CO2 capture by partial oxidation of natural gas · Air and steam are used to produce a synthesis gas from natural gas by a combination of partial oxidation and steam reforming. The

CO2 CAPTURE VIA PARTIAL OXIDATION OF NATURAL

GAS

Report Number PH3/21 April 2000

This document has been prepared for the Executive Committee of the Programme. It is not a publication of the Operating Agent, International Energy Agency or its Secretariat.

Title: CO2 capture via partial oxidation of natural gas Reference number: PH3/21 Date issued: April 2000 Other remarks:

Background Previous work by IEA GHG established pre-combustion decarbonisation (PCD) as a potential method of reducing CO2 emissions from natural gas–fired power stations. The PCD approach to emission reduction is now accepted as an alternative to CO2 capture by scrubbing the power station flue gas (post-combustion decarbonisation). A recent study by IEA GHG of the leading options for capture of CO2 in power generation (report PH3/14) concluded that, for natural gas feeds, that there was little to chose between pre- and post- combustion capture options.1 The objective of this study was to assess processing options that might improve the attractiveness of PCD as a method of reducing CO2 emissions. In the original study on this technology (report PH2/19) a process based on catalytic partial oxidation (CAPO), using air abstracted from the gas turbine’s compressor, appeared the most promising of the options examined. Although, as far as we are aware, a commercial-scale CAPO unit does not exist there appears to be no major technical reasons why such a unit could not be operated.2 In this study the CAPO process is taken as a reference case and potential improvements to the process assessed. The opportunity was taken to re-examine the relative merits of air- and oxygen-blown synthesis gas units.

Approach adopted The study was organised in stages as follows:

• a reference design was produced to serve as a marker against which potential improvements could be compared. This design was based on the process developed previously by IEA GHG and Statoil as reported in PH2/19.

• a number of potential process improvements were identified and the engineering judgement of the contractor was used to assess their potential impact. The conclusions of this screening analysis are reported.

• Five process variations were developed incorporating potential process improvements and two of these ‘improved’ processes were examined in more detail.

The study was done by Fluor Daniel Inc. of California, USA.

1 The option of post-combustion scrubbing the flue gas using an amine-based solvent appeared marginally the cheapest for the present state-of-the-art. 2 Existing commercial applications of synthesis gas technology are based on steam reforming and/or oxygen-blown partial oxidation. Air-blown reactors are used commercially as secondary reformers in ammonia plant. If air is used, the synthesis gas will contain large quantities of nitrogen which, in most applications, is undesirable. In a PCD application, the presence of nitrogen is not a major disadvantage as it is needed to facilitate combustion of the hydrogen-rich gas in the gas turbine (see later in this report and the forthcoming report PH3/12). It is interesting to note that the developers of the Syntroleumtm version of the Fischer-Tropsch process suggest that using an air-blown process to produce synthesis gas is an effective way to reduce costs.

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Results and discussion The objective of the study was to assess processing options that could improve the attractiveness of the pre-combustion decarbonisation process in natural gas-fired power plants. Figure S1 illustrates the main components of the reference process used as a basis to assess the effectiveness of potential modifications. Air and steam are used to produce a synthesis gas from natural gas by a combination of partial oxidation and steam reforming. The catalytic air-blown partial oxidation (CAPO) reactor is a single pressure vessel containing two reaction stages.3 The reactor pressure is dictated by the discharge pressure of the gas turbine’s air compressor. Shift conversion is used to convert the CO content of the synthesis gas to CO2 which is then captured in an amine-based scrubbing system. The hydrogen-rich fuel gas is fired in the gas turbine of a combined cycle to produce electricity. Figure S1: schematic of the reference CAPO process processing options The following is a summary of the processing options considered in the report: 1. Syngas reactor. A wide range of technologies are in common use for the production of synthesis gases. Syngas can be produced in a steam-methane reformer, this is the method by which most hydrogen is made commercially; air/oxygen is not required. The other main process option is partial oxidation in which oxygen is one of the reactants. There are many hybrids of these two basic reaction routes such as for example, autothermal reforming, one variant of which is used as the reference process in this study. The choice of technology for synthesis gas production is often a marginal decision4 but it is generally agreed that steam-methane reforming of natural gas is cheaper than oxygen-blown partial oxidation. Partial oxidation using oxygen is preferred for production of synthesis gases if the feedstock is heavier 3 The reactor consists of a refractory-lined vessel having an inlet conical combustion zone, in which combustion/partial oxidation reactions take place, and a lower cylindrical volume containing a nickel-based reforming catalyst, in which the reforming reaction mainly occurs and the reactants reach equilibrium. 4 A comment received from an expert, about the previous study on this topic, was that “alternative routes to the production of synthesis gas from methane can be even found on the same site”.

Shiftconversion

Syngas reactor

Steamcycle

air

natural gas

gas turbine

CO2 to storage

CO2capture

H2O, N 2 tostack

H2-richfuel gas

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than natural gas, or if a high CO:H2 ratio is required. If partial oxidation with air can be used, it appears to be less expensive than steam-methane reforming. The use of advanced synthesis gas technology, i.e. a two-stage reactor system involving a gas-heated reformer (GHR), is assessed in some detail.5 A gas-heated reformer gives a significant gain in the efficiency of synthesis gas generation- this is reflected in both the calorific value of the raw syngas and the hydrogen content of the fuel gas (see table S1). However, in the overall power generation process, much of this gain is lost by a need to add large amounts of steam to the gas turbine feed to minimise emissions of NOx (see later). The extra cost of the more complicated reactor system off-sets a small overall gain in process efficiency, with the result that the power generated is more expensive (see tables 3-1, 8-4, and fig 8-1in main report). The conclusion is that there is no advantage to be gained by use of complicated (advanced) reactor systems to raise the efficiency of synthesis gas production, if the resulting hydrogen-rich fuel has to be diluted before it can be used in the gas turbine (see later). Table S1: Synthesis and fuel gas compositions

Component (mole%)

Reference CAPO case ‘Improved’ process with gas -heated reformer

raw syngas fuel gas raw syngas fuel gas H2 27.5 50.7 33.4 59.6 N2 31.6 45.5 26.0 36.0 H2O 26.3 0.1 25.0 0.7 CO 8.0 0.3 10.2 0.8 CO2 5.8 2.1 4.8 0.5 CH4 + 0.5 0.7 0.4 2.0 Ar 0.4 0.6 0.3 0.4

Total 100.1 100.0 100.1 100.0 calorific value

(MJ/kg) (LHV) 6.5 8.6 8.6 13.4

2. Air vs. Oxygen. The use of oxygen (i.e. absence of nitrogen) limits the sensible heat carried out of the reactor by the syngas product. However, steam has to be added to an oxygen-blown partial oxidation unit to moderate the temperature. In an air-blown CAPO unit the nitrogen acts as a temperature moderator. At lower temperatures, less sensible heat is contained in the syngas and it is more likely that an air-blown configuration will show advantages. This study endorsed the conclusion of previous work (PH2/19) that the use of an oxygen-blown partial oxidation unit is less attractive than the reference CAPO case. The large power consumption of an ASU 6 and the relatively low efficiency of the partial oxidation unit, which operates at a higher temperature than the CAPO unit, lower the overall plant efficiency. In this study, Fluor Daniel calculate an efficiency of 40.7% for an oxygen-blown version of the process and 42.8% for the reference air-blown CAPO plant. 3. Shift conversion. The shift conversion reaction is exothermic. A two-stage shift system, as in the reference design, enables a high conversion to be achieved and most of the heat to be recovered at a high temperature. The efficiency of the plant is increased if this heat is used to pre-heat the fuel gas before it enters the gas turbine. A detailed design is needed to optimise capital expenditure on heat recovery against the reduced cost of power obtained through the increased efficiency. 5 There are various proprietary versions of this technology the development of which is largely based on ammonia plant technology introduced by ICI in the late 1980s. It usually involves the hot synthesis gas produced in one reactor being used to provide heat for reforming reactions in another reactor. (UHDE have developed a version of this process where both stages take place in a single vessel). 6 Air Separation Unit

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4. Combustion of hydrogen-rich fuel gas. In this study, acting on the advice of GE, large amounts of steam are added to the fuel gas as a method of controlling NOx emissions (see table S2). This represents a significant loss in potential efficiency, particularly in the case of a gas-heated reformer which efficiently produces a syngas rich in hydrogen. As discussed in the main report, if turbine considerations allow, it is likely to be cost-effective to use a selective catalytic reduction (SCR) unit on the turbine’s exhaust gases for NOx control.7 This is a key area for further development and is discussed in detail in a forthcoming report on gas turbines (PH3/12). Table S2: Steam added to reduce NOx production in the gas turbine

reference CAPO case gas -heated reformer Process efficiency (LHV) 42.8 44.1 Hydrogen content of fuel gas (mol%) 51 60 Calorific value LHV (MJ/kg) 8.6 13.4 Steam diluent required (kg steam /MJ fuel gas)

0.04 0.057

kg steam added / kg of fuel gas 0.34 0.76 5. Fuel gas temperature. Significant performance improvements can be obtained by supplying the fuel gas to the turbine at elevated temperatures. A supply temperature of 399°C was selected arbitrarily on the basis of it being a GE standard. GE offers a fuel system for temperatures up to 538°C but it tends to be expensive and the implications for a hydrogen-rich fuel gas are not known. 6. Operating pressure. In the reference case, the operating pressure of about 30 bar is determined by the discharge of the gas turbine’s air compressor. On the advice of ICI Synetex, the operating pressure considered in this study was limited to a maximum of about 50 bar.8 Fluor Daniel recommend careful consideration of the potential for increasing the pressure, particularly in the case of a CAPO reactor. They point out, for instance, that increasing the pressure, decreases the size of the process equipment but also increases ‘methane slip’ i.e. the concentration of methane in the syngas. In the study report, a high pressure CAPO case (50 bar) is compared with an ‘improved’ medium pressure (30 bar) CAPO process. The high pressure case shows no performance advantage as a gain in output from an expansion turbine is offset by an increase in the power requirement for the air compressor; the process efficiency in both cases is 43.8%. The cost implications are not clear. Fluor’s experience with integrated gasification combined cycles is that, when the pressure is raised over 30 bar, the increase in cost of equipment due to increased wall thickness is not quite offset by the decrease in size. However, costs could be lower, if an increase in pressure means that the number of process trains could be reduced. The overall conclusion is that increases in operating pressure could be cost-effective but are not likely to give a radical improvement in cost and performance figures. (Note that there has been a progressive history of increased operating pressure in the development of chemical processes based on synthesis gas; modern ammonia plants operate with reformer pressures in the range of 30 to 45 bar.) 7. CO2 capture. CO2 capture using a physical solvent is not likely to be cost-effective unless the operating pressure is increased significantly. This is because the presence of N2 in the syngas reduces the partial pressure of CO2. This study endorsed the previous view that the best commercially 7 SCR is relatively inexpensive, typically the specific capital cost is in the region of 25$/kWe. A disadvantage is that the system requires the handling of ammonia and introduces the possibility of ammonia emissions (see main report). 8 This is because of concerns about metal dusting (see later). A main cause of carbon formation is the Boudouard reaction 2CO � C + CO2. Pressure drives the reaction equilibrium to the right.

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available solvent would be based on methyl di-ethanolamine (MDEA) which is generally described as a physio-chemical solvent (physical and chemical absorption take place).9 8. methane slip and/or by-pass. About 8% of the natural gas feed can by-pass the synthesis gas and CO2 capture sections of the plant and be fed to the gas turbine whilst the maintaining the overall reduction in CO2 emissions at 85%. This would improve the combustion characteristics of the fuel gas and increase the efficiency of the process by about a percentage point at little or no capital cost.10 This raises a general query about the optimum level of CO2 capture; it is obvious that the cost of power generation can be significantly reduced if more natural gas by-passes the capture process but at a penalty of reducing the level of CO2 capture. The optimum level of CO2 emission reduction can be queried for any CO2 capture process; this has been raised before by IEA GHG but not investigated. Other potential improvements to the process discussed in the report include: natural gas conditioning and pressurisation, refrigeration to aid optimal heat recovery, fuel gas humidification, and modifications aimed at optimum combination of fuel gas and gas turbine. It is recommended that consideration be given to processing schemes using at least a part of the gas turbine exhaust to supply combustion air to the reformer.11 comparison with previous performance results In table S3 process efficiencies calculated in this study are compared with previous work on air-blown PCD processes. In all cases the level of CO2 capture achieved is 85% or greater. Table S3: A comparison of calculated efficiencies for air-blown PCD processes

Study Contractor % efficiency (LHV)12 This report (PH3/21):

reference CAPO case Fluor Daniel 42.8 an ‘improved’ process with a GHR Fluor Daniel 44.1

Precombustion decarbonisation study (PH2/19): Foster Wheeler

49.6

Studies for others by Foster Wheeler13 Foster Wheeler

46.3 to 46.9

Apart from turbine assumptions, the differences in efficiency are probably largely due to assumptions about the amount of steam required in the process and the temperature levels at which it is used. 9 Note that a purely chemical solvent, monoethanolamine (MEA), is generally accepted as the most suitable established solvent for CO2 capture from post-combustion flue gases. 10 It is likely that provision of a natural gas supply and facilities for blending with the hydrogen-rich fuel gas will be needed for start up and shut down. Alternatively, the synthesis gas reactor could be designed for a reduced conversion of methane but this would lead to higher methane partial pressures in the subsequent processing steps making CO2 capture more difficult. 11 This was one of the process configurations considered in earlier work (PH2/19). It is possible that the concept could be successfully used to recycle heat to the upper part of the combined cycle but a detailed evaluation would be needed to assess the options. 12 The turbine combined cycle used in this study is based on GE’s MS9001FA. Depending on the application, GE quote efficiencies of 56.7% and 57.1% for this machine in a combined cycle. In the Foster Wheeler study for IEA GHG, the combined cycle was assumed to have an efficiency of 60% (based on new ‘H’ units, the first delivery of which is scheduled for this year). In the Foster Wheeler studies for others, lower efficiency gas turbine cycles were assumed but the details are not known. GE have reported that firing temperatures need to be reduced when burning hydrogen-rich fuels. This will lower the efficiency. 13 Private communication

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There are three main steam requirements, in each of which gains in efficiency can be traded against the cost of investment; a detailed design would be needed to optimise the overall integrated process: A) Steam to the synthesis gas reactors. The amount of steam which must be added is significantly

more than is theoretically required for process reasons, i.e. to convert CH4 to H2 by steam reforming.14 One major reason for the excess steam is to avoid ‘metal dusting’. Metal dusting occurs when metals exposed to synthesis gases having a high CO/CO2 content disintegrate into a fine metal powder and metal oxide particles mixed with elemental carbon. The phenomenon is not fully understood and has caused serious problems in ammonia and methanol plants. The solution is to avoid metal surfaces within the temperature range 400°C to 850°C and to increase the steam/carbon ratio to decrease the carbon concentration in the synthesis gas. Advice was taken from ICI on the appropriate steam level for the designs in this study but, as the design of this process is not established, particularly with the use of an air-blown partial oxidation unit, the amount of steam required is speculative.

B) Steam to the CO2 recovery reboiler. This is a large user of steam and there are many

permutations and sources of steam and temperature level. In this study Fluor Daniel used designs, that in their judgement, were a “decent compromise between expense and efficiency”.

C) Steam supply to the gas turbine to minimise NOx emissions. Large quantities of steam are used.

(See table S2). As discussed earlier, this can have a major effect on turbine efficiency. This topic is dealt with in detail in the forthcoming report on turbines (PH3/12). In this study, steam is extracted directly from the steam turbine. Foster Wheeler, in commenting on these results, have said that: “we integrated and utilised heat in a better way by humidification of the gas turbine fuel gas using low-grade heat”

The overall conclusion is that the efficiencies calculated in this study are probably conservative and that an efficiency in the region of 46% can be expected for a cost-effective process design using an F-series gas turbine.15 cost implications As discussed in an earlier report (PH3/14), much of the cost penalty for CO2 capture is attributable to capital charges. The specific cost of the complete plant derived in this study is compared with previous work in table S4. These cost estimates are remarkably consistent given the relatively low level of estimating effort involved in scoping studies and the number of engineering design judgements made. The reader is cautioned that these figures are based on similar methodologies for the outline estimation of costs and that the overall accuracy is not likely to be better than +/-30% Table S4: A comparison of specific costs for air-blown PCD processes

Study Contractor Specific cost (US$/kWe)

This report (PH3/21): reference CAPO case Fluor Daniel 1 070 an ‘improved’ process with a GHR Fluor Daniel 1110

Precombustion decarbonisation study (PH2/19) Foster 940 14 The steam-to-carbon ratio in conventional methane reforming aimed at maximising the production of hydrogen is around 3:1. The large amount of excess steam is primarily there to drive the conversion reaction in the direction of hydrogen production. 15 In support of this statement, note that a recent patent application (WO 99/41188) on a version of this process specifies “overall efficiency up to a level of at least 45%”.

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Wheeler Table S5 gives a breakdown of specific costs for the reference CAPO process and an ‘improved’ process with a GHR. It can be seen that, even with the reduced quantities of gas involved in pre-combustion decarbonisation (as opposed to post-combustion decarbonisation), the cost of CO2 capture and processing equipment is considerable. Note that in terms of making decisions on processing options the differences between the reference CAPO and the ‘improved case’ are very marginal. Table S5: Specific costs (US$/kWe)

Process area Specific cost (US$/kWe) reference CAPO ‘improved’ case + GHR Oxidant supply and fuel gas processing 72 67 Syngas generation 144 163 CO2 capture and regeneration 177 180 CO2 compression/liquefaction 70 71 Combined cycle power block 323 337 General facilities 282 286 TOTAL 1068 1105

Expert Group and other comments

One reviewer said that the air vs. oxygen debate is still far from clear. In particular he noted the development of membrane technologies for the production of enriched air could be usefully applied. Two reviewers noted that, as steam consumption in the gas turbine to reduce NOx production is a problem, the use of selective catalytic reduction technology on the turbine’s exhaust gas might be a cheaper and better solution. It was suggested that metal dusting problems in the reactor feed/effluent exchanger could be avoided altogether by generating HP steam instead of preheating feed. (A counter-argument to this is that steam generation is indicative of lost efficiency as it is energy that bypasses the gas turbine.)

Major conclusions

The precombustion decarbonisation route to a reduction of emissions from gas-fired power generation can be built up from various combinations of the component technologies. The differences between versions of the processes are not great. No major potential improvements on the air-blown CAPO process, as defined in previous studies, were found. This conclusion refers to near-term or state-of-the-art technology; in the longer term, developments aimed at improving synthesis gas technology such at ceramic membrane reformers could have a major impact. The efficiency of a cost-effective CAPO process is about 46% (LHV). This is comparable to the efficiency of 47%, reported in the leading options study (PH3/14), for post-combustion decarbonisation using amine scrubbing. Neither process is obviously superior to the other. Both processes have the potential for improvement. The optimum version of the PCD process depends on the requirements for the hydrogen-rich fuel to be burnt in the gas turbine. The reference air-blown CAPO process can produce a fuel gas relatively cheaply which can probably be burnt in existing gas turbines without a major loss of performance.

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A gas-heated reformer could increase the cost-effectiveness of synthesis gas production and the hydrogen concentration in the fuel gas but, with existing turbine technology, no advantage accrues in the overall process as the fuel gas has to be diluted by large amounts of steam to reduce NOx formation. Alternatively, use of selective catalytic reduction on the gas turbine’s exhaust may be a cost-effective way of reducing emissions of NOx from a turbine using hydrogen-rich fuel. Within the major constraint of keeping the capital cost as low as possible, the following are the key areas for potential performance improvement in the process: • Minimise steam production. Steam production is a direct indicator of efficiency lost. • Introduce as much sensible heat as possible into the fuel and air entering the gas turbine. These key areas are essentially ‘systems integration’ issues, the effective development of which requires a considerably more detailed design than is possible in scoping studies like this. As with other technologies involving the production of synthesis gas, their application should lead to progressive improvements. Any significant development of the process is likely to lead to competitive proprietary versions.

Recommendations IEA GHG should continue to promote this technology as a potential approach to reducing the emissions of CO2 from power generation. Further work on the development of this process must be at a more detailed (and expensive) level in order to integrate the ‘system’ design options in (something approaching) the optimum manner. Further work at the scoping level (i.e. as in conventional IEA GHG reports) is not recommended as it is unlikely to add significantly to the present level of knowledge. As a general issue, IEA GHG should consider work aimed at assessment of the optimum level of CO2 reduction in CO2 capture processes.

CRE Group, Ltd. FLUOR DANIEL, INCElectricity Production and CO2 Capture via Partial Oxidation of Natural Gas Contract 04592200

D:\Power_generation\A P O Study\REPORT.doc

Electricity Production and CO2 Capturevia Partial Oxidation of Natural Gas

Presented to

CRE Group, LTD.

by

`Fluor Daniel, Inc

One Fluor Daniel DriveAliso Viejo, CA 92698

Principal Investigators:

Ashok RaoDavid Francuz

Jeffrey ScherffiusEdward West

November 1999


D:\Power_generation\A P O Study\REPORT.doc

DISCLAIMER

The following estimate/report is based in part on information not within Fluor Daniel’scontrol. It is believed that the estimates and conclusions contained therein will bereliable under the conditions and subject to the qualifications set forth. However, FluorDaniel does not warrant or guarantee their accuracy. Use of such estimates/reportsshall, therefore, be at the user’s sole risk. Such use shall constitute a release andagreement to defend and indemnify Fluor Daniel from and against any liability (includingbut not limited to liability for special, indirect or consequential damages) in connectionwith such use, whether liability is asserted to arise in contract, negligence, strict liabilityor other theory of law.


D:\Power_generation\A P O Study\REPORT.doc i

TABLE OF CONTENTS

STUDY BACKGROUND ......................................................................................................... 1

EXECUTIVE SUMMARY ........................................................................................................ 2

1.0 INTRODUCTION ......................................................................................................... 7

1.1 Study Objectives.......................................................................................................... 7

1.2 Study Approach ........................................................................................................... 8

1.2.1 Task 1 - Prepare and assess a Base Case Design .............................................. 8

1.2.2 Task 2 - Consideration of Processing Options ..................................................... 8

2.0 PROCESS DESIGN BASIS ....................................................................................... 10

2.1 Site Conditions........................................................................................................... 10

2.2 Products..................................................................................................................... 10

2.2.1 Electric Power .................................................................................................... 10

2.2.2 Recovered CO2.................................................................................................. 10

2.3 Raw Materials ............................................................................................................ 11

2.3.1 Natural Gas ........................................................................................................ 11

2.3.2 Municipal Water.................................................................................................. 11

2.4 Plant Configuration .................................................................................................... 11

2.5 Environmental Criteria ............................................................................................... 11

3.0 SCREENING ANALYSES.......................................................................................... 12

3.1 Unit Process Considerations...................................................................................... 12

3.1.1 Fuel Gas Compression....................................................................................... 12

3.1.2 Single-stage Shift System .................................................................................. 13

3.1.3 Alternative Methods for Capture of Carbon Dioxide ........................................... 13

3.1.4 LiBr Refrigeration ............................................................................................... 14

3.1.5 Humidification..................................................................................................... 14

3.1.6 Gas Turbine........................................................................................................ 15

3.1.7 Alternative Reactors and Reactor Arrangements ............................................... 16

3.1.8 Replacing Air Feed to POX by Oxygen or Enriched-air ...................................... 17

3.1.9 Increased Operating Pressure............................................................................ 19

3.2 Configuration Development ....................................................................................... 20

3.2.1 Base Case (Case 1) ........................................................................................... 20

3.2.2 Case 2................................................................................................................ 20

3.2.3 Case 3................................................................................................................ 23

3.2.4 Case 4................................................................................................................ 25


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3.2.5 Case 5................................................................................................................ 27

3.3 Conclusions and Recommendations.......................................................................... 27

4.0 BASE CASE PROCESS DESIGN ............................................................................. 34

5.0 IMPROVED CATALYTIC AIR PARTIAL OXIDATION PROCESS DESIGN .............. 48

6.0 IMPROVED CATALYTIC AIR PARTIAL OXIDATION WITH GAS HEATED

REFORMER PROCESS DESIGN ........................................................................................ 52

7.0 COST ESTIMATION.................................................................................................. 57

7.1 Plant Cost Estimate ................................................................................................... 57

7.1.1 Estimating Criteria and Methodology.................................................................. 57

7.1.2 Engineering, Procurement, and Construction Schedule..................................... 66

7.2 Technical and Financial Assessment Criteria ............................................................ 66

8.0 DISCUSSION OF RESULTS ..................................................................................... 71

8.1 Screening Analyses ................................................................................................... 71

8.2 Performance Summary - Improved Cases................................................................. 72

8.3 Cost Estimates........................................................................................................... 77

8.4 Conclusions and Recommendations.......................................................................... 83

APPENDIX............................................................................................................................ 85


D:\Power_generation\A P O Study\REPORT.doc 1

STUDY BACKGROUND

Two basic options are available for carbon dioxide capture in gas turbine combined cycle

plants:

• Capture of the carbon dioxide from the combustion flue gases

• Capture of the carbon dioxide from the gas turbine fuel before combustion

In the first option, which is also known as post-combustion decarbonization, the carbon

dioxide may be recovered utilizing a process such as an amine wash system. However,

the disadvantage is that the carbon dioxide partial pressure in the feed gas to the wash

system is very low and results in the treatment of a large volume of gas (flue gas) with

significant energy consumption.

The second option that is also referred to as pre-combustion decarbonization, consists

of reforming or partially oxidizing the fuel (natural gas), then shifting the synthesis gas to

convert the carbon monoxide to carbon dioxide, followed by removing the carbon dioxide

and combusting the low carbon content fuel to produce power. The advantage with this

scheme is that the carbon dioxide is captured from a stream with a high partial pressure

of carbon dioxide, making it possible to reduce the size of the carbon dioxide removal

equipment and the amount of steam required for the solvent stripping operation. On the

other hand, the reforming and/or partial oxidation processes operate at high

temperatures while the shift reaction is exothermic, leading to a degradation of a portion

of the chemical energy of the fuel into thermal energy. This results in directly

transferring a portion of the fuel bound energy into the less efficient bottoming cycle

rather than into the topping cycle (gas turbine) of the combined cycle power plant.

For the success of the second option, the overall plant should be optimally configured

and the heat recovery from the raw syn gas properly integrated.



EXECUTIVE SUMMARY

The raw synthesis gas is produced by a combination of reforming and partial oxidation,

followed by shifting which results in a raw fuel gas containing carbon dioxide; 85 percent

of the carbon dioxide is captured before the resulting hydrogen rich gas is fired in a gas

turbine combined cycle.

A number of potential process improvements were considered in outline and an

assessment of their potential impact was made based on estimated overall plant thermal

efficiency and/or engineering judgement.

The results indicated that:

• Non-catalytic O2 blown partial oxidation process showed no benefits over catalytic

air partial oxidation (CAPO).

• Addition of the gas heated reformer (GHR) increases the thermal efficiency of the

plant slightly, the significant gains made in increasing the thermal efficiency of the

syn gas generation plant being offset mostly by the increased demand for steam by

the gas turbines for NOx control.

• Increasing operating pressure of the CAPO + GHR configuration reduces the overall

thermal efficiency of the plant quite significantly because of the increased steam

demand of the reformers (in order to avoid metal dusting within the GHR).

• Consequently, the use of a physical solvent process such as Selexol for carbon

dioxide capture may not be advantageously incorporated.

Performance improvements may be realized by incorporating the following design

options:

• By-passing a portion of the natural gas around the syn gas generation plant and

combining it with the fuel gas down stream of the carbon dioxide removal unit while

increasing the carbon dioxide capture level in the amine (activated MDEA) unit such

that the overall reduction in the plant carbon dioxide emissions is maintained at the

85 percent level. Another option of “by-passing” the natural gas would be to increase

the level of unconverted methane in the reformer effluent, i.e., have a significant

“methane-slip” and not by-pass around the reactors. This scheme would however,

increase the volume of gas to be treated down stream of the reformers, decrease the



partial pressure of the reactants in the shift reactors, decrease the partial pressure of

the CO2 in the CO2 removal unit, and increase the amount of chemical energy

conversion to heat in the reformer reactor(s) which would be carried out from the

reactor(s) by the methane. Note that the gas turbine will need the option of starting

up on natural gas and therefore the cost of a natural gas by-pass system is

negligible.

• Supplying the fuel gas to the gas turbines at a higher temperature. A temperature of

399 oC was selected. General Electric offers three types of fuel systems, one for fuel

gas temperatures up to 233 oC, one for gas temperatures up to 399 oC and finally

one for temperatures up to 538 oC (which tends to be a very expensive system). The

399 oC temperature was achieved by first heating the fuel gas against boiler

feedwater and then preheating it against the high temperature shift reactor effluent

gas.

The thermal and environmental performance and cost estimates were developed for the

improved cases which were configured based on the findings of the screening phase of

this study (“Improved CAPO Case” and “Improved CAPO+GHR Case”) as well as the

Base Case.

The Improved CAPO+GHR Case has the highest thermal efficiency at 44.07 percent

based on the natural gas lower heating value. The net plant output of 677 MW for this

case is lower than that for the Improved CAPO Case which is 749 MW because a large

amount of steam is removed from the steam turbine and utilized in the gas turbine for

NOx control in the Improved CAPO+GHR (2-stage reactor system) Case. The gas

turbines in this CAPO+GHR case are fully loaded because the extraction rate of the air

(used in the CAPO) is lower than the other cases. Using a CAPO+GHR gives

significant gain in efficiency of syn gas generation but the overall gain is not as

significant because more steam has to be added to the gas turbine to keep the NOx

level down. The steam demand of the gas turbine is increased for the CAPO+GHR case

in order to compensate for the reduced N2 content of the fuel gas (N2 content of the fuel

gas is lower because less oxidant is demanded by the more efficient CAPO+GHR

system). Note that steam as well as N2 function as thermal diluents in the combustor of

the gas turbine and reduce the flame temperature and thus, the thermal NOx formation.



The specific steam demand expressed in kg steam per MJ fuel input to the gas turbine,

increases from 0.04 kg/MJ for the Improved CAPO case to 0.057 kg/MJ for the

CAPO+GHR case. Thus, the result is an efficiently produced syn gas which is richer in

hydrogen than for the single CAPO reactor system but its benefits are not able to be

registered because of combustion considerations in the turbine.

While comparing the efficiency of the cases developed in this study to those published

elsewhere, it should be noted the NOx emission for these cases was set at the very low

value of 10 ppmv (dry, 15 percent O2 basis) which demanded a large amount of steam

injection into the gas turbines. Steam injection into the gas turbine degrades the thermal

efficiency of the combined cycle significantly.

The plant costs range from USD 1059/kW corresponding to the Improved CAPO Case to

USD 1105/kW corresponding to the Improved CAPO+GHR Case. The Improved

CAPO+GHR Case uses less air for the reformers and produces less volume of syn gas

(because of the reduced concentration of N2 in the gas). The resulting decrease in cost

of equipment down stream and upstream of the reformers for the Improved CAPO+GHR

Case was more than offset by the cost of the GHR. In the two CAPO only cases, the

feed/effluent interchanger down stream of the CAPO is expensive on a unit of heat

transferred basis but because of the very large temperature differences prevailing

between the gas streams exchanging the heat, the required heat transfer area is quite

low resulting in a lower absolute cost for this unit when compared to the GHR.

The cost of electricity calculated for the three cases are similar, in the range of USD

0.047/kWh to USD 0.048/kWh for natural gas priced at USD 2.00/GJ. The Improved

CAPO case shows a slightly lower cost than the other cases in the entire range of fuel

costs from USD 1.58/GJ to USD 2.64/GJ because the improved CAPO Case’s slightly

lower thermal efficiency is more than off-set by its lower plant cost when compared to

the Improved CAPO+GHR Case.

The steam demand of the reformer is increased by as much as 35 percent when the

pressure is increased from 30 barA to 46 barA in order to avoid metal dusting at the

higher pressure. In order to take full advantage of operating the reformer at higher



pressure, developmental work directed at reducing the steam consumption of the

reformer or developing materials that are not susceptible to metal dusting or a

combination of the two would be required.

Using a GHR results in an efficiently produced syn gas which is richer in hydrogen than

a single CAPO reactor system but its benefits are not able to be realized because of

combustion considerations in the turbine which demands significant amounts of steam to

keep the NOx levels down. Use of a Selective Catalytic Reduction unit (SCR) which

may cost USD 2 ½ to 3 million per gas turbine, installed down stream of the gas turbine

for NOx control should be investigated in the future. A disadvantage of the SCR system

is that it requires the handling and storage of ammonia. Catalysts are being developed

that do not require ammonia such as those installed within the combustor of the gas

turbine. Steam injection may still be required but to a limited extent. Research in the

area of gas turbine combustors aimed at minimizing the steam addition while burning H2

rich gas would be useful.

Cycles such as the Humid Air Turbine (HAT) or the Inter-cooled Steam Injected (ISTIG)

cycle should be investigated in addition to the combined cycle because these cycles

consist of or depend on introducing water vapor or steam into the combustor of the gas

turbine and NOx control is a natural outcome of these cycles without penalizing the

thermal efficiency of the cycle.

A configuration consisting of reforming the natural gas (steam-methane reforming)

utilizing at least a portion of the gas turbine exhaust to supply the combustion air to the

reformer should also be investigated. Another concept that should be considered

consists of utilizing a catalyst doped ceramic membrane reformer which is under

development. The membrane separates the H2 as it is formed such that the H2 from the

reaction mixture is constantly removed which results in maximizing the conversion of the

methane at the lower temperatures prevalent in the gas turbine exhaust.

The results of this study point the direction towards another configuration that may be

worthy of evaluation, a high pressure CAPO (without the GHR) with an expander to

recover the potential energy of the high pressure syn gas. Note that without the GHR,



the steam demand of the reformer does not increase as the pressure is increased. The

hot gas leaving the low temperature shift unit could be fed to the expander to recover

power as well as cool the gas. In this configuration, a chemical solvent such as the

amine wash would be suitable for CO2 removal.

Increasing the gas turbine fuel gas temperature to 399 oC showed a significant increase

in the thermal efficiency of the plant. However, the steam addition to the gas turbine had

to be also increased in order to control the NOx at the 10 ppmv (dry, 15 percent O2

basis). Thus, the full benefit of utilizing a higher fuel temperature was not realized. A

trade-off in order to optimize the fuel temperature is another area for further

investigation.



1.0 INTRODUCTION

This report documents the findings of a study consisting of a techno-economic

assessment of processing options for a power generation process in which

natural gas is converted to the hydrogen rich fuel gas by the pre-combustion

decarbonization process, commissioned by the International Energy Agency (IEA)

of Cheltenham UK as part of the IEA Greenhouse Gas (GHG) R& D Programme.

The raw synthesis gas is produced by a combination of reforming and partial

oxidation, followed by shifting which results in a raw fuel gas containing carbon

dioxide; carbon dioxide will be captured before the resulting hydrogen rich gas is

fired in the gas turbine combined cycle such that the carbon dioxide emissions to

the atmosphere are reduced by 85 percent. This process has been studied

previously by IEA GHG and has been shown to have the potential for commercial

applications with the ultimate goal of minimizing emission of carbon dioxide to

atmosphere while allowing the continued use of natural gas to produce electricity.

1.1 Study Objectives

The objective of the study was to assess processing options that may

improve the attractiveness of the pre-combustion decarbonization

process in natural gas fired power plants. The storage of captured

carbon dioxide was not to be considered in this study.

The initial study activity was to re-assess the process design developed in

a previous study sponsored by IEA GHG (“Reference Case”) to confirm

and/or modify the previous conclusions. A “Base Case” process design

was then developed which was the preparatory work for the main

objective of the study, consisting of seeking improvements to the process.



1.2 Study Approach

After the Base Case was established, improvements to the design were

sought and evaluated. Revised designs were then developed which

incorporated modifications considered likely to improve the overall

attractiveness of the process. The study was organized in two tasks as

follows:

1.2.1 Task 1 - Prepare and assess a Base Case Design

In the Base Case design, air and steam are used to partially oxidize

natural gas. Air is extracted from the compressor of the (combined cycle)

gas turbine, compressed and preheated in an interchanger and supplied

to the partial oxidation reactor. Air, steam, and natural gas are mixed in

the partial oxidation reactor where both partial oxidation and steam

reforming reactions take place; the reactants are brought close to

equilibrium by passing through a bed of catalyst. The reactor operates at

a pressure dictated by the gas turbine pressure ratio. Following 2-stage

shift conversion of the raw synthesis gas, carbon dioxide is captured

using a commercially available amine scrubbing process. The synthesis

gas, which is essentially a mixture of hydrogen and nitrogen, is then burnt

in the gas turbine of a combined cycle system to produce electricity. A

significant amount of steam is injected into the gas turbine to limit the

NOx emission to 10 ppmvd (15 percent excess O2 basis).

The Base Case consisted of utilizing essentially established “state-of-the-

art” technology.

1.2.2 Task 2 - Consideration of Processing Options

In this task, which was the main objective of the study, potential

improvements to the process were sought and assessed. There were

two stages to this task.



In the first stage of this task (Task 2A), a number of potential process

improvements were considered in outline and an assessment of their

potential impact was made based on estimated overall plant thermal

efficiency and/or engineering judgement.

The potential process improvements for consideration included the

following:

• Replacing the Air Feed to the catalytic air partial oxidation (CAPO)

reactor by either Oxygen or Enriched-air

• Fuel Gas Conditioning

• Alternative reactors and reactor arrangements

• A single-stage shift system

• Alternative methods for capture of carbon dioxide

• LiBr refrigeration

• Fuel Gas Compression

• Increased operating pressure

• Humidification

• Modifications aimed at optimum combination of fuel gas and gas

turbine

In the second stage of this task (Task 2B), modifications selected from

the Task 2A activity aimed at improving on the Base Case were adopted

and two revised process designs produced. These revised processes

were assessed and compared with the Base Case.



2.0 PROCESS DESIGN BASIS

The process design basis for this study is presented in the following section.

2.1 Site Conditions

Location NE coast of The Netherlands

Elevation over sea level 5 M

Heat balancecalculations

Air (60% R.H.) 15 °C

Seawater Supply (maximum) 12 °C

Seawater Supply (average) 8 °C

Seawater Return (max. allow) 19 °C

Surface Condenser Pressure 2.0 cmHgA

2.2 Products

2.2.1 Electric PowerNominal Combinedcycle (50Hz) output

800 Mwe

2.2.2 Recovered CO2Target recovery 85 %

Delivery state Liquid

Delivery point Plant limits

Compositionspecifications

CO2 >99.5 mole %

H2 <0.5 mole %

Pressure 90 Bar abs.

Temperature ≤30 °C



2.3 Raw Materials

2.3.1 Natural GasComposition CO2 1.8 Mole %

N2 0.4

CH4 83.9

C2H6 9.2

C3H8 3.3

C4H10+ 1.4

Total 100

Maximum sulfur 30 mg/Nm3

Pressure 52 Bar abs.

sensitivity to 20 Bar abs.

Temperature 15 °C

Price 2 USD/GJ (LHV)

2.3.2 Municipal WaterPressure 11 BarA

2.4 Plant ConfigurationPlant Heat Rejection Once through seawater cooling

Plant Water Makeup Municipal water

Gas Turbine Two trains (General Electric MS9001FA)

Steam Turbine Single train

2.5 Environmental CriteriaNOx < 10 ppmv (dry, 15 % O2)

Particulate matter < 25 mg/Nm3



3.0 SCREENING ANALYSES

A number of potential process improvements were considered in outline and an

assessment of their potential impact was made based on estimated overall plant

thermal efficiency and/or engineering judgement. The Reference Case was also

revisited and its configuration modified during this screening analyses resulting in

the Base Case. The process modifications considered involved either “state-of-

the-art” technology or those requiring little process development with no obvious

major technical problems to be solved.

In assessing the potential attractiveness of process improvements, factors

considered included cost, power generation efficiency, carbon dioxide capture

efficiency, confidence that the process will work and the level of development

required, safety and environmental implications for the general acceptability of

the process, and the applicability of the process e.g. in terms of restraints on

size, or availability of required components and materials.

3.1 Unit Process Considerations

3.1.1 Fuel Gas Compression

A previous design (the Reference Case) developed for a Catalytic

Autothermal Partial Oxidation (CAPO) based combined cycle with CO2

capture consisted of extracting the air from the gas turbine compressor

and using it directly in the CAPO followed by syn gas compression

upstream of the CO2 removal unit. All the configurations (including that

for the Base Case) that were utilized and evaluated in this study

consisted of compression of the air extracted from the gas turbine for the

CAPO, instead. The extracted air was cooled, compressed and again

heated resulting in some additional equipment. On the other hand, the

compression power and size of the compression equipment were reduced

since the gas volume to be compressed was almost half in the case of air



compression when compared to synthesis gas compression.

Furthermore, air compressors are simpler in design when compared to

compressors for syn gas containing large concentrations of hydrogen.

This selected configuration also results in a syn gas generation plant

design that operates at a higher (and thus near a more optimum)

pressure than the case where air compression is not utilized for a given

gas turbine pressure ratio when considering the trade-off between the

size of equipment and equipment wall thickness.

3.1.2 Single-stage Shift System

Adiabatic shift reactors operating at the highest exit temperatures

compatible with the catalyst result in maximizing the power generation

efficiency. The temperature of heat generated by the shift reactors is high

and may either be recovered for generation of steam at a high pressure

and exported to the bottoming cycle of the combined cycle where it may

be converted to power at a high efficiency, or the heat may be utilized for

preheating the fuel gas to a high temperature before it is combusted in

the gas turbines, whereby the overall efficiency of the plant is increased.

On the other hand, when a single stage shift reactor is utilized in a plant

designed to achieve the 85 percent carbon dioxide capture, operation of

the reactor near isothermal conditions is required resulting in the

generation of heat by the reactor at a much lower temperature. The

conversion of this lower temperature heat to power (by the generation of

lower pressure steam) is less efficient.

Thus, all the configurations evaluated in this study (including the Base

Case) consisted of utilizing the adiabatic shift reactors.

3.1.3 Alternative Methods for Capture of Carbon Dioxide

In addition to solvent based carbon dioxide capture processes such as

use of chemical solvents (amine) and physical solvents (Selexol),



cryogenic separation and Pressure Swing Adsorption (PSA) or

combination of membranes and PSA were considered. The energy

consumption of a cryogenic process is high while with a PSA; a significant

portion of the hydrogen is rejected by the PSA. Recovery of the hydrogen

from the “tail gas” of the PSA may be achieved but requires a significant

amount of power. Thus only the amine (activated MDEA) based and

Selexol based solvents were further evaluated in this study.

3.1.4 LiBr Refrigeration

Use of LiBr adsorption refrigeration, which may be operated utilizing low

temperature heat was considered in order to recover low temperature

heat and produce chilled water for cooling the gas streams (air of the

CAPO and the product carbon dioxide) before and during the

compression operations and the liquefaction operation in case of the

carbon dioxide in order to minimize the power consumption. However,

during the development of the process configuration of the various cases

it was determined that the useful low temperature heat generated by the

plant could all be recovered for deaerator feedwater heating in addition to

recovering the low temperature heat from the gas turbine exhaust gas.

Furthermore, with the very cold seawater available for cooling, the extra

capital required by the LiBr refrigeration did not appear justifiable.

Thus, LiBr refrigeration was dropped from any further consideration and

the configurations that were further evaluated in this study did not include

the LiBr refrigeration.

3.1.5 Humidification

Natural gas and/or fuel gas humidification provides means for capturing

low temperature heat from the process and also a means of disposing

waste water (dissolved solids content of the water, if any, should be

closely monitored). Thus, natural gas humidification could reduce the



amount of steam removed from the combined cycle for injection into the

CAPO, while fuel gas humidification could reduce the amount of steam

removed from the combined cycle for injection into the gas turbine for

NOx control.

However as mentioned previously, during the development of the process

configuration of the various cases, it was determined that the useful low

temperature heat generated by the plant could all be recovered for

deaerator feedwater heating in addition to recovering the low temperature

heat from the gas turbine exhaust gas.

Thus, humidification was dropped from any further consideration and the

configurations that were further evaluated in this study did not include the

humidification operation. In future studies where configurations are

developed that generate significant amounts of low temperature heat,

humidification of the fuel gas should be given consideration.

3.1.6 Gas Turbine

Aeroderivative gas turbines were also given consideration in this initial

screening phase of the study. The General Electric LM6000 was chosen

as being representative of an aero-derivative gas turbine. For a nominal

700 MW power plant, as many as 10 to 11 trains may be required while

with the heavy frame engines, as represented by the General Electric MS

9001FA, only two trains would be required. Thus, economies of scale are

maximized when these heavy frame engines are used instead.

Furthermore, for combined cycle applications, the combination of firing

temperature and pressure ratio of the MS 9001FA is nearer to the

optimum which provides a higher overall plant thermal efficiency.

General Electric has stated that this machine has the capability of air

extraction (for the CAPO), while the small combustor of the LM6000 may

not be suitable for a low calorific value fuel gas. Furthermore, the LM



6000 with its significantly higher pressure ratio would require the

reformers to be operated at a higher pressure which would increase the

steam demand of the reformers in cases incorporating the CAPO+GHR

configuration in order to avoid metal dusting in the GHR. This increased

steam demand of the reformers reduces the overall plant thermal

efficiency as shown later in this section.

The next generation gas turbines such as the steam cooled General

Electric MS9001H may be an excellent machine for combined cycle

applications in the future. However, its higher pressure ratio could

potentially be a disadvantage in these applications. General Electric has

stated that they would not be able to provide any support in terms of cost

and performance for the gas turbine since it is still under development.

General Electric has done significant amount of work on burning low

calorific value gas with the combustor used in the MS 9001FA machine.

Note that the MS 9001FA is a geometric scale-up of the MS 7001FA (a

60 Hz machine) which has been demonstrated in coal based IGCC plants

burning low calorific value fuel gas. General Electric has stated that only

confirmation tests would be required with MS 9001FA machine to burn

the hydrogen rich fuel gas. Note however that in IGCC plants, the

combustibles in the fuel gas consist of primarily hydrogen and carbon

monoxide whereas in the plants where carbon dioxide is removed prior to

combustion, the combustible in the gas turbine fuel gas is primarily

hydrogen which has a much higher flame speed that carbon monoxide.

Other candidate engines include the large heavy frame engines offered

by Siemens/Westinghouse, Mitsubishi and Asea Brown Boveri.

3.1.7 Alternative Reactors and Reactor Arrangements

Alternate arrangements consisting of recovering and recycling the heat

contained in the reactor effluent (gas heated reformers with internal heat



exchangers) can increase the thermal efficiency of the partial oxidation

unit. Thus, a variant consisting of utilizing the CAPO unit effluent to

supply the endothermic heat required by a Gas Heated (steam-methane)

Reformer (GHR) located downstream of the CAPO was evaluated. The

advantage of this reactor arrangement is that a majority of the exhaust

heat from the CAPO is captured and converted into chemically bound

energy in the GHR with the potential for a higher overall plant thermal

efficiency since more energy is made available for conversion to power in

a combined cycle and not transferred directly into only the bottoming

cycle by the generation of steam. Cases 2, 3, 5 and 6 were developed for

this screening analysis and the results of this evaluation are discussed

later in this section.

Catalytic versus non-catalytic Autothermal Partial Oxidation (POX) was

also evaluated in this screening analysis task of the study. The CAPO

has the advantage of operating at a lower temperature than the non-

catalytic POX resulting in a higher “cold gas efficiency” defined by the

ratio of the chemically bound energy contained in the reactor effluent gas

(syn gas) and the natural gas feed. Thus, with the CAPO, the potential

exists for a higher overall plant thermal efficiency since more energy is

available for conversion to power in a combined cycle and not transferred

directly into the bottoming cycle. On the other hand the non-catalytic

POX unit is simpler in design and does not require significant amount of

steam injection into the reactor. Potential exists for soot formation,

however. Case 4 was developed to evaluate the non-catalytic POX and

the results of this evaluation are discussed later in this section.

3.1.8 Replacing Air Feed to POX by Oxygen or Enriched-air

The thermal efficiency of the POX reactor is higher with oxygen since a

smaller portion of the fuel bound chemical energy is transformed to heat,

which is carried out of the POX by the effluent. The thermal efficiency is

the lowest with air since the POX effluent flow rate is the highest due to

the nitrogen present in the air. Furthermore, the size of the POX as well



as the equipment down stream of the POX is minimized when oxygen is

used as the POX oxidant. The carbon dioxide capture process also

benefits from the higher partial pressure of the carbon dioxide with an

oxygen blown POX. However, the oxygen blown POX plant with its air

separation unit requires significant amount of electric power for its

operation as well as capital expenditure. Thus, a trade-off exists and this

trade-off becomes significant as the POX operating temperature

increases.

A review of the findings made by Fluor Daniel and others in previous

techno-economic evaluations of comparing oxygen versus enriched air

versus air blown systems has shown that oxygen of 95 mole percent

purity is the oxidant of choice for high temperature POX applications.

Furthermore, an air separation unit operating at elevated pressure such

that the nitrogen separated from the air is available at pressure for

injection into the gas turbine for NOx is preferred over a conventional low

pressure air separation unit (ASU) from an economic basis. The nitrogen

stream is used as a thermal diluent and also serves as additional motive

fluid in the expander section of the gas turbine. The advantage with this

ASU configuration together with nitrogen injection to the gas turbine over

a configuration utilizing a conventional low pressure ASU, is that the

overall power generation efficiency (of the oxygen blown plant) is

increased, while the concerns of burning a hydrogen rich fuel and NOx

emissions are mitigated.

For lower operating temperatures, however, the air blown configuration is

preferred since the sensible heat carried by the effluent is lower than a

high temperature non-catalytic POX, while the inefficiency and capital

cost associated with air separation are eliminated. Thus, an oxygen

blown non-catalytic POX was developed (Case 4) to evaluate this

configuration and the results of this evaluation are discussed later in this

section.



3.1.9 Increased Operating Pressure

Increasing the operating pressure decreases the size of the process

equipment in the syn gas generation plant but also increases the

methane leakage from the autothermal partial oxidation unit according to

the equilibrium constraint of the reforming reaction: p3H2 pCO/(pCH4 pH2O) =

K p, where p i is the partial pressure of component i and K p is the

equilibrium constant. As can be seen from the above expression, the

concentration of methane in the reformer effluent increases as the

pressure increases.

On the other hand, a physical solvent process for carbon dioxide capture

that has potential to reduce energy consumption may be advantageously

applied. A fuel gas expander located upstream of the gas turbine may

also be included in the design in order to enhance the efficiency. Two

configurations for inclusion of a fuel gas expander in high pressure partial

oxidation plants are possible, one option consists of utilizing a hot gas

expander to maximize the power produced by the expander while a

second option consists of utilizing a cold gas expander to recover power

as well as provide refrigeration which may be utilized in the physical

solvent based carbon dioxide capture process. It was determined,

however that the refrigeration made available downstream of the

expander utilizing the “Cold Gas Expander” configuration was only a

minor fraction of the total amount of refrigeration required by the physical

solvent (Selexol) process and thus this concept was dropped from further

consideration. Cases 3, 4 and 5 were developed to evaluate the higher

operating pressure of the syn gas generation plant utilizing the “Hot Gas

Expander” which maximizes the power generation by the expander. The

results of this evaluation are discussed later in this section.



3.2 Configuration Development

The cases developed based on these above considerations are described

in the following. The results of a quantitative analysis of their thermal

efficiencies are also presented in order to select the cases for further

analyses.

3.2.1 Base Case (Case 1)

The Base Case as depicted in Figure 3 - 1 consists of two General

Electric MS 9001FA gas turbines and air extraction from the gas turbine

compressor to supply air required by the CAPO. The extracted air is

cooled, compressed and then preheated against the hot extraction air

before being fed into the CAPO. The CAPO unit operates at a pressure

of 30 barA in order to supply the fuel gas to the gas turbine at its required

pressure. Syn gas compression is not utilized.

Two beds of adiabatic shift reactors convert the carbon monoxide to

carbon dioxide while an amine (activated MDEA) unit removes the carbon

dioxide in order to reduce the carbon dioxide emissions from the flue gas

by 15 percent. The NOx emission from the gas turbines is limited to 10

ppmvd (15 percent O2 basis) by steam injection. The natural gas that is

assumed to be supplied at 52 barA is preheated and expanded in a turbo-

generator to a pressure as required by the CAPO.

3.2.2 Case 2

Case 2 (depicted in Figure 3 - 2) is similar to the Base Case except that

the CAPO is followed by a GHR. Thus, the reforming process is

JULY 1999

DESIGNED BY

CHECKED BY

PROCESS ENGR.

LEAD ENGR. / SPECIALIST

PROJECT

CLIENT

APP. DATE

APP. DATE

APP. DATE

APP. DATE

REV.DRAWING NUMBERSCALE

NONE

FIGURE 1 OVERALL BLOCK FLOW DIAGRAM

BASE (CAPO) CASE`FLU O R D AN IELD. FRANCUZ

A. RAO

FIGURE 1 A

DATEREV.NO.

REVISION DESCRIPTION DRAWN APP.

REFERENCE DWG. NO.DWG. NO.DWG. NO. REFERENCE DWG. NO.

CHECK

COMPRESSEDAIR SUPPLY

CATALYTICPARTIAL

OXIDATION(CAPO)30 BarA

COCONVERSION

(SHIFT)

CO2RECOVERY

(AMINE)

GASTURBINES

CO2LIQUEFACTION

STEAMTURBINE

HRSGS

NATURAL GASEXPANSION &PURIFICATION

EXHAUST

NOx CONTROL STEAM

STEAM

NATURALGAS

FROMPIPELINE52 BarA

EXTRACTION AIR

CO2PRODUCT

HYDROGENRICH GAS

COMPRESSION

PREHEATED FUEL GAS

RECYCLE GAS

JULY 1999

DESIGNED BY

CHECKED BY

PROCESS ENGR.


PROJECT

CLIENT

APP. DATE

APP. DATE

APP. DATE

APP. DATE


NONE


CAPO +GHR CASE`FLU O R D AN IELD. FRANCUZ

A. RAO

FIGURE 2 A

DATEREV.NO.



CHECK


CATALYTICPARTIAL

OXIDATION(CAPO) & GAS

HEATEDREFORMER

(GHR) 30 BarA

COCONVERSION

(SHIFT)

CO2RECOVERY

(AMINE)

GASTURBINES

CO2LIQUEFACTION

STEAMTURBINE

HRSGS

NATURAL GASEXPANSION &PURIFICATION

EXHAUST

NOx CONTROL STEAM

STEAM

NATURALGAS

FROMPIPELINE52 BarA

EXTRACTION AIR

CO2PRODUCT

HYDROGENRICH GAS

COMPRESSION

PREHEATED FUEL GAS

RECYCLE GAS



conducted partially in the CAPO and partially in the GHR to recover the

exhaust heat from the CAPO as chemical energy. Two General Electric

MS 9001FA gas turbines with air extraction from the gas turbine

compressor to supply air required by the CAPO are utilized and the NOx

emission is limited to 10 ppmvd (15 percent O2 basis) by steam injection.

The natural gas that is assumed to be supplied at 52 barA is preheated

and expanded in a turbo-generator to a pressure as required by the

CAPO. An amine (activated MDEA) process removes the CO2.

This case quantifies the advantages of adding a GHR. As seen from the

performance estimates presented in Table 3-1 (included at the end of

Section 3), the overall plant thermal efficiency is increased only slightly

over the Base Case. The cold gas efficiency was increased resulting in

less air being required by the CAPO. This, however, resulted in a fuel

gas with lower concentration of nitrogen, and to limit the NOx emission

from the gas turbine to 10 ppmvd (15 percent O2 basis), the steam

injection rate to the gas turbine had to be increased. This increased

steam extraction rate from the steam turbine decreased the combined

cycle efficiency and offset to a significant extent the gains made in

increasing the efficiency of syn gas generation plant. This case requires

an additional reactor (GHR) but the size of the equipment in the syn gas

generation plant is much smaller because of the lower concentration of

nitrogen in the syn gas. Furthermore, the power consumption of the

oxidant feed supply to the CAPO unit as well as the carbon dioxide

removal unit are reduced. The gas turbines are more fully loaded than in

the Base Case since the air extraction rate is lower for Case 2.

3.2.3 Case 3

Case 3 (depicted in Figure 3 - 3) is similar to Case 2 except that the syn

gas generation occurs at high pressure, near 46 barA. The pressure was

limited to 46 barA based on recommendations provided by ICI Synetex.

JULY 1999

DESIGNED BY

CHECKED BY

PROCESS ENGR.


PROJECT

CLIENT

APP. DATE

APP. DATE

APP. DATE

APP. DATE


NONE

FIGURE 3 OVERALL BLOCK FLOW DIAGRAMHIGH PRESSURE CAPO +GHR CASE`FLU O R D AN IEL

D. FRANCUZ

A. RAO

FIGURE 3 A

DATEREV.NO.



CHECK


CATALYTICPARTIAL


HEATEDREFORMER

(GHR) 46 BarA

COCONVERSION

(SHIFT)

CO2RECOVERY(SELEXOL)

GASTURBINES

CO2LIQUEFACTION

STEAMTURBINE

HRSGS

NATURAL GASPURIFICATION

EXHAUST

NOx CONTROL STEAM

STEAM

NATURALGAS

FROMPIPELINE52 BarA

EXTRACTION AIR

CO2PRODUCT

PREHEATED FUEL GAS

RECYCLE GAS

FUEL GASPREHEAT AND

EXPANSION

HYDROGENRICH GAS

COMPRESSION



Increasing the operating pressure of the GHR above 46 barA should be

carefully evaluated taking into account the increase in the cost of the

GHR and the increase in its steam demand. The CO2 is removed in a

Selexol process and the syn gas leaving the Selexol process is preheated

and expanded in a turbo-generator to a pressure as required by the gas

turbine. The thermal efficiency of this plant is significantly lower than any

of the cases evaluated, the reason being that the steam and power

consumption of the carbon dioxide removal process were significantly

increased over the other cases. The operating pressure of the system

was not high enough to take full advantage of the Selexol process. Note

that the feed gas to the CO2 removal process contains a significant

concentration of N2 that reduces the feed gas CO2 partial pressure which

is the principal driving force for absorption in the column of a physical

solvent process. Furthermore, the steam consumption of the reforming

unit increased significantly over the lower pressure cases in order to

avoid metal dusting in the GHR. The steam demand for the reformer is

calculated based on the Boudard reaction equilibria considerations. This

increased steam demand further decreased the overall plant efficiency.

3.2.4 Case 4

Case 4 (depicted in Figure 3 - 4) is similar to Case 3 except that the

reactor consists of a high pressure (83 barA) non-catalytic partial

oxidation (POX) unit and the oxidant consists of 95 percent purity oxygen

supplied by an elevated pressure air separation unit. The high pressure

nitrogen generated in the air separation unit is further compressed and is

injected into the gas turbine for NOx control. Steam is also supplied to

the gas turbine since the nitrogen does not meet the entire thermal diluent

requirement of the gas turbine.

The large parasitic power consumption of the ASU as well as the lower

cold gas efficiency of the POX unit, which operates at a much higher

JULY 1999

DESIGNED BY

CHECKED BY

PROCESS ENGR.


PROJECT

CLIENT

APP. DATE

APP. DATE

APP. DATE

APP. DATE


NONE


HIGH PRESSURE POX CASE`FLU O R D AN IELD. FRANCUZ

A. RAO

FIGURE 4 A

DATEREV.NO.



CHECK

ELEVATEDPRESSURE

AIRSEPARATION

PARTIALOXIDATION

(NON-CATALYTIC)

83 BarA

COCONVERSION

(SHIFT)

CO2RECOVERY(SELEXOL)

GASTURBINES

CO2LIQUEFACTION

STEAMTURBINE

HRSGS

NATURAL GASCOMPRESSION

ANDPURIFICATION

EXHAUST

ADDITIONAL NOxCONTROL STEAM

STEAM

NATURALGAS

FROMPIPELINE52 BarA

EXTRACTION AIR

CO2PRODUCT

PREHEATED FUEL GAS

RECYCLE GAS

FUEL GASPREHEAT AND

EXPANSION

HYDROGENRICH GAS

COMPRESSION

95% O2

HIGH PRESSURE NITROGEN FOR NOx CONTROL



temperature than the CAPO and the GHR, contributed to lowering the

overall plant efficiency.

3.2.5 Case 5

Case 5 (depicted in Figure 3 - 5) attempted to combine the advantageous

design features of the previous cases. The design consists of a CAPO

followed by a GHR operating near 46 barA. The high pressure syn gas is

expanded in a turbo-generator after exiting the second shift reactor (the

shift reactor provided the preheating of the gas). The lower pressure gas

is then fed to an amine (activated MDEA) unit for CO2 removal since as

seen from case 3, the Selexol process caused a decrease in the plant

efficiency. This case however, did not show a thermal efficiency gain

over the lower pressure Case 2, the gain made by increasing the output

of the expander (by expansion of the gas after the shift unit) being offset

by the increased steam demand of the GHR and the gas turbine.

3.3 Conclusions and Recommendations

Based on the results developed from these five cases, the following

conclusions may be made:

• Non-catalytic O2 blown partial oxidation process showed no benefits

over CAPO.

• Increasing operating pressure of the CAPO + GHR configuration

reduces the overall thermal efficiency of the plant quite significantly

because of the increased steam demand of the reformer (in order to

avoid metal dusting within the GHR).

JULY 1999

DESIGNED BY

CHECKED BY

PROCESS ENGR.


PROJECT

CLIENT

APP. DATE

APP. DATE

APP. DATE

APP. DATE


NONE


HIGH PRESSURE CAPO +GHR CASE WITHSHIFTED GAS EXPANSION

`FLU O R D AN IELD. FRANCUZ

A. RAO

FIGURE 5 A

DATEREV.NO.



CHECK


CATALYTICPARTIAL


HEATEDREFORMER

(GHR) 46 BarA

COCONVERSION

(SHIFT)

CO2RECOVERY

(AMINE)

GASTURBINES

CO2LIQUEFACTION

STEAMTURBINE

HRSGS


EXHAUST

NOx CONTROL STEAM

STEAM

NATURALGAS

FROMPIPELINE52 BarA

EXTRACTION AIR

CO2PRODUCT

HYDROGENRICH GAS

COMPRESSION

PREHEATED FUEL GAS

RECYCLE GAS

SYN GASEXPANSION



• Consequently, the use of a physical solvent process such as Selexol

for carbon dioxide capture may not be advantageously incorporated.

• Addition of the GHR increases the thermal efficiency of the plant

slightly, the significant gains made in increasing the thermal efficiency

of the syn gas generation plant being offset mostly by the increased

demand for steam by the gas turbines for NOx control.

Based on these findings, two additional cases (“Improved CAPO Case”

and “Improved CAPO+GHR Case”) were developed which improved upon

the Base Case and Case 2, respectively. These cases are depicted in

Figures 3 - 6 and 3 - 7. The improvements consisted of:

• by-passing a portion of the natural gas around the syn gas generation

plant and combining it with the fuel gas down stream of the carbon

dioxide removal unit while increasing the carbon dioxide capture level

in the amine unit such that the overall reduction in the plant carbon

dioxide emissions is maintained at the 85 percent level.

• supplying the fuel gas to the gas turbines at a higher temperature. A

temperature of 399 oC was selected. General Electric offers three

types of fuel systems, one for fuel gas temperatures up to 233 oC, one

for gas temperatures up to 399 oC and finally one for temperatures up

to 538 oC (which tends to be a very expensive system). The 399 oC

temperature was achieved by first heating the fuel gas against boiler

feedwater and then preheating it against the high temperature shift

effluent gas. Heating the gas to 538 oC would require a significantly

more expensive fuel system and also a heat exchanger with

expensive metallurgy to preheat the fuel gas. Furthermore, the

resulting increase in efficiency would be marginal since the heat

required for preheating the fuel gas would be at the expense of high

pressure steam generation; only a 0.5 percent reduction in the net

heat rate may be expected.



These cases show a significant increase in the thermal efficiency over the

Base Case and were selected for further analyses consisting of

estimating the plant costs and calculating the levelized cost of electricity.

JULY 1999

DESIGNED BY

CHECKED BY

PROCESS ENGR.


PROJECT

CLIENT

APP. DATE

APP. DATE

APP. DATE

APP. DATE


NONE


IMPROVED CAPO CASE`FLU O R D AN IELD. FRANCUZ

A. RAO

FIGURE 6 A

DATEREV.NO.



CHECK


CATALYTICPARTIAL

OXIDATION(CAPO)

REFORMER(GHR) 30 BarA

COCONVERSION

(SHIFT)

CO2RECOVERY

(AMINE)

GASTURBINES

CO2LIQUEFACTION

STEAMTURBINE

HRSGS


EXHAUST

NOx CONTROL STEAM

STEAM

NATURALGAS

FROMPIPELINE52 BarA

EXTRACTION AIR

CO2

PRODUCT

HYDROGENRICH GAS

COMPRESSION

HOT FUEL GAS 399 °C

RECYCLE GAS

NATURAL GASPREHEAT AND

EXPANSION

FUEL GASPREHEAT

NATURAL GAS BY-PASS

JULY 1999

DESIGNED BY

CHECKED BY

PROCESS ENGR.


PROJECT

CLIENT

APP. DATE

APP. DATE

APP. DATE

APP. DATE


NONE


IMPROVED CAPO +GHR CASE`FLU O R D AN IELD. FRANCUZ

A. RAO

FIGURE 7 A

DATEREV.NO.



CHECK


CATALYTICPARTIAL


HEATEDREFORMER

(GHR) 30 BarA

COCONVERSION

(SHIFT)

CO2RECOVERY

(AMINE)

GASTURBINES

CO2LIQUEFACTION

STEAMTURBINE

HRSGS


EXHAUST

NOx CONTROL STEAM

STEAM

NATURALGAS

FROMPIPELINE52 BarA

EXTRACTION AIR

CO2

PRODUCT

HYDROGENRICH GAS

COMPRESSION

HOT FUEL GAS 399 °C

RECYCLE GAS

NATURAL GASPREHEAT AND

EXPANSION

FUEL GASPREHEAT

NATURAL GAS BY-PASS



Table 3-1Thermal Performance Estimates

CaseDesignation

BaseCase

(Case 1)

2 3 4 5 ImprovedCAPOCase

ImprovedCAPO +

GHR CaseReactor

TypeCAPO CAPO+

GHRCAPO+

GHRPOX CAPO+

GHRCAPO CAPO+

GHRReactor

Pressure,barA

30 30 46 83 46 30 30

Oxidant Air Air Air 95% O2 Air Air AirCO2

RemovalAmine Amine Selexol Selexol Amine Amine Amine

Natural GasBypass, %

0 0 0 0 0 7.84 7.73

Fuel Input,MW (LHV)

1740 1565 1533 1708 1571 1708 1537

GT Power,MW

544 568 570 549 572 566 572

ST Power,MW

289 179 95 257 169 272 174

ExpanderPower, MW

1.5 1.2 11 16 18 1.3 1.2

GrossPower, MW

835 748 676 822 759 839 747

Aux. Power,MW

90 75 129 127 83 90 70

Net Power,MW

745 673 547 695 676 749 677

ThermalEfficiency,

% LHV42.8 43.1 35.7 40.7 43.0 43.8 44.1



4.0 BASE CASE PROCESS DESIGN

The Base Case consists of two General Electric MS 9001FA gas turbines and air

extraction from the gas turbine compressor to supply air required by the CAPO.

The extracted air is cooled, compressed and then preheated against the hot

extraction air before being fed into the CAPO. The CAPO unit operates at a

pressure of 30 barA in order to supply the fuel gas to the gas turbine at its

required pressure. Syn gas compression is not utilized.

Two beds of adiabatic shift reactors convert the carbon monoxide to carbon

dioxide while an amine (activated MDEA) unit removes the carbon dioxide in

order to reduce the carbon dioxide emissions from the flue gas by 85 percent.

The NOx emission from the gas turbines is limited to 10 ppmvd (15 percent O2

basis) by steam injection. The natural gas that is assumed to be supplied at 52

barA is preheated and expanded in a turbo-generator to a pressure as required

by the CAPO.

The material balance for the Base Case is attached in the Appendix.

Oxidant Supply

Refer to Process Flow Diagram, Oxidant Supply Base Case (drawing number

592200-001) for the following discussion. This drawing depicts the compression

and preheating of the oxidant (air) supply.

Air extracted from the gas turbine air compressor at 404 oC is cooled in the

CAPO Air Preheater (10-HX-1). The air is further cooled by preheating deaerator

feedwater in the Deaerator Feed Water Heater (10-HX-2). The cooled air is

passed through the Extraction Air Knock Out (KO) Drum (10-VE-1), which

separates out condensate from the cooled air. The condensate from the

compressed air is of good quality and is routed to the make-up water tank.



The cooled air is then fed to the CAPO Air Compressor driven by a Motor (10-

CM/MO-1) which compresses the air up to the pressure as required for the

CAPO Reactor. The compressed air after preheating to 378 oC against the

extracted air in 10-HX-1 is supplied to the CAPO.

Syn Gas Generation

Refer to Process Flow Diagram, Syn Gas Generation Plant Base Case (drawing

number 592200-002) for the following discussion. This drawing depicts the

natural gas expansion, syn gas generation, and syn gas cooling steps.

Natural gas enters the plant from the pipeline at 52 barA. The natural gas is

preheated against boiler feedwater in the Natural Gas Heater (20-HX-1) before

the pressure is reduced in the Natural Gas Expander (20-EX-1), which produces

electric power.

A portion of the hydrogen rich fuel gas from the CO2 Recovery Plant is

compressed in the Motor Driven Recycle Hydrogen Compressor (20-CM/MO-1)

and is mixed with the expanded natural gas to maintain a H2 concentration of 2

to 5 mole percent in the feed gas. The feed gas consisting of the natural

gas/hydrogen mixture is preheated in the Desulfurizer Feed Preheater (20-HX-2)

by cooling a portion of the hot effluent from the High Temperature CO Converter.

Trace amounts of sulfur in the natural gas are then removed in the Natural Gas

Purifier (20-RR-2).

High pressure steam from the combined cycle plant is then added to the purified

natural gas. The natural gas/hydrogen/steam mixture is preheated by cooling the

effluent from the CAPO reactor in the CAPO Feed Preheater (20-HX-3).

Preheated air from the Oxidant Supply unit (described previously) and the pre-

heated natural gas/hydrogen/steam mixture enter the top of the CAPO (20-RR-1)

and flow down through the catalyst bed where the methane is reacted with steam



to form H2, CO and CO2. The oxidation reaction is exothermic and the reactor is

adiabatic so that the effluent from the reactor at 895 oC is much hotter than the

feed which is at 512 oC. As described above the feed to the reactor is heated by

cooling the effluent. Additional heat is recovered from the CAPO effluent by

producing high pressure (HP) steam in the HP Steam Generator (20-HX-4). The

steam is superheated in the Heat Recovery Steam Generators (HRSGs) before it

is sent to the steam turbine.

The catalyst in the CAPO is a nickel based shaped catalyst (pelleted with holes)

with approximate size of 1 cm diameter and 2 cm length. Note that the exact set

of conditions used in the study for the CAPO may not have been demonstrated

on a large scale. However, similar units have been built as secondary reformers

in large ammonia plants with air feed. Furthermore, the temperature conditions

below the burner and above the catalyst for the CAPO are less severe than

oxygen based secondary reformers or auto thermal reformers which in addition

have been run at lower steam addition rates.

The carbon monoxide in the syn gas is converted to hydrogen in two catalytic

stages via the “CO shift” reaction (CO + H2O ⇋ H2 + CO2).

The syn gas from the HP Steam Generator which is within the initiation

temperature range of the High Temperature CO Converter (20-RR-3) enters the

reactor directly from the steam generator. The reaction is exothermic and the

catalytic reactor is adiabatic so that the effluent temperature is hotter than the

feed temperature. A portion of the effluent from the High Temperature CO

Converter is used to heat the feed to the Natural Gas Purifier as described

previously. The remaining syn gas from the High Temperature CO Converter is

used to preheat the fuel gas from the CO2 Recovery Plant in the Fuel Gas Heater

(20-HX-5). The preheated fuel gas is then sent to the Combined Cycle Plant

(CCP). This recovery of heat into the topping cycle increases the Thermal

efficiency of the plant.



The syn gas from the two preheaters is recombined and sent to the Low

Temperature CO Converter (20-RR-4) where the CO concentration is further

reduced. The CO concentration is reduced from approximately 8 volume percent

to less than 0.5 volume percent in the two converters.

The syn gas is cooled in two stages before it is sent to the CO2 recovery unit.

First, the syn gas provides heat for stripping the amine solvent (CO2 Recovery

Unit) in the Gas Heated Reboiler (30-HX-7A/B). The syn gas is then further

cooled by heating deaerator feedwater in the Deaerator Feed Heater (20-HX-6).

Water that condenses from the syn gas in the cooling process is separated from

the syn gas in the Process Gas Separator (20-VE-1). The condensate is

polished in Condensate Polishing Unit (50-ME-3) before it is recycled to the

Demineralized Water Tank.

The syn gas from the Process Gas Separator is sent to the CO2 recovery plant.

CO2 Recovery Plant

Refer to Process Flow Diagram, CO2 Recovery Base Case (drawing number

592200-003) for the following discussion. This drawing depicts the amine CO2

removal from the syn gas and recovery.

Syn Gas Treating

The syn gas from the syn gas generation plant is mixed with repressurized flash

gas and enters the bottom of the CO2 Absorber (30-VE-1). The syn gas is

counter-currently contacted with the amine (activated MDEA) in a two stage

packed tower. The amine absorbs the CO2 from the syn gas and flows out the

bottom of the absorber. The clean syn gas flows out the top of the absorber and

is sent to the Syn Gas Generation Plant for preheating.



Amine Circulation

The rich (CO2 bearing) amine leaving the bottom of the absorber is flashed in the

HP Flash Drum (30-VE-2) to recover a portion of the hydrogen that is co-

absorbed. The flashed gas is repressurized in the two-stage Recycle Gas

Compressor (30-CM/MO-1). The compressor has an air-cooled Compressor

Intercooler (30-HX-1) which rejects the inter-stage heat to the atmosphere.

The lean solvent from the bottom of the CO2 Stripper heats the amine solvent

leaving the HP Flash Drum in the Rich/Lean Solvent Interchanger (30-HX-2).

The rich solvent is then introduced into the top of the CO2 Stripper (30-VE-3).

Steam strips the amine solvent of the CO2 as it flows down through two packed

sections. A gas-heated reboiler and a steam-heated reboiler are used to raise

the stripping steam in the vessel. Heat is recovered from the syn gas in the Gas

Heated Reboiler (30-HX-7). LP steam is used in the Steam Heated Reboiler

(30-HX-6) to provide the additional heat necessary for stripping the solvent.

The hot vapor exiting the top of the stripper is sent to the air-cooled Stripper

Overhead Condenser (30-HX-5). The steam is condensed and separated from

the CO2 in the Stripper Condensate Accumulator (30-VE-4). The CO2 leaving the

top of the accumulator is sent to the Compression, Drying, and Liquefaction

Plant. The condensate is pumped back to the top of the CO2 Stripper via the

Stripper Condensate Pump (30-PV-3).

The amine solvent exits the CO2 stripper from two locations. Semi-lean solvent

is collected on a chimney tray after having the bulk of the CO2 removed in the

first stripping section. The Semi-Lean Solvent Pump (30-PV-2) transfers the

semi-lean solvent from the chimney tray of the CO2 Stripper to the middle section

of the CO2 Absorber. The semi-lean solvent is air-cooled in the Semi-Lean

Solvent Cooler (30-HX-4) before it enters the column. The lean solvent from the

bottom of the CO2 Stripper is transferred via the Lean Solvent Pump (30-PV-1) to

the top section of the CO2 absorber. The solvent is cooled as it passes through



the Rich/Lean Solvent Interchanger and the air-cooled Lean Solvent Cooler

(30-HX-3) before it re-enters the CO2 absorber.

CO2 Compression, Drying and Liquefaction

Refer to Process Flow Diagram, CO2 Compression, Drying and Liquefaction

Plant - Base Case (drawing number 592200-004) for the following discussion.

This drawing depicts CO2 compression in a multi-stage compressor, drying in a

glycol contactor and liquefaction by condensation. Drying of the CO2 is required

in order to avoid corrosion in the pipeline. Note that wet CO2 causes severe

corrosion of carbon steel.

The CO2 from the CO2 recovery plant is cooled against closed loop cooling water

in the Compressor Inlet Cooler (40-HX-1). Condensate is separated from the

CO2 gas in the Compressor Inlet KO Drum (40-VE-1). The CO2 is compressed in

the first two stages of the CO2 Compressor (40-CM-1). After exiting the first

stage of compression, the CO2 is cooled against closed loop cooling water in the

Compressor Intercooler (40-HX-2). Condensate is separated from the CO2 in the

Compressor KO Drum #1 (40-VE-2) before it enters the second compressor

stage. The condensate from the intercooler KO drum is combined with the

condensate from the inlet KO drum and sent to water treatment.

The CO2 leaving the second stage of the compressor is cooled against closed

loop cooling water in the Compressor Aftercooler (40-HX-3). The cooled,

compressed CO2 passes through the Compressor KO Drum #2 (40-VE-3) before

it is sent to the CO2 Drying Package (40-ME-1).

The CO2 Drying Package is a vendor-supplied glycol type, drying unit. The CO2

is contacted with glycol in a counter-current, packed bed column (Glycol

Contactor). The glycol absorbs the water vapor and then rejects the water vapor

to the atmosphere in the steam heated Glycol Regenerator. The glycol is cooled

with closed loop cooling water before it is re-circulated to the Glycol Contactor.



The dried CO2 from the drying package is sent to the third stage of the CO2

Compressor where it is compressed to the pressure required for liquefaction.

The dried and compressed CO2 is then liquefied in the CO2 Condenser

(40-HX-4). Seawater is used as the cooling medium for the condenser. The

liquid CO2 flows into the Liquid CO2 Surge Drum (40-VE-4) which feeds the CO2

Product Pumps (40-PU-1). Inerts are vented from the surge drum through the

Blow-Off System.

Steam and Power Generation

Refer to Process Flow Diagram, Steam and Power Generation - Base Case

(drawing number 592200-005) for the following discussion. This drawing depicts

the gas turbine, heat recovery steam generator (HRSG), steam turbine, and

boiler feedwater (BFW) circulation. The combined cycle unit consists of two

General Electric MS 9001FA gas turbines, a reheat steam turbine, two reheat

HRSGs, condenser, deaerator, condensate pumps, and BFW supply pumps.

Atmospheric air is filtered and introduced into the Gas Turbine (50-GT-1) air

compressor. A portion of the air is extracted from the compressor and sent to the

syn gas generation plant as oxidant to the CAPO. Steam injected from the

steam turbine is injected into the combustor of the gas turbine for NOx control.

The hot combustion gases are expanded in the gas turbine expander which

drives the compressor and the generator (see Table 8-1 at the end of section 8.2

for the gas turbine electric power output for the Base Case).

The exhaust gas exits the turbine and is ducted to the HRSG (50-HR-1). The

exhaust gas supplies heat to the steam system via the HRSG. The arrangement

of heat recovery coils in the HRSG in the direction of gas flow are as follows:

• Superheater and reheater in parallel

• HP evaporator

• High temperature BFW economizer

• MP evaporator

• Low temperature BFW economizers



• LP evaporator/integral deaerator

• Condensate heater.

The HP BFW is preheated in a low temperature economizer and a high

temperature economizer. A portion of the HP BFW is sent to the Syn Gas Cooler

in the Syn Gas Generation Plant where HP steam is produced. The remaining

HP BFW is sent to the HP evaporator in the HRSG. The HP steam from the syn

gas plant is combined with the HP steam produced in the HRSG and is

superheated. The superheated HP steam enters the HP section of the Steam

Turbine Generator (50-ST-1). A portion of the HP steam is extracted from the

HP section of the steam turbine and sent to the CAPO reactor in the syn gas

plant.

A portion of the steam exiting the HP section of the steam turbine is sent to the

gas turbine for NOx control and the remaining portion is sent to the HRSG and

reheated before it is returned to the steam turbine. Low pressure steam is

extracted from the steam turbine and desuperheated for use in the CO2 recovery

plant amine solvent reboiler. It is assumed that the plant will be operated mostly

under base load conditions. However, during part load operation, when the

steam cycle pressure “float” down to lower values, HP steam would be required

for gas turbine injection for NOx control.

The exhaust from the surface condenser is condensed in the seawater cooled

surface condenser (50-HX-1). The polished process condensate from the syn

gas plant is injected into the surface condenser. The vacuum Condensate Pump

(50-PU-2A/B) extracts the condensate from the surface condenser. The

condensate is sent through the syn gas plant for preheating and then to the

HRSGs for further preheating before it returns to the Deaerator (50-DA-1). A

portion of the boiler feedwater is sent from the deaerator to the LP steam coil in

the HRSG to produce LP steam for deaeration of the BFW.



The MP BFW is preheated in the low temperature economizer and is sent to the

MP Evaporator in the HRSG, which produces steam for regenerating the glycol in

the CO2 drying unit.

General Facilities

The general facilities required for a grass roots facility are divided:

• Demineralizer/condensate polisher

• Raw water storage

• Nitrogen

• Instrument air

• Fuel gas supply

• Closed loop cooling water supply

• Seawater cooling water supply

• Fire water

• Flare

• Roads, paving, lighting, fencing and buildings

• Auxiliary boiler

• Miscellaneous material handling

• Sanitary wastewater treatment

• In-plant electrical distribution

• Communications

• Uninterruptible power supply

• Gas turbine generator step-up transformers

• Steam turbine generator step-up transformers

• Emergency generator

• Continuous emissions monitoring (for HRSGs)

• Distributed Control System

• Chemical fire protection

• Interconnecting pipeway

• Process analyzers

• Switchyard.



5.0 IMPROVED CATALYTIC AIR PARTIAL OXIDATION PROCESS DESIGN

The Improved Catalytic Air Partial Oxidation Process as depicted in Figure 3 - 6

has similar configuration as the Base Case. The major difference is as follows:

A portion of the natural gas bypasses the syn gas generation plant and combines

with the fuel gas down stream of the carbon dioxide removal unit. The CO2

capture level in the amine unit is such that the overall reduction in the plant CO2

emissions is maintained at the 85 percent level.

Refer to the Base Case process descriptions for the following

• Oxidant Supply

• CO2 Recovery

• CO2 Compression, Drying and Liquefaction Plant

• Steam and Power Generation, and

• General Facilities.

Syn Gas Generation

Refer to Process Flow Diagram, Syn Gas Generation Plant Improved CAPO

Case (drawing number 592200-006) for the following discussion. This drawing

depicts the natural gas expansion, syn gas generation, and syn gas cooling

steps.

Natural gas enters the plant from the pipeline at 52 barA. The natural gas is

preheated against boiler feedwater in the Natural Gas Heater (20-HX-1) before


electric power.



About 8 percent of the expanded natural gas by-passes the Syn Gas Generation

Plant. This natural gas is mixed with the fuel gas from the CO2 removal plant and

preheated against boiler feedwater in the Fuel Gas Preheater (20-HX-7). The

fuel gas mixture is further heated against hot syn gas in the Fuel Gas Heater (20-

HX-5) and is then sent to the gas turbines.


compressed in the Motor driven Recycle Hydrogen Compressor (20-CM/MO-1)

and is mixed with the remaining expanded natural gas. The natural




Purifier (20-RR-2).



effluent from the CAPO reactor in the CAPO Feed Preheater (20-HX-3).


heated natural gas/hydrogen/steam mixture enter the top of the CAPO (20-RR-1)

and flow down through the catalyst bed where the methane is reacted with steam

to form H2, CO and CO2. The oxidation reaction is exothermic and the reactor is

adiabatic so that the effluent from the reactor at 895 oC is much hotter than the

feed which is at 514 oC. As described above the feed to the reactor is heated by

cooling the effluent. Additional heat is recovered from the CAPO effluent by

producing high pressure (HP) steam in the HP Steam Generator (20-HX-4). The

steam is superheated in the HRSGs before it is sent to the steam turbine.











above. The remaining syn gas from the High Temperature CO Converter is used

to preheat the fuel gas from the CO2 Recovery Plant in the Fuel Gas Heater (20-

HX-5). The preheated fuel gas is then sent to the Combined Cycle Plant. This

recovery of heat into the topping cycle increases the thermal efficiency of the

plant.



reduced. The CO concentration is reduced from approximately 8 volume percent

to less than 0.5 volume percent in the two converters.












6.0 IMPROVED CATALYTIC AIR PARTIAL OXIDATION WITH GAS HEATEDREFORMER PROCESS DESIGN

The Improved Catalytic Air Partial Oxidation Process with Gas Heated Reformer

as depicted in Figure 3 - 7 has similar configuration as the Improved CAPO Case

with the addition of a gas heated reformer downstream of the catalytic partial

oxidation reactor.

Refer to the Base Case process descriptions for the following

• Oxidant Supply

• CO2 Recovery

• CO2 Compression, Drying and Liquefaction Plant

• Steam and Power Generation, and

• General Facilities.

Syn Gas Generation

Refer to Process Flow Diagram, Syn Gas Generation Plant Improved CAPO +

GHR Case (drawing number 592200-007) for the following discussion. This

drawing depicts the natural gas expansion, syn gas generation, and syn gas

cooling steps.

Natural gas enters the plant from the pipeline at about 52 barA. The natural gas

is preheated against boiler feedwater in the Natural Gas Heater (20-HX-1) before


electric power.

About 8 percent of the expanded natural gas by-passes the Syn Gas Generation

Plant. This natural gas is mixed with the fuel gas from the CO2 plant and

preheated against boiler feedwater in the Fuel Gas Preheater (20-HX-7). The

fuel gas mixture is further heated against hot syn gas in the Fuel Gas Heater (20-

HX-5) and is then sent to the gas turbines.




compressed in the Motor driven Recycle Hydrogen Compressor (20-CM/MO-1)

and is mixed with the remaining expanded natural gas. The natural




Purifier (20-PR-2).



effluent from the Advanced Gas Heated Reformer (GHR) (20-RR-2) in the CAPO

Feed Preheater (20-HX-3).

The fuel/steam mixture then passes through the GHR (20-RR-2) where a portion

of the natural gas is steam reformed into CO, CO2 and H2. The exhaust from the

CAPO reactor provides the heat for the reformer.


reformed natural gas/hydrogen/steam mixture from the GHR enter the top of the

CAPO (20-RR-1) and flow down through the catalyst bed where the remaining

methane is converted. The reactions occurring within the CAPO are exothermic

while those in the GHR are endothermic on a net basis. The effluent after

providing heat for the GHR is at 518 oC while the feed to the GHR is at 411 oC.

Note that the difference in the inlet and outlet temperatures is much lower than

that in the cases with only the CAPO.

As described above the feed to the reactor system is heated by cooling the

effluent from the CAPO+GHR system. Additional heat is recovered by producing

high pressure (HP) steam in the HP Steam Generator (20-HX-4). The steam is

then superheated in the HRSGs before it is supplied to the steam turbine.



The catalyst in the GHR is similar to that in the CAPO which is a nickel based

shaped catalyst (pelleted with holes) with approximate size of 1 cm diameter and

2 cm length.









above. The remaining syn gas from the High Temperature CO Converter is used

to preheat the fuel gas from the CO2 Recovery Plant in the Fuel Gas Heater (20-

HX-5). The preheated fuel gas is then sent to the Combined Cycle Plant. This

recovery of heat into the topping cycle increases the thermal efficiency of the

plant.



reduced. The CO concentration is reduced from approximately 10 volume

percent to less than 1 volume percent in the two converters.














7.0 COST ESTIMATION

7.1 Plant Cost Estimate

The criteria and methodology for estimating the plant installed costs are

presented in the following. Next the technical and financial assessment

criteria used in calculating the levelized cost of electricity is presented.

7.1.1 Estimating Criteria and Methodology

Total Plant Installed Cost

The total plant installed cost is the cost required for the EPC contractor to

design and construct the plant. The plant-installed cost estimate has

been developed using a combination of equipment-factored and unit

capacity-factored estimating techniques with major equipment items

priced in-house.

Unit capacity-factored estimating is based on multiplying the cost of a unit

for which the direct construction costs are known by the ratio of the new

unit's capacity to the capacity of the known unit. Capacity ratios are

adjusted by an exponent chosen on the basis of the unit type. The costs

are adjusted for design differences, location and time frame. This

technique was used to estimate the costs of all the process units and

most of the general facilities for all cases.

Equipment-factored estimates are based on the capacity of individual

items of machinery and equipment. When developing equipment-

factored cost estimates, the cost of each equipment item (vessels,

compressors, turbines, pumps, exchangers, etc.) is determined by using

in-house pricing, data from previous projects or estimated by Fluor Daniel

proprietary estimating cost model programs. The costs for the bulk



materials (concrete, piping, electrical cable, etc.) and field labor man-

hours required for their installation are then factored, based on an

appropriate equipment parameter (duty, size, weight, etc.), to determine

the total direct construction cost. The factors for bulk materials and labor

man-hours are based on the design requirements of each equipment item

and Fluor Daniel's historical experience. The equipment factored

technique was used to estimate the costs of all the process units

(reformer oxidant supply, syn gas generation, CO2 removal, CO2

compression/drying/liquefaction) for the CAPO reactor and improved

CAPO + GHR reactor cases and the combined cycle power block for all

cases.

Pricing of major equipment was obtained in-house for combustion turbine

generators, heat recovery steam generators, steam turbine generator,

etc.

The balance of the equipment was primarily priced from recent Fluor

Daniel project equipment cost data by adjusting each reference

equipment price to the capacity, metallurgy, etc. required for this study.

A few of the general facilities that were not estimated by capacity

factoring were included as cost allowances.

All capital cost estimates reflect instantaneous mid 1999 dollars and are

comprised of the following components:

• Direct Field Material Costs;

• Direct Field Labor Costs (subcontract basis);

• Subcontract Supply and Erect Costs;

• Vendor Representative Costs;

• Ocean Freight/Marine Insurance Costs;

• Heavy Haul/Heavy Lift Costs;

• Indirect Field Costs;

• Home Office Costs;



• Initial Catalysts and Chemicals;

• Contingency;

• Fee.

The sum of these components is defined as the total plant installed cost.

The capital cost estimate does not include owner's costs, which are

estimated separately. The components of the capital cost estimate are

discussed individually in the following paragraphs.

Direct Field Material Costs

Direct field material costs are the costs of the permanent physical plant

facilities and include the following elements:

00 Account Excavation and Civil

Includes roads, asphalt paving and clean structural fill.

10 Account Concrete and Fireproofing

Includes foundations, concrete structures, retaining walls,

floor slabs, concrete area paving and concrete fireproofing.

20 Account Structural Steel

Includes steel structures, pipe racks, handrails, ladders,

stairs, platforms and miscellaneous supports.

30 Account Buildings

Includes plant buildings including framing, walls, exterior

cladding and roofing, HVAC, electrical lighting and power,

interior fixtures and finishes.

40 Account Equipment

Includes tanks, vessels, combustion turbines, steam

turbines, compressors, heat exchangers, heat recovery

steam generators (HRSGs), pumps, material processing



equipment, material handling equipment, and

miscellaneous equipment (e.g. filters, strainers, etc.)

50 Account Piping

Includes aboveground and underground pipe, fittings,

valves, flanges, gaskets, shoes and guides, specialty

items, and non-destructive examination.

60 Account Electrical

Includes major electrical equipment (transformers,

breakers, switches, bus bars, etc.), power wire and cable,

conduit, cable tray, push-button stations, welding and

power receptacles, lighting, grounding, and instrument wire

and cable.

70 Account Instrumentation

Includes process instruments, analyzers, process

connections, control house consoles, distributed control

systems, and instrument mountings.

83 Account Paint

Includes paint and coatings for plant equipment, piping,

and structures.

85 Account Insulation

Includes insulation for plant equipment and piping

The direct field material costs are based on worldwide procurement of

material. We have assumed that there will be some material purchased

outside of the Netherlands such as rotating equipment, specialty

equipment, DCS, etc. For that material, charges have been included for

inland freight to the shipping port of export, and inland freight from the

port of entry to the northeast Netherlands coastal project site. Ocean



freight costs for this imported material are covered separately as

discussed below. Freight to job site is included in the balance of material

purchased within the Netherlands.

Subcontract Labor Costs

The subcontract labor costs were derived by multiplying estimates of field

labor man-hours by a composite "all-in" wage rate. The direct field labor

man-hours were estimated according to Fluor Daniel's U.S. Gulf Coast

standard base. These man-hours are then applied to the Northeast

Netherlands coastal site location using essentially the same labor

productivity level as the U.S. Gulf Coast. The site labor productivity

includes such items as:

• manpower availability, turnover, and skill level;

• congestion, access, and available equipment laydown areas;

• weather;

• work week schedule;

• type of project;

• special safety conditions/considerations;

• work methodology and techniques/rework;

• degree of direct supervision;

• labor/management relationships;

• material delivery and shortages.

The subcontract wage rate is an average wage rate for an appropriate

mix of construction crafts. It includes salaries or wages, payroll burdens,

travel/living allowance, construction equipment and tools, field (foreman)

supervision, and subcontractor overhead and profit. The wage rate was

developed from recent project cost data supplied by Fluor Daniel Haarlem

office. An adequate supply of craft manpower is assumed to be available

in the area to support this construction project.



Subcontract Supply and Erect Costs

Subcontract supply and erect costs include equipment and materials

furnished by the local subcontractors such as field fabricated tanks,

seawater cooling system, pond liners, switchyard, etc. These subcontract

costs also include all installation labor and indirect costs, overhead and

profit of the subcontractors.

Heavy Haul/Heavy Lift Equipment

Costs for heavy haul and heavy lift equipment and cranes to move heavy

machinery and equipment are included as a factor of total equipment

cost.

Vendor Representatives

Vendor representative costs are included in the estimate as a factor of

total equipment cost.

Ocean Freight/Marine Insurance

Ocean freight/marine insurance costs are included as a factor of total

imported material cost.

Heavy Haul/Heavy Lift Equipment

Costs for heavy haul and heavy lift equipment and cranes to move heavy

machinery and equipment are included as an allowance.

Indirect Field Costs

Indirect field costs were factored from direct field costs based on Fluor

Daniel's construction experience with projects of similar magnitude.

These costs cover the following areas:

• Construction Management - Total cost of the construction supervisory

staff provided by the management contractor. Included are salaries,

allowances, burdens, benefits, overhead, and international expenses.



• Other Indirects/Services - Costs for the managing contractor's

temporary field offices, maintenance of the temporary construction

facilities, office services, office equipment and supplies, warehouse

and laydown area, guard and medical facility and services, rigging,

heavy haul road and road maintenance, and cranes.

Home Office Costs

Home office costs were factored from total field costs based on Fluor

Daniel's historical experience with adjustments made to reflect this

specific type of plant. These costs cover the following areas:

• Engineering, design and procurement work-hours, and labor costs.

• Office expenses such as computer costs, reproduction and

communication costs, and travel.

• Office burdens, benefits, and overhead costs and fee.

Initial Catalysts and Chemicals

Costs for initial catalysts and chemicals include all initial fill materials

calculated on the basis of unit price per unit volume or unit price per unit

weight of each type of catalyst or chemical times total volume or total

weight of catalyst or chemical contents inside each applicable plant

pressure vessel or storage tank (i.e., reactors, tubular reformer, guard

beds, filters, chemical/solvent storage tanks, etc.).

Contingency

Contingency is a special monetary provision in the estimate to cover

uncertainties or unforeseeable elements of time/cost within the scope of

the project. Experience indicates that contingency will be spent, but it

cannot be identified at this stage where it will be spent. Costs associated

with the following items are included in contingency:

• Material cost changes (other than scope changes)

• Labor rate changes



• Labor productivity changes

• Design changes (other than scope changes)

• Errors and omissions

• Schedule slippage

• Estimate inaccuracies.

Contingency excludes the following items:

• Changes in government regulations

• Scope changes

• Major design changes

• Catastrophic events

• Extreme weather

• Unknown field conditions

• Labor strikes.

Contingency is added to limit the probability that the final plant costs will

exceed the estimated cost plus the contingency. A contingency

allowance of 7 percent for the combined cycle power block (area 50) and

25 percent for the balance of plant has been added to the base estimate

for each case based on Fluor Daniel's experience with this type of project

at this stage of development. This results in a composite 19 percent total

plant contingency allowance. Note that IEA GHG typically uses a 10

percent contingency but in this case, given the state of development, a

higher level of contingency was deemed appropriate.

Contractor’s Fee

Fee is included as a cost allowance.

Estimate Accuracy

Estimate accuracy is + or - 30%.



Estimate Exclusions:

• Value added tax (VAT)

• Import duties

• Excise taxes

• Owner’s contingency

• Hazardous waste and underground or above ground obstruction

removal

• Construction camp and catering costs

• Environmental offsets or mitigation

• Environmental permits

• Startup costs

• Working capital

• Operating/capital spare parts

• Transmission line costs outside the plant battery limits

• Natural gas pipeline/CO2 product pipeline costs to/from plant battery

limits

• Land, rights of way, surveys and fees.

• Soil compaction/dewatering, unusual foundation requirements, piling,

etc.

• Field insurance

• Licensor costs, royalties and associated fees paid to licensors

• Project development – legal, financial, etc., consultants

• Plant mobile equipment/office furniture/laboratory equipment/shop

equipment

• Works outside of plot limits

• Site clearing & site preparation of “future areas”

• Escalation beyond mid 1999

• Delays in issuance of work permits

• Approvals for moving of heavy loads on local roads

• Scope changes

• Performance tests



• Owner/operator costs

• Interest during construction

7.1.2 Engineering, Procurement, and Construction Schedule

Figure 7-1 represents the engineering, procurement and construction

(EPC) schedule typical for all plant cases cost estimated. It is assumed

that each case would have essentially the same time duration for each

major element as shown on this schedule. Note that this schedule is

based on the current delivery time for the gas turbines reflecting the

current market demand for these engines. If a delivery time for the gas

turbines from previous years is used instead, the construction schedule

could take 36 months instead of the 44-month period shown here.

7.2 Technical and Financial Assessment Criteria

This section contains a general list of technical and financial factors for

appraisal of the cases derived mostly from criteria supplied by IEA GHG

programme.

Plant Life

Design life is used as a basis for economic appraisal that is assumed to

be 25 years.

Load Factor

Load factor, which is defined as the achieved output as a percentage of

rated/nameplate capacity, is assumed to be 90 percent for all the cases.

Cost of Debt

Since money is required during design, construction and commissioning

i.e. before any returns on sales are achieved, all capital requirements are



Figure 7 - 1



treated as debt at the same discount rate used to derive capital charges.

No allowance for grants, cheap loans etc. is made.

Capital Charges and Inflation

Discounted cash flow calculations are expressed at a discount rate of 10

percent and, to illustrate sensitivity, at 5 percent. All annual expenditures

are assumed to be incurred at the end of the year. Inflation assumptions

are not made. No allowance is made for escalation of fuel, labor, or other

costs relative to each other.

Currency

The results of the study are expressed in USD applicable to a specific

year (mid 1999). (Converting USD costs to a local currency equivalent

involves more than using the current exchange rate and members of the

IEA GHG programme will need to take their own views on appropriate

rates.)

Commissioning and Working Capital

Commissioning is defined as the period between the construction period

and the start of the 1st year of operation. Working capital includes raw

materials in store, catalysts, chemicals etc. A total of five months of

commissioning with a three-month commissioning period extending

beyond the mechanical completion is allowed for all plant. Sufficient

storage for 15 days operation at rated capacity is allowed for raw

materials, products, and consumables (except for natural gas and other

gaseous fuels in which case provision is made for an alternative supply of

fuel). No allowance is made for receipts from sales in this period.



Decommissioning

Decommissioning is cost associated with final shut down of the plant,

long term provisions and “making good” the Site. This is included to

facilitate comparison with technologies where decommissioning can be a

significant proportion of project cost. However, for this study, the

decommissioning costs were not included. Note that these plants may be

expected to operate well beyond the 25-year life assumed for economic

analysis purposes with proper operating and maintenance practices.

Location

The standard site for IEA GHG studies is on the NE coast of The

Netherlands; this appears to give costs that are in the middle of the range

for OECD member countries. A green field site with no special civil works

implications is assumed and the plant is assumed to be on the North East

coast of The Netherlands. Adequate plant and facilities to make the plant

self sufficient in site services are included in the investment costs. A cost

of 5 percent of the installed plant cost (overnight construction) is assumed

to cover land purchase, surveys, general site preparation etc.

Taxation and Insurance

Since the treatment of these items differ markedly from country to

country, a simple treatment is used which can be readily adapted to suit

the circumstances of individual members consisting of an allowance of 1

percent of the installed plant cost (overnight construction) to cover

specific services e.g. local rates. Taxation on profits is not included in the

assessments. Insurance is taken as 1 percent per annum of the installed

plant cost (overnight construction)



Miscellaneous Fees

This fee covers process/patent fees, fees for agents or consultants, legal

and planning costs etc. A total of 2 percent of installed plant cost

(overnight construction) is used.

Maintenance

Maintenance includes routine, breakdown and any major refurbishment

activities. Routine and breakdown maintenance are allowed for at 2

percent per annum of installed plant cost (overnight build) for these plants

which handle essentially gases and liquids (except for catalysts during

charging and change-out and some solid chemicals).

Operating Labor

The cost of maintenance labor is assumed to be covered by item above,

i.e. 2 percent per annum of installed plant cost. An average operating

labor rate of USD 27.4 per man-hour is assumed. Operating labor is

assumed to work 1960 hour/year in a four shift pattern. An allowance of

20 percent of the operating labor direct costs is made to cover

supervision. A further 60 percent of direct labor costs are allowed to

cover administration and general overheads (i.e. total cost = direct cost +

0.2 cost + 0.6 cost)

Fuels and Raw Materials

The cost of natural gas delivered by pipeline to site is assumed to be

USD 2/GJ (LHV). Cost of power is calculated for a range of fuel prices.

Fresh water cost is assumed to be USD 0.19 per cubic meter.



8.0 DISCUSSION OF RESULTS

8.1 Screening Analyses

A number of potential process improvements were considered in outline

and an assessment of their potential impact was made based on

estimated overall plant thermal efficiency and/or engineering judgement,

as discussed in detail in Section 3.

The results indicated that:

• Non-catalytic O2 blown partial oxidation process showed no benefits

over CAPO.

• Increasing operating pressure of the CAPO + GHR configuration

reduces the overall thermal efficiency of the plant quite significantly

because of the increased steam demand of the reformers in order to

avoid metal dusting in the GHR.

• Consequently, the use of a physical solvent process such as Selexol

for carbon dioxide capture may not be advantageously incorporated.

• Addition of the GHR increases the thermal efficiency of the plant

slightly, the significant gains made in increasing the thermal efficiency

of the syn gas generation plant being offset mostly by the increased

demand for steam by the gas turbines for NOx control.

Performance improvements may be realized by incorporating the

following design options:

• By-passing a portion of the natural gas around the syn gas generation

plant and combining it with the fuel gas down stream of the carbon

dioxide removal unit while increasing the carbon dioxide capture level

in the amine unit such that the overall reduction in the plant carbon

dioxide emissions is maintained at the 85 percent level. Another

option of “by-passing” the natural gas would be to increase the level of

unconverted methane in the reformer effluent, i.e., have a significant

“methane-slip” and not by-pass around the reactors. This scheme



would however, increase the volume of gas to be treated down stream

of the reformers, decrease the partial pressure of the reactants in the

shift reactors, decrease the partial pressure of the CO2 in the CO2

removal unit, and increase the amount of chemical energy conversion

to heat in the reformer reactor(s) which would be carried out from the

reactor(s) by the methane.

• Supplying the fuel gas to the gas turbines at a higher temperature. A

temperature of 399 oC was selected. General Electric offers three

types of fuel systems, one for fuel gas temperatures up to 233 oC, one

for gas temperatures up to 399 oC and finally one for temperatures up

to 538 oC (which tends to be a very expensive system). The 399 oC

temperature was achieved by first heating the fuel gas against boiler

feedwater and then preheating it against the high temperature shift

reactor effluent gas.

8.2 Performance Summary - Improved Cases

The thermal and environmental performance summaries for the improved

cases which were configured based on the findings of the screening

phase of this study (“Improved CAPO Case” and “Improved CAPO+GHR

Case”) as well as the Base Case are presented in Tables 8-1 and 8-2.

As can be seen from the results, the Improved CAPO+GHR Case has the

highest thermal efficiency. The net output of the plant for this case is the

lowest because a large amount of steam is removed from the steam

turbine and utilized in the gas turbine for NOx control. The gas turbines in

this case are fully loaded since the air extraction rate (for the CAPO) is

lower than the other cases. Using a GHR (2-stage reactor system) gives

significant gains in the efficiency of syn gas generation but the overall

gain is not as significant because more steam has to be added to the gas

turbine to keep the NOx levels down. The result is an efficiently produced

syn gas which is richer in hydrogen than for the single CAPO reactor



system but its benefits are not able to be registered because of

combustion considerations in the turbine.

While comparing the efficiency of the cases developed in this study to

those published elsewhere, it should be noted the NOx emission for these

cases was set at the very low value of 10 ppmv (dry, 15 percent O2

basis) which demanded a large amount of steam injection into the gas

turbines. Steam injection into the gas turbine degrades the thermal

efficiency of the combined cycle significantly.

The BASF activated MDEA process was used for CO2 removal and

generation of a waste stream consisting of degraded amine solution is not

expected from this plant.

The power generated by the natural gas expander is listed separately in

Table 8-1, which may be deleted for cases where the natural gas supply

pressure is lower and an expander may not be incorporated in the design

of the plant. Note that the impact of deleting the natural gas heater on the

steam cycle and consequently the output of the steam turbine will be

insignificant since the heat required by the heater was supplied by heat

recovered from the HRSG down stream of its low pressure steam

generator.

Table 8-3 summarizes the major heat rejections from the plant for each of

the three cases. Note that the latent heat of the water vapor is included in

the energy content of the gas streams leaving the plant. As can be seen

from the table, the HRSG stack gas for the Improved CAPO+GHR Case

is very high because of the larger amount of steam added to the gas

turbine.



Table 8-1Thermal Performance Comparison

Case Designation Base Case Improved CAPO Case Improved CAPO + GHRCase

ReactorType

CAPO CAPO CAPO+GHR

Reactor Pressure,barA

30 30 30

Oxidant Air Air Air

CO2 Removal Amine Amine Amine

Natural GasBypass, %

0 7.84 7.73

Fuel Input, MW(LHV)

1740 1708 1537

GT Power, MW 544 566 572

ST Power, MW 289 272 174

Expander Power,MW 1.5 1.3 1.2

Gross Power, MW 835 839 748

Auxiliary Power,MW 90 90 70

Net Power, MW 745 749 677

ThermalEfficiency, % LHV 42.8 43.8 44.1



Table 8-2Environmental Performance


ReactorType

CAPO CAPO CAPO+GHR

Fuel Input, MW(HHV)

1922 1888 1698

Fuel Input, MW(LHV)

1740 1708 1537


ThermalEfficiency, % HHV 38.8 39.7 39.9

ThermalEfficiency, % LHV 42.8 43.8 44.1

CO2 Emission,g/kW

72.7 71.6 71.6

NOx Emissions,ppmv (dry, 15%O2 basis)Mg/Nm3 (dry, 15%O2 basis)

10

25

10

25

10

25



Table 8-3Plant Major Heat Rejections

(Basis for Stream Energy Content: 15.9 oC and Liquid Water)


ReactorType

CAPO CAPO CAPO+GHR

Fuel Input, MW(HHV)

1922 1888 1698

Fuel Input, MW(LHV)

1740 1708 1537


ThermalEfficiency, % HHV 38.8 39.7 39.9HRSG Stack Gas

Temp, oC 127 89 100

HRSG Stack GasHeat, MW 562 542 597

SurfaceCondenser HeatRejection, MW 339 301 176

CO2 SeparationUnit HeatRejections, MW 144 159 137

CO2 Compression& LiquefactionHeat Exchangers,MW 67 66 59

Total Major HeatRejections, MW 1112 1068 969

Total Major HeatRejections, % ofFuel HHV 57.9 56.6 57.1

Total Major HeatRejections, % ofFuel LHV 63.9 62.5 63.0



8.3 Cost Estimates

The plant cost estimate summaries for the Improved CAPO Case and

Improved CAPO+GHR Case as well as the Base Case are presented in

Table 8-4 broken down by major plant sections. Note that cost of the

natural gas heater and expander are shown separately which may be

deleted for cases where the natural gas supply pressure is lower and an

expander may not be incorporated in the design of the plant.

The costs on a USD/kW basis are lowest for the Improved CAPO Case.

The Improved CAPO+GHR Case uses less air for the reformers and

produces less volume of syn gas (because of the reduced concentration

of N2 in the gas). The resulting decrease in cost of equipment down

stream and upstream of the reformers for the Improved CAPO+GHR

Case was more than offset by the cost of the GHR. In the two CAPO only

cases, the feed/effluent interchanger down stream of the CAPO is

expensive on a unit of heat transferred basis but because of the very

large temperature differences prevailing between the streams exchanging

the heat, the required heat transfer area is quite low resulting in a lower

absolute cost for this unit when compared to the GHR.

The levelized cost of electricity calculated for the three cases is presented

for various fuel costs and discounted cash flow (DCF) rates in Figures 8-1

through 8-4. Two sets of curves have been generated, one for the 44

month construction schedule and the other for the 36 month construction

schedule. The cost of electricity calculated for the three cases are similar

with the Improved CAPO case showing slightly lower cost than the other

cases in the entire range of fuel costs. The improved CAPO Case’s

slightly lower thermal efficiency is more than offset by its lower plant cost

when compared to the Improved CAPO+GHR Case.



CRE GROUP, LTDNortheast NetherlandsIE Greenhouse Gas Study

ImprovedImproved CAPO + GHR

Area Area Descriptions CAPO Reactor CAPO Reactor Reactors

10 Reformer Oxidant Supply 53,947 49,852 41,688

20 Syngas Generation -- Feed Heater & Expander 4,295 2,450 3,782

20 Syngas Generation -- Balance of Unit 103,206 98,808 110,437

21 CO2 Removal 131,985 140,468 122,216

40 CO2 Compression / Drying / Liquefaction 52,562 51,755 48,329

50 Combined Cycle Power Block 240,553 241,634 228,208

60 General Facilities 209,994 207,929 193,833

Total Installed Cost 796,542 792,896 748,493

Net Power Output, MW = 745.34 748.69 677.35

$/KW = 1,069 1,059 1,105

Table 8 - 4` FLUO R DAN IEL

Rev. 0 29-July-99

(Thousand US Dollars)Instantaneous Mid 1999 Costs

Estimate Summary Comparison

With CO2 RemovalReformed Syngas Fired Combined Cycle Plant



Figure 8-1 (44 Month Construction Schedule, DCF = 10%)

COE Versus Natural Gas Cost

42.5

45

47.5

50

52.5

1.6 1.8 2 2.2 2.4 2.6 2.8

Natural Gas Cost, USD/GJ

CO

E, c

ents

/kW

h

Base Case CAPO + GHR CAPO



Figure 8-2(44 Month Construction Schedule, DCF = 5%)


33353739414345

1.6 1.8 2 2.2 2.4 2.6 2.8







41

43

45

47

49

51

53

1.6 1.8 2 2.2 2.4 2.6 2.8


CO

E, c

ents

/kW

h






33

35

37

39

41

43

1.6 1.8 2 2.2 2.4 2.6 2.8


CO

E, c

ents

/kW

h


CRE Group, Ltd. FLElectricity Production and CO2 Capture via Partial Oxidation of Natural Gas C


8.4 Conclusions and Recommendations

Increasing operating pressure of the reforming process reduces the overall thermal efficiency of

the plant quite significantly because of the increased steam demand of the reformers.

Consequently, the use of a physical solvent process such as Selexol for carbon dioxide capture

may not be advantageously incorporated. Use of a turbo-expander to recover power from the

syn gas could not be taken advantage of either.

The steam demand of the reformer is increased in order to avoid metal dusting. In order to take

full advantage of operating the reformer at higher pressure, developmental work directed at

reducing the steam consumption of the reformer or developing materials that are not susceptible

to metal dusting or a combination of the two would be required.

Using a GHR results in an efficiently produced syn gas which is richer in hydrogen than a single

CAPO reactor system but its benefits are not able to be realized because of combustion

considerations in the turbine which demands significant amounts of steam to keep the NOx

levels down. Use of a Selective Catalytic Reduction unit (SCR) which may cost USD 2 ½ to 3

million per gas turbine, installed down stream of the gas turbine for NOx control should be

investigated in the future. A disadvantage of the SCR system is that it requires the handling and

storage of ammonia. Catalysts are being developed that do not require ammonia such as those

installed within the combustor of the gas turbine. Steam injection may still be required but to a

limited extent. Research in the area of gas turbine combustors aimed at minimizing the steam

addition while burning H2 rich gas would be useful.

Cycles such as the Humid Air Turbine (HAT) or the Inter-cooled Steam Injected (ISTIG) cycle

should be investigated in addition to the combined cycle because these cycles consist of or

depend on introducing water vapor or steam into the combustor of the gas turbine and NOx

control is a natural outcome of these cycles without penalizing the thermal efficiency of the

cycle.

A configuration consisting of reforming the natural gas (steam-methane reforming) utilizing at

least a portion of the gas turbine exhaust to supply the combustion air to the reformer should

also be investigated. Another concept that should be considered consists of utilizing a catalyst

CRE Group, Ltd. FLElectricity Production and CO2 Capture via Partial Oxidation of Natural Gas C


doped ceramic membrane reformer which is under development. The membrane separates the

H2 as it is formed such that the H2 from the reaction mixture is constantly removed which

results in maximizing the conversion of the methane at the lower temperatures prevalent in the

gas turbine exhaust.

The results of this study point the direction towards another configuration that may be worthy of

evaluation, a high pressure CAPO (without the GHR) with an expander to recover the potential

energy of the high pressure syn gas. Note that without the GHR, the steam demand of the

reformer does not increase as the pressure is increased. The hot gas leaving the low

temperature shift unit could be fed to the expander to recover power as well as cool the gas. In

this configuration, a chemical solvent such as the amine wash would be suitable for CO2

removal.

Increasing the gas turbine fuel gas temperature to 399 oC showed a significant increase in the

thermal efficiency of the plant. However, the steam addition to the gas turbine had to be also

increased in order to control the NOx at the 10 ppmv (dry, 15 percent O2 basis). Thus, the full

benefit of utilizing a higher fuel temperature was not realized. A trade-off in order to optimize

the fuel temperature is another area for further investigation.



APPENDIX 1



TABLE A-1STREAM DATA

BASE CASESTREAM NUMBER AND DESCRIPTION

1 2 3 4 5 6 7Component Extraction Air Preheated Air Natural Gas Expanded HP Steam Syn Gas Syn Gas From

From Gas To CAPO From Natural Gas From CCP From CAPO High Temp.Turbine Pipeline CO Convertermol% mol% mol% mol% mol% mol% mol%

H2 0.00 0.00 0.00 0.00 0.00 27.49 33.61H2O 0.99 0.25 0.00 0.00 100.00 26.27 20.15N2 77.34 77.92 0.40 0.40 0.00 31.59 31.59O2 20.74 20.89 0.00 0.00 0.00 0.00 0.00

CO 0.00 0.00 0.00 0.00 0.00 7.96 1.84CO2 0.00 0.00 1.80 1.80 0.00 5.80 11.92CH4 0.00 0.00 83.90 83.90 0.00 0.51 0.51C2H6 0.00 0.00 9.20 9.20 0.00 0.00 0.00C3H8 0.00 0.00 3.30 3.30 0.00 0.00 0.00NC4H10 0.00 0.00 1.40 1.40 0.00 0.00 0.00

Ar 0.93 0.94 0.00 0.00 0.00 0.38 0.38

Total kg/sec 186.96 186.08 37.18 37.18 82.72 307.98 307.98bar abs 13.8 33.18 52 34.68 34.1 30.69 28.5°C 404 378 15 106 381 895 400Mole. Wt. 28.85 28.93 19.40 19.40 18.02 19.15 19.15LHV (dry), kJ/kg 0 0 46,870 46,870 0 6,450 5,832Enthalpy, GJ/h 281 252 6,925 6,958 928 8,507 7,365



TABLE A-1, CONTINUEDSTREAM DATA

BASE CASESTREAM NUMBER AND DESCRIPTION

8 9 10 11 12 13Component Syn Gas From Syn Gas to Fuel Gas CO2 From Preheated CO2

Low Temp. CO2 From CO2 Recovery Plant Fuel Gas to ProductCO Converter Recovery Plant Recovery Plant CCP

mol% mol% mol% mol% mol% mol%H2 35.23 43.19 50.72 0.14 50.72 0.15H2O 18.53 0.12 0.14 8.98 0.14 0.00N2 31.59 38.72 45.48 0.05 45.48 0.05O2 0.00 0.00 0.00 0.00 0.00 0.00

CO 0.22 0.27 0.32 0.00 0.32 0.00CO2 13.54 16.60 2.05 90.83 2.05 99.80CH4 0.51 0.63 0.74 0.00 0.74 0.00C2H6 0.00 0.00 0.00 0.00 0.00 0.00C3H8 0.00 0.00 0.00 0.00 0.00 0.00NC4H10 0.00 0.00 0.00 0.00 0.00 0.00

Ar 0.38 0.47 0.55 0.00 0.55 0.00

Total kg/sec 307.98 254.6 168.84 89.34 166.84 85.86bar abs 27.34 26.45 26.3 1.1 26.12 90°C 225 25 45 45 281 21Mole. Wt. 19.15 19.41 15.12 41.61 15.12 43.94LHV (dry), kJ/kg 5,683 5,685 8,576 8 8,576 8Enthalpy, GJ/h 6,989 6,147 6,171 42 6,383 -74




IMPROVED CAPO CASESTREAM NUMBER AND DESCRIPTION

1 2 3 4 4A 4B 5 6Component Extraction Air Preheated Air Natural Gas Expanded Natural Gas Natural Gas HP Steam Syn Gas

From Gas To CAPO From Natural Gas to CAPO to Gas Turbine From CCP From CAPOTurbine Pipelinemol% mol% mol% mol% mol% mol% mol% mol%

H2 0.00 0.00 0.00 0.00 0.00 0.00 0.00 27.60H2O 0.99 0.24 0.00 0.00 0.00 0.00 100.00 26.26N2 77.34 77.92 0.40 0.40 0.40 0.40 0.00 31.45O2 20.74 20.90 0.00 0.00 0.00 0.00 0.00 0.00

CO 0.00 0.00 0.00 0.00 0.00 0.00 0.00 7.96CO2 0.00 0.00 1.80 1.80 1.80 1.80 0.00 5.78CH4 0.00 0.00 83.90 83.90 83.90 83.90 0.00 0.57C2H6 0.00 0.00 9.20 9.20 9.20 9.20 0.00 0.00C3H8 0.00 0.00 3.30 3.30 3.30 3.30 0.00 0.00NC4H10 0.00 0.00 1.40 1.40 1.40 1.40 0.00 0.00

Ar 0.93 0.94 0.00 0.00 0.00 0.00 0.00 0.38

Total kg/sec 167.56 166.78 36.5 36.5 30.3 2.86 74.76 277.18bar abs 16.05 35.3 52 36.37 36.37 36.37 35.77 32.2°C 406 381 15 108.9 108.9 108.9 387 895Mole. Wt. 28.85 28.93 19.40 19.40 19.40 19.40 18.02 19.11LHV (dry), kJ/kg 0 0 46,870 46,870 46,870 46,870 0 6,518Enthalpy, GJ/h 175 228 6,798 6,832 5,671 535 841 7,716




IMPROVED CAPO CASESTREAM NUMBER AND DESCRIPTION

7 8 9 10 11 12 13Component Syn Gas From Syn Gas From Syn Gas to Fuel Gas CO2 From Preheated CO2

High Temp. Low Temp. CO2 From CO2 Recovery Plant Fuel Gas to ProductCO Converter CO Converter Recovery Plant Recovery Plant CCP

mol% mol% mol% mol% mol% mol% mol%H2 33.72 35.34 43.12 51.57 0.15 50.63 0.16H2O 20.14 18.52 0.57 0.35 8.72 0.67 0.00N2 31.45 31.45 38.38 45.89 0.03 45.07 0.03O2 0.00 0.00 0.00 0.00 0.00 0.00 0.00

CO 1.85 0.22 0.27 0.32 0.00 0.32 0.00CO2 11.89 13.52 16.50 0.48 91.10 0.50 99.81CH4 0.57 0.57 0.70 0.83 0.00 2.06 0.00C2H6 0.00 0.00 0.00 0.00 0.00 0.14 0.00C3H8 0.00 0.00 0.00 0.00 0.00 0.05 0.00NC4H10 0.00 0.00 0.00 0.00 0.00 0.02 0.00

Ar 0.38 0.38 0.46 0.56 0.00 0.54 0.00

Total kg/sec 277.18 277.18 230.04 145.26 87.5 146.72 84.2bar abs 30.27 29 28.5 28 1.1 27.6 90°C 400 225 55 45 45 398.45 21Mole. Wt. 19.11 19.11 19.36 14.62 41.68 14.70 43.94LHV (dry), kJ/kg 5,896 5,743 5,745 9,085 9 9,835 9Enthalpy, GJ/h 6,688 6,346 5,636 5,611 41 6,461 -74




IMPROVED CAPO + GHR CASESTREAM NUMBER AND DESCRIPTION

1 2 3 4 4A 4B 5 6Component Extraction Air Preheated Air Natural Gas Expanded Natural Gas Natural Gas HP Steam Syn Gas

From Gas To CAPO From Natural Gas to CAPO to Gas Turbine From CCP From CAPOTurbine Pipelinemol% mol% mol% mol% mol% mol% mol% mol%

H2 0.00 0.00 0.00 0.00 0.00 0.00 0.00 33.36H2O 0.99 0.25 0.00 0.00 0.00 0.00 100.00 25.00N2 77.34 77.92 0.40 0.40 0.40 0.40 0.00 26.00O2 20.74 20.89 0.00 0.00 0.00 0.00 0.00 0.00

CO 0.00 0.00 0.00 0.00 0.00 0.00 0.00 10.16CO2 0.00 0.00 1.80 1.80 1.80 1.80 0.00 4.82CH4 0.00 0.00 83.90 83.90 83.90 83.90 0.00 0.35C2H6 0.00 0.00 9.20 9.20 9.20 9.20 0.00 0.00C3H8 0.00 0.00 3.30 3.30 3.30 3.30 0.00 0.00NC4H10 0.00 0.00 1.40 1.40 1.40 1.40 0.00 0.00

Ar 0.93 0.94 0.00 0.00 0.00 0.00 0.00 0.31

Total kg/sec 115.94 115.4 32.84 32.84 30.3 2.54 67.2 214.9bar abs 15.64 35.3 52 36.37 36.37 36.37 35.77 30.26°C 401 376 15 108.9 108.9 108.9 387 517.97Mole. Wt. 28.85 28.93 19.40 19.40 19.40 19.40 18.02 17.61LHV (dry), kJ/kg 0 0 46,870 46,870 46,870 46,870 0 8,557Enthalpy, GJ/h 173 156 6,116 6,146 5,671 475 756 6,784




IMPROVED CAPO + GHR CASESTREAM NUMBER AND DESCRIPTION

7 8 9 10 11 12 13Component Syn Gas From Syn Gas From Syn Gas to Fuel Gas CO2 From Preheated CO2

High Temp. Low Temp. CO2 From CO2 Recovery Plant Fuel Gas to ProductCO Converter CO Converter Recovery Plant Recovery Plant CCP

mol% mol% mol% mol% mol% mol% mol%H2 40.39 42.98 50.48 60.70 0.15 59.57 0.16H2O 17.97 15.38 0.61 0.38 8.72 0.72 0.00N2 26.00 26.00 30.54 36.73 0.03 36.04 0.03O2 0.00 0.00 0.00 0.00 0.00 0.00 0.00

CO 3.13 0.54 0.64 0.76 0.00 0.75 0.00CO2 11.85 14.44 16.95 0.50 91.10 0.52 99.81CH4 0.35 0.35 0.41 0.49 0.00 1.76 0.00C2H6 0.00 0.00 0.00 0.00 0.00 0.14 0.00C3H8 0.00 0.00 0.00 0.00 0.00 0.05 0.00NC4H10 0.00 0.00 0.00 0.00 0.00 0.02 0.00

Ar 0.31 0.31 0.37 0.44 0.00 0.43 0.00

Total kg/sec 214.9 214.9 182.22 106.02 78.46 107.12 75.5bar abs 28.42 27.26 26.75 26.28 1.1 25.91 90°C 425 250 55 45 45 400 21Mole. Wt. 17.61 17.61 17.54 12.27 41.68 12.40 43.94LHV (dry), kJ/kg 7,604 7,294 7,295 12,524 9 13,353 9Enthalpy, GJ/h 6,532 6,231 5,659 5,637 36 6,345 -65



Documents

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