation work being the next largest (67 MWe). Second law efficiency for air separation unit

ranged from 30.56% in the base case to 286. in the integrated case.


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Appendix A

Figure 20 - Equilibrium Values versus Temperature for water gas shift reaction (CO + H2O

CO2+H2)


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Gasification States

State T P X_Co2 X_CO X_H2O X_H2

1 1200 40 0.685 0.315 1.178 0.222

2 600 40 0.933 0.067 1.430 0.470

3 200 40 0.933 0.067 1.430 0.470

4 200 40 0.999 0.001 1.864 0.536

5 200 40 0.000 0.000 1.864 0.536

6 200 40 0.000 0.000 0.000 0.536

Figure 21- Gasification and Power Cycle Concentrations, Temperatures, and Pressures

Integration of Cryogenic Air Separation and IGCC with Carbon Capture

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