CO2 Capture by Partial Oxidation of Natural

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

Syngas reactor Shift CO to storage conversion 2

CO2 capture

H2O, N 2 to stack H2-rich fuel gas

Steam cycle

natural gas air gas turbine

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”.

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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 Reference CAPO case ‘Improved’ process with gas -heated (mole%) 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 6.5 8.6 8.6 13.4 (MJ/kg) (LHV)

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 7 the turbine’s exhaust gases for NOx control. 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 0.04 0.057 gas) 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) whic h 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 49.6 Wheeler Studies for others by Foster Wheeler13 Foster 46.3 to 46.9 Wheeler

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.

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

Electricity Production and CO2 Capture via Partial Oxidation of Natural Gas

Presented to

CRE Group, LTD.

by ` Fluor Daniel, Inc One Fluor Daniel Drive Aliso Viejo, CA 92698

Principal Investigators:

Ashok Rao David Francuz Jeffrey Scherffius Edward West

November 1999

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DISCLAIMER

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

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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

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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.

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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

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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.

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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

D:\Power_generation\A P O Study\REPORT.doc 4 CRE Group, Ltd. FLUOR DANIEL, INC Electricity Production and CO2 Capture via Partial Oxidation of Natural Gas Contract 04592200 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,

D:\Power_generation\A P O Study\REPORT.doc 5 CRE Group, Ltd. FLUOR DANIEL, INC Electricity Production and CO2 Capture via Partial Oxidation of Natural Gas Contract 04592200 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.

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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.

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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.

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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.

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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 balance Air (60% R.H.) 15 °C calculations Seawater Supply (maximum) 12 °C

Seawater Supply (average) 8 °C

Seawater Return (max. allow) 19 °C

Surface Condenser Pressure 2.0 cm HgA

2.2 Products

2.2.1 Electric Power Nominal Combined 800 Mwe cycle (50Hz) output

2.2.2 Recovered CO2 Target recovery 85 %

Delivery state Liquid

Delivery point Plant limits

Composition CO2 >99.5 mole % specifications H2 <0.5 mole %

Pressure 90 Bar abs.

Temperature ≤30 °C

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2.3 Raw Materials

2.3.1 Natural Gas Composition 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 Water Pressure 11 BarA

2.4 Plant Configuration Plant Heat Rejection Once through seawater cooling

Plant Water Makeup Municipal water

Gas Turbine Two trains (General Electric MS 9001FA)

Steam Turbine Single train

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

Particulate matter < 25 mg/Nm3

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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

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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),

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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

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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

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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

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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

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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.

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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 3 the equilibrium constraint of the reforming reaction: p H2 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.

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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

D:\Power_generation\A P O Study\REPORT.doc 20 COMPRESSED EXTRACTION AIR AIR SUPPLY

PREHEATED FUEL GAS

CATALYTIC PARTIAL CO CO2 GAS EXHAUST OXIDATION CONVERSION RECOVERY HRSGS TURBINES (CAPO) (SHIFT) (AMINE) 30 BarA

CO2 CO2 LIQUEFACTION PRODUCT

NOx CONTROL STEAM STEAM STEAM TURBINE

NATURAL GAS NATURAL GAS RECYCLE GAS HYDROGEN EXPANSION & RICH GAS FROM PURIFICATION COMPRESSION PIPELINE 52 BarA

DESIGNED BY D. FRANCUZ FIGURE 1 FLU O R DANIEL CHECKED BY PROCESS ENGR. APP. DATE OVERALL BLOCK FLOW DIAGRAM DRAWN REV. DATE REVISION DESCRIPTION APP. NO. CHECK ` LEAD ENGR. / SPECIALIST APP. DATE BASE (CAPO) CASE A. RAO PROJECT APP. DATE JULY 1999

CLIENT APP. DATE SCALE DRAWING NUMBER REV. DWG. NO. REFERENCE DWG. NO. DWG. NO. REFERENCE DWG. NO. NONE FIGURE 1 A COMPRESSED EXTRACTION AIR AIR SUPPLY

PREHEATED FUEL GAS

CATALYTIC PARTIAL OXIDATION CO CO2 GAS EXHAUST (CAPO) & GAS CONVERSION RECOVERY HRSGS TURBINES HEATED (SHIFT) (AMINE) REFORMER (GHR) 30 BarA

CO2 CO2 LIQUEFACTION PRODUCT

NOx CONTROL STEAM STEAM STEAM TURBINE

NATURAL GAS NATURAL GAS RECYCLE GAS HYDROGEN EXPANSION & RICH GAS FROM PURIFICATION COMPRESSION PIPELINE 52 BarA

DESIGNED BY D. FRANCUZ FIGURE 2 FLU O R DANIEL CHECKED BY PROCESS ENGR. APP. DATE OVERALL BLOCK FLOW DIAGRAM DRAWN REV. DATE REVISION DESCRIPTION APP. NO. CHECK ` LEAD ENGR. / SPECIALIST APP. DATE CAPO +GHR CASE A. RAO PROJECT APP. DATE JULY 1999

CLIENT APP. DATE SCALE DRAWING NUMBER REV. DWG. NO. REFERENCE DWG. NO. DWG. NO. REFERENCE DWG. NO. NONE FIGURE 2 A CRE Group, Ltd. FLUOR DANIEL, INC Electricity Production and CO2 Capture via Partial Oxidation of Natural Gas Contract 04592200

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.

D:\Power_generation\A P O Study\REPORT.doc 23 COMPRESSED EXTRACTION AIR AIR SUPPLY

PREHEATED FUEL GAS

CATALYTIC PARTIAL OXIDATION CO CO2 FUEL GAS GAS EXHAUST (CAPO) & GAS CONVERSION RECOVERY PREHEAT AND HRSGS TURBINES HEATED (SHIFT) (SELEXOL) EXPANSION REFORMER (GHR) 46 BarA

CO2 CO2 LIQUEFACTION PRODUCT

NOx CONTROL STEAM STEAM STEAM TURBINE

NATURAL GAS HYDROGEN NATURAL GAS RECYCLE GAS RICH GAS PURIFICATION FROM COMPRESSION PIPELINE 52 BarA

DESIGNED BY D. FRANCUZ FIGURE 3 FLU O R DANIEL CHECKED BY PROCESS ENGR. APP. DATE OVERALL BLOCK FLOW DIAGRAM DRAWN REV. DATE REVISION DESCRIPTION APP. NO. CHECK ` LEAD ENGR. / SPECIALIST APP. DATE HIGH PRESSURE CAPO +GHR CASE A. RAO PROJECT APP. DATE JULY 1999

CLIENT APP. DATE SCALE DRAWING NUMBER REV. DWG. NO. REFERENCE DWG. NO. DWG. NO. REFERENCE DWG. NO. NONE FIGURE 3 A CRE Group, Ltd. FLUOR DANIEL, INC Electricity Production and CO2 Capture via Partial Oxidation of Natural Gas Contract 04592200

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

D:\Power_generation\A P O Study\REPORT.doc 25 HIGH PRESSURE NITROGEN FOR NOx CONTROL ELEVATED PRESSURE AIR EXTRACTION AIR SEPARATION PREHEATED FUEL GAS 95% O2

PARTIAL OXIDATION CO CO2 FUEL GAS GAS EXHAUST (NON- CONVERSION RECOVERY PREHEAT AND HRSGS TURBINES CATALYTIC) (SHIFT) (SELEXOL) EXPANSION 83 BarA

CO2 CO2 LIQUEFACTION PRODUCT

ADDITIONAL NOx CONTROL STEAM STEAM STEAM TURBINE

NATURAL NATURAL GAS GAS HYDROGEN COMPRESSION RECYCLE GAS RICH GAS AND FROM COMPRESSION PIPELINE PURIFICATION 52 BarA

DESIGNED BY D. FRANCUZ FIGURE 4 FLU O R DANIEL CHECKED BY PROCESS ENGR. APP. DATE OVERALL BLOCK FLOW DIAGRAM DRAWN REV. DATE REVISION DESCRIPTION APP. NO. CHECK ` LEAD ENGR. / SPECIALIST APP. DATE HIGH PRESSURE POX CASE A. RAO PROJECT APP. DATE JULY 1999

CLIENT APP. DATE SCALE DRAWING NUMBER REV. DWG. NO. REFERENCE DWG. NO. DWG. NO. REFERENCE DWG. NO. NONE FIGURE 4 A CRE Group, Ltd. FLUOR DANIEL, INC Electricity Production and CO2 Capture via Partial Oxidation of Natural Gas Contract 04592200

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).

D:\Power_generation\A P O Study\REPORT.doc 27 COMPRESSED EXTRACTION AIR AIR SUPPLY

PREHEATED FUEL GAS

CATALYTIC PARTIAL OXIDATION CO CO2 GAS EXHAUST (CAPO) & GAS CONVERSION RECOVERY HRSGS TURBINES HEATED (SHIFT) (AMINE) REFORMER (GHR) 46 BarA

SYN GAS CO2 CO2 EXPANSION LIQUEFACTION PRODUCT

NOx CONTROL STEAM STEAM STEAM TURBINE

NATURAL GAS HYDROGEN NATURAL GAS RECYCLE GAS RICH GAS PURIFICATION FROM COMPRESSION PIPELINE 52 BarA

DESIGNED BY D. FRANCUZ FIGURE 5 FLU O R DANIEL CHECKED BY OVERALL BLOCK FLOW DIAGRAM PROCESS ENGR. APP. DATE DRAWN REV. DATE REVISION DESCRIPTION APP. HIGH PRESSURE CAPO +GHR CASE WITH NO. CHECK ` LEAD ENGR. / SPECIALIST APP. DATE SHIFTED GAS EXPANSION A. RAO PROJECT APP. DATE JULY 1999

CLIENT APP. DATE SCALE DRAWING NUMBER REV. DWG. NO. REFERENCE DWG. NO. DWG. NO. REFERENCE DWG. NO. NONE FIGURE 5 A CRE Group, Ltd. FLUOR DANIEL, INC Electricity Production and CO2 Capture via Partial Oxidation of Natural Gas Contract 04592200

• 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.

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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.

D:\Power_generation\A P O Study\REPORT.doc 30 COMPRESSED EXTRACTION AIR AIR SUPPLY HOT FUEL GAS 399 °C

FUEL GAS PREHEAT

CATALYTIC PARTIAL CO CO2 OXIDATION GAS EXHAUST CONVERSION RECOVERY HRSGS (CAPO) TURBINES (SHIFT) (AMINE) REFORMER (GHR) 30 BarA

CO2 CO2 LIQUEFACTION PRODUCT NOx CONTROL STEAM STEAM STEAM TURBINE

HYDROGEN NATURAL GAS RECYCLE GAS RICH GAS PURIFICATION COMPRESSION

NATURAL GAS BY-PASS

NATURAL GAS NATURAL GAS PREHEAT AND FROM EXPANSION PIPELINE DESIGNED BY D. FRANCUZ FIGURE 6 52 BarA FLU O R DANIEL CHECKED BY PROCESS ENGR. APP. DATE OVERALL BLOCK FLOW DIAGRAM DRAWN REV. DATE REVISION DESCRIPTION APP. NO. CHECK ` LEAD ENGR. / SPECIALIST APP. DATE IMPROVED CAPO CASE A. RAO PROJECT APP. DATE JULY 1999

CLIENT APP. DATE SCALE DRAWING NUMBER REV. DWG. NO. REFERENCE DWG. NO. DWG. NO. REFERENCE DWG. NO. NONE FIGURE 6 A COMPRESSED EXTRACTION AIR AIR SUPPLY HOT FUEL GAS 399 °C

FUEL GAS PREHEAT

CATALYTIC PARTIAL OXIDATION CO CO2 GAS EXHAUST (CAPO) & GAS CONVERSION RECOVERY HRSGS TURBINES HEATED (SHIFT) (AMINE) REFORMER (GHR) 30 BarA

CO2 CO2 LIQUEFACTION PRODUCT NOx CONTROL STEAM STEAM STEAM TURBINE

HYDROGEN NATURAL GAS RECYCLE GAS RICH GAS PURIFICATION COMPRESSION

NATURAL GAS BY-PASS

NATURAL GAS NATURAL GAS PREHEAT AND FROM EXPANSION PIPELINE DESIGNED BY D. FRANCUZ FIGURE 7 52 BarA FLU O R DANIEL CHECKED BY PROCESS ENGR. APP. DATE OVERALL BLOCK FLOW DIAGRAM DRAWN REV. DATE REVISION DESCRIPTION APP. NO. CHECK ` LEAD ENGR. / SPECIALIST APP. DATE IMPROVED CAPO +GHR CASE A. RAO PROJECT APP. DATE JULY 1999

CLIENT APP. DATE SCALE DRAWING NUMBER REV. DWG. NO. REFERENCE DWG. NO. DWG. NO. REFERENCE DWG. NO. NONE FIGURE 7 A CRE Group, Ltd. FLUOR DANIEL, INC Electricity Production and CO2 Capture via Partial Oxidation of Natural Gas Contract 04592200

Table 3-1 Thermal Performance Estimates

Case Base 2 3 4 5 Improved Improved Designation Case CAPO CAPO + (Case 1) Case GHR Case Reactor CAPO CAPO+ CAPO+ POX CAPO+ CAPO CAPO+ Type GHR GHR GHR GHR Reactor 30 30 46 83 46 30 30 Pressure, barA Oxidant Air Air Air 95% O2 Air Air Air CO2 Amine Amine Selexol Selexol Amine Amine Amine Removal Natural Gas 0 0 0 0 0 7.84 7.73 Bypass, % Fuel Input, 1740 1565 1533 1708 1571 1708 1537 MW (LHV) GT Power, 544 568 570 549 572 566 572 MW ST Power, 289 179 95 257 169 272 174 MW Expander 1.5 1.2 11 16 18 1.3 1.2 Power, MW Gross 835 748 676 822 759 839 747 Power, MW Aux. Power, 90 75 129 127 83 90 70 MW Net Power, 745 673 547 695 676 749 677 MW Thermal Efficiency, 42.8 43.1 35.7 40.7 43.0 43.8 44.1 % LHV

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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.

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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

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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.

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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.

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

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

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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.

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

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

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• 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.

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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.

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

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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.

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 remaining expanded natural gas. 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 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.

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

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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 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.

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.

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6.0 IMPROVED CATALYTIC AIR PARTIAL OXIDATION WITH GAS HEATED REFORMER 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 the pressure is reduced in the Natural Gas Expander (20-EX-1), which produces 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.

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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 remaining expanded natural gas. 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-PR-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 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.

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

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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.

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-4) 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 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.

The syn gas from the two preheaters is recombined and sent to the Low Temperature CO Converter (20-RR-5) where the CO concentration is further reduced. The CO concentration is reduced from approximately 10 volume percent to less than 1 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.

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

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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

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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;

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• 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

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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

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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.

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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.

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• 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

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• 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%.

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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

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• 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

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Figure 7 - 1

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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.

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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)

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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.

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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

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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

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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.

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Table 8-1 Thermal Performance Comparison

Case Designation Base Case Improved CAPO Case Improved CAPO + GHR Case Reactor CAPO CAPO CAPO+ Type GHR Reactor Pressure, 30 30 30 barA Oxidant Air Air Air CO2 Removal Amine Amine Amine Natural Gas 0 7.84 7.73 Bypass, % Fuel Input, MW 1740 1708 1537 (LHV) 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 Thermal Efficiency, % LHV 42.8 43.8 44.1

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Table 8-2 Environmental Performance

Case Designation Base Case Improved CAPO Case Improved CAPO + GHR Case Reactor CAPO CAPO CAPO+ Type GHR Fuel Input, MW 1922 1888 1698 (HHV) Fuel Input, MW 1740 1708 1537 (LHV) Net Power, MW 745 749 677 Thermal Efficiency, % HHV 38.8 39.7 39.9 Thermal Efficiency, % LHV 42.8 43.8 44.1 CO2 Emission, 72.7 71.6 71.6 g/kW NOx Emissions, ppmv (dry, 15% O2 basis) 10 10 10 Mg/Nm3 (dry, 15% O2 basis) 25 25 25

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Table 8-3 Plant Major Heat Rejections (Basis for Stream Energy Content: 15.9 oC and Liquid Water)

Case Designation Base Case Improved CAPO Case Improved CAPO + GHR Case Reactor CAPO CAPO CAPO+ Type GHR Fuel Input, MW 1922 1888 1698 (HHV) Fuel Input, MW 1740 1708 1537 (LHV) Net Power, MW 745 749 677 Thermal Efficiency, % HHV 38.8 39.7 39.9 HRSG Stack Gas Temp, oC 127 89 100 HRSG Stack Gas Heat, MW 562 542 597 Surface Condenser Heat Rejection, MW 339 301 176 CO2 Separation Unit Heat Rejections, MW 144 159 137 CO2 Compression & Liquefaction Heat Exchangers, MW 67 66 59 Total Major Heat Rejections, MW 1112 1068 969 Total Major Heat Rejections, % of Fuel HHV 57.9 56.6 57.1 Total Major Heat Rejections, % of Fuel LHV 63.9 62.5 63.0

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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.

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Table 8 - 4 CRE GROUP, LTD ` FLU O R DANIEL Northeast Netherlands IE Greenhouse Gas Study Rev. 0 29-July-99 Reformed Syngas Fired Combined Cycle Plant With CO2 Removal

Estimate Summary Comparison Instantaneous Mid 1999 Costs (Thousand US Dollars)

Improved Improved 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

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Figure 8-1 (44 Month Construction Schedule, DCF = 10%)

COE Versus Natural Gas Cost 52.5

50

47.5

45 COE, cents/kWh 42.5 1.6 1.8 2 2.2 2.4 2.6 2.8 Natural Gas Cost, USD/GJ Base Case CAPO + GHR CAPO

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Figure 8-2 (44 Month Construction Schedule, DCF = 5%)

COE Versus Natural Gas Cost

45 43 41 39 37 35 33 1.6 1.8 2 2.2 2.4 2.6 2.8 Natural Gas Cost, USD/GJ

Base Case CAPO + GHR CAPO

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Figure 8-3 (36 Month Construction Schedule, DCF = 10%)

COE Versus Natural Gas Cost 53

51

49

47

45 COE, cents/kWh 43

41 1.6 1.8 2 2.2 2.4 2.6 2.8 Natural Gas Cost, USD/GJ Base Case CAPO + GHR CAPO

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Figure 8-4 (36 Month Construction Schedule, DCF = 5%)

COE Versus Natural Gas Cost 43

41

39

37

COE, cents/kWh 35

33 1.6 1.8 2 2.2 2.4 2.6 2.8 Natural Gas Cost, USD/GJ Base Case CAPO + GHR CAPO

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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

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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.

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APPENDIX 1

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TABLE A-1 STREAM DATA BASE CASE STREAM NUMBER AND DESCRIPTION 1234567 Component 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 Converter mol% mol% mol% mol% mol% mol% mol%

H2 0.00 0.00 0.00 0.00 0.00 27.49 33.61

H2O 0.99 0.25 0.00 0.00 100.00 26.27 20.15

N2 77.34 77.92 0.40 0.40 0.00 31.59 31.59

O2 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 1.84

CO2 0.00 0.00 1.80 1.80 0.00 5.80 11.92

CH4 0.00 0.00 83.90 83.90 0.00 0.51 0.51

C2H6 0.00 0.00 9.20 9.20 0.00 0.00 0.00

C3H8 0.00 0.00 3.30 3.30 0.00 0.00 0.00

NC4H10 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 82.72 307.98 307.98 bar abs 13.8 33.18 52 34.68 34.1 30.69 28.5 °C 404 378 15 106 381 895 400 Mole. Wt. 28.85 28.93 19.40 19.40 18.02 19.15 19.15 LHV (dry), kJ/kg 0 0 46,870 46,870 0 6,450 5,832 Enthalpy, GJ/h 281 252 6,925 6,958 928 8,507 7,365

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TABLE A-1, CONTINUED STREAM DATA BASE CASE STREAM NUMBER AND DESCRIPTION 8 9 10 11 12 13 Component Syn Gas From Syn Gas to Fuel Gas CO2 From Preheated CO2 Low Temp. CO2 From CO2 Recovery Plant Fuel Gas to Product CO Converter Recovery Plant Recovery Plant CCP mol% mol% mol% mol% mol% mol%

H2 35.23 43.19 50.72 0.14 50.72 0.15

H2O 18.53 0.12 0.14 8.98 0.14 0.00

N2 31.59 38.72 45.48 0.05 45.48 0.05

O2 0.00 0.00 0.00 0.00 0.00 0.00 CO 0.22 0.27 0.32 0.00 0.32 0.00

CO2 13.54 16.60 2.05 90.83 2.05 99.80

CH4 0.51 0.63 0.74 0.00 0.74 0.00

C2H6 0.00 0.00 0.00 0.00 0.00 0.00

C3H8 0.00 0.00 0.00 0.00 0.00 0.00

NC4H10 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 85.86 bar abs 27.34 26.45 26.3 1.1 26.12 90 °C 225 25 45 45 281 21 Mole. Wt. 19.15 19.41 15.12 41.61 15.12 43.94 LHV (dry), kJ/kg 5,683 5,685 8,576 8 8,576 8 Enthalpy, GJ/h 6,989 6,147 6,171 42 6,383 -74

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TABLE A-2 STREAM DATA IMPROVED CAPO CASE STREAM NUMBER AND DESCRIPTION 12344A4B56 Component 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 CAPO Turbine Pipeline mol% mol% mol% mol% mol% mol% mol% mol%

H2 0.00 0.00 0.00 0.00 0.00 0.00 0.00 27.60

H2O 0.99 0.24 0.00 0.00 0.00 0.00 100.00 26.26

N2 77.34 77.92 0.40 0.40 0.40 0.40 0.00 31.45

O2 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 7.96

CO2 0.00 0.00 1.80 1.80 1.80 1.80 0.00 5.78

CH4 0.00 0.00 83.90 83.90 83.90 83.90 0.00 0.57

C2H6 0.00 0.00 9.20 9.20 9.20 9.20 0.00 0.00

C3H8 0.00 0.00 3.30 3.30 3.30 3.30 0.00 0.00

NC4H10 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 74.76 277.18 bar abs 16.05 35.3 52 36.37 36.37 36.37 35.77 32.2 °C 406 381 15 108.9 108.9 108.9 387 895 Mole. Wt. 28.85 28.93 19.40 19.40 19.40 19.40 18.02 19.11 LHV (dry), kJ/kg 0 0 46,870 46,870 46,870 46,870 0 6,518 Enthalpy, GJ/h 175 228 6,798 6,832 5,671 535 841 7,716

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TABLE A-2, CONTINUED STREAM DATA IMPROVED CAPO CASE STREAM NUMBER AND DESCRIPTION 7 8 9 10 11 12 13 Component 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 Product CO 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 0.16

H2O 20.14 18.52 0.57 0.35 8.72 0.67 0.00

N2 31.45 31.45 38.38 45.89 0.03 45.07 0.03

O2 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 0.00

CO2 11.89 13.52 16.50 0.48 91.10 0.50 99.81

CH4 0.57 0.57 0.70 0.83 0.00 2.06 0.00

C2H6 0.00 0.00 0.00 0.00 0.00 0.14 0.00

C3H8 0.00 0.00 0.00 0.00 0.00 0.05 0.00

NC4H10 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 84.2 bar abs 30.27 29 28.5 28 1.1 27.6 90 °C 400 225 55 45 45 398.45 21 Mole. Wt. 19.11 19.11 19.36 14.62 41.68 14.70 43.94 LHV (dry), kJ/kg 5,896 5,743 5,745 9,085 9 9,835 9 Enthalpy, GJ/h 6,688 6,346 5,636 5,611 41 6,461 -74

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TABLE A-3 STREAM DATA IMPROVED CAPO + GHR CASE STREAM NUMBER AND DESCRIPTION 12344A4B56 Component 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 CAPO Turbine Pipeline mol% mol% mol% mol% mol% mol% mol% mol%

H2 0.00 0.00 0.00 0.00 0.00 0.00 0.00 33.36

H2O 0.99 0.25 0.00 0.00 0.00 0.00 100.00 25.00

N2 77.34 77.92 0.40 0.40 0.40 0.40 0.00 26.00

O2 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 10.16

CO2 0.00 0.00 1.80 1.80 1.80 1.80 0.00 4.82

CH4 0.00 0.00 83.90 83.90 83.90 83.90 0.00 0.35

C2H6 0.00 0.00 9.20 9.20 9.20 9.20 0.00 0.00

C3H8 0.00 0.00 3.30 3.30 3.30 3.30 0.00 0.00

NC4H10 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 67.2 214.9 bar abs 15.64 35.3 52 36.37 36.37 36.37 35.77 30.26 °C 401 376 15 108.9 108.9 108.9 387 517.97 Mole. Wt. 28.85 28.93 19.40 19.40 19.40 19.40 18.02 17.61 LHV (dry), kJ/kg 0 0 46,870 46,870 46,870 46,870 0 8,557 Enthalpy, GJ/h 173 156 6,116 6,146 5,671 475 756 6,784

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TABLE A-3, CONTINUED STREAM DATA IMPROVED CAPO + GHR CASE STREAM NUMBER AND DESCRIPTION 7 8 9 10 11 12 13 Component 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 Product CO 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 0.16

H2O 17.97 15.38 0.61 0.38 8.72 0.72 0.00

N2 26.00 26.00 30.54 36.73 0.03 36.04 0.03

O2 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 0.00

CO2 11.85 14.44 16.95 0.50 91.10 0.52 99.81

CH4 0.35 0.35 0.41 0.49 0.00 1.76 0.00

C2H6 0.00 0.00 0.00 0.00 0.00 0.14 0.00

C3H8 0.00 0.00 0.00 0.00 0.00 0.05 0.00

NC4H10 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 75.5 bar abs 28.42 27.26 26.75 26.28 1.1 25.91 90 °C 425 250 55 45 45 400 21 Mole. Wt. 17.61 17.61 17.54 12.27 41.68 12.40 43.94 LHV (dry), kJ/kg 7,604 7,294 7,295 12,524 9 13,353 9 Enthalpy, GJ/h 6,532 6,231 5,659 5,637 36 6,345 -65

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