decarbonizedhydrogenandelectricityfromnaturalgasfaculty.jsd.claremont.edu/emorhardt/159/pdfs/2006/...apparently...

International Journal of Hydrogen Energy 30 (2005) 701–718www.elsevier.com/locate/ijhydene

Decarbonized hydrogen and electricity from natural gas

Stefano Consonni∗,1, Federico Viganò1

1Princeton Environmental Institute, Princeton University, Princeton, NJ 08544-1033, USA

Available online 12 September 2004

Abstract

This paper discusses configuration, attainable performances and thermodynamic features of stand-alone plants for the co-production of de-carbonized hydrogen and electricity from natural gas (NG) based on commercially available technology.

We focus on the two basic technologies currently used in large industrial applications: fired tubular reformer (FTR) andauto-thermal reformer (ATR). In both cases we assume that NG is pre-heated and humidified in a saturator providing water forthe reforming reaction; this reduces the amount of steam to be bled from the power cycle and increases electricity production.Outputs flows are made available at conditions suitable for transport via pipeline: 60 bar for pure hydrogen, 150 bar for pureCO2. To reduce hydrogen compression power requirements reforming is carried out at relatively high pressures: 25 bar forFTR, 70 bar for ATR. Reformed gas is cooled and then passed through two water–gas shift reactors to optimize heat recoveryand maximize the conversion to hydrogen. In plants with CO2 capture, shifted gas goes through an amine-based chemicalabsorption system that removes most of the CO2. Pure hydrogen is obtained by pressure swing absorption (PSA), leaving apurge gas utilized to fire the reformer (in FTR) and to boost electricity production.

For the power cycle we consider conventional steam cycles (SC) and combined cycles (CC). The scale of plants based ona CC is determined by the gas turbine. To maintain NG input within the same range (around 1200 MW), we considered aGeneral Electric 7FA for ATR, a 6FA for FTR. The scale of plants with SC is set by assuming the same NG input of thecorresponding CC plant.

Heat and mass balances are evaluated by a model accounting for the constraints posed by commercial technology, as wellas the effects of scale. Results show that, from a performance standpoint, the technologies of choice for the production ofde-carbonized hydrogen from NG are FTR with SC or ATR with CC. When operated at high steam-to-carbon ratios, the latterreach CO2 emissions chargeable to hydrogen of 10–11 kg of CO2 per GJLHV —less than 20% of NG—with an equivalentefficiency of hydrogen production in excess of 77%.� 2004 International Association for Hydrogen Energy. Published by Elsevier Ltd. All rights reserved.

Keywords:Hydrogen production; Combined cycles; CO2 capture; Efficiency; Natural gas

1. Background and scope

Concern about rising concentrations of greenhouse gasesin the atmosphere has spurred research into various ways

∗ Corresponding author. Fax: +39-02-2399-3940.E-mail address:[email protected](S. Consonni).

1 Current address: Dipartimento di Energetica, Politecnico diMilano, P.za L. Da Vinci 32, Milano 20133, Italy.

0360-3199/$30.00� 2004 International Association for Hydrogen Energy. Published by Elsevier Ltd. All rights reserved.doi:10.1016/j.ijhydene.2004.07.001

of capturing and storing CO2 before it is emitted. Due tostrong scale effects on performance and cost, this technol-ogy is applied most naturally to large, point source emittersof CO2 such as fossil fuel-based plants generating electric-ity or synthetic fuels. So far, most of the attention has beendirected to the “de-carbonization” of electricity from fossilfuels, which can be realized according to three different cap-ture strategies: “pre-combustion”[1–4], “post-combustion”[5–7], and “oxy-fuel” [8,9,40]. However, electricity produc-tion is responsible for only∼ 30% of global CO2 emissions,

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702 S. Consonni, F. Viganò / International Journal of Hydrogen Energy 30 (2005) 701–718

Nomenclature

E electric power, MWF fuel, MWH hydrogen output, MWLHVmbl flow bled from steam turbine, kg/sp pressure, barT temperature,◦C or KWLHV LHV power, MW�E specific CO2 emissions, kg/MWh�H specific CO2 emissions, kg/GJLHV�E net electric power/nat gas LHV power�H H2 LHV power/nat gas LHV power�Ex Exergy (or 2nd-Law) efficiency�∗

Ex Exergy efficiency, net of CO2 capture

Acronyms

ASU air separation unitCC combined cycleATR auto-thermal reformers

FTR fried tubular reformersGT gas turbineLHV lower heating value, MJ/kgHP, IP, LP high/intermediate/low pressureHRSG heat recovery steam generatorHT, LT high/low temperatureNG natural gasPSA pressure swing adsorptionRH, SH reheat, superheatSC steam cycleS/C steam-to-carbon ratio (molar)ST steam turbineTIT turbine inlet temperatureWGS water-gas shiftWGSR water–gas shift reactor

Subscripts

E ElectricityH Hydrogenref Reforming

and this share is expected to fall[10]; the remainder comesfrom fossil fuels used directly in distributed applicationssuch as in transportation, residences, commercial and indus-trial facilities, where CO2 capture, transport, and disposalwould be problematic and extremely costly.

In order to stabilize atmospheric CO2 at 450 ppm—an ap-parently acceptable target—the carbon emissions predictedby the “business-as-usual” IPCC IS92a scenario[11] for theperiod 1990–2100 should be cut by about 50% and, by 2100,annual emissions would have to be reduced to about 3 Gi-gatons of carbon per year[12,13]. Even if the power sectorcould be completely de-carbonized, the achievement of thisemission level would require a five-fold reduction of carbonemissions from fuels used directly[10]—a level achievableonly by a massive use of synthetic, de-carbonized fuels. De-spite its many severe handicaps—low density, large com-pression work, wide flammability limits, low ignition energy,containment problems, safety concerns, etc.—hydrogen isattractive both because it is completely de-carbonized andbecause it is most suited to fuel cells.

In a carbon-constrained world, the production of hydro-gen from fossil fuels (other than via electrolysis, nuclear en-ergy or other carbon-free processes) makes sense only if itsproduction is associated to carbon capture. This has been in-vestigated in a number of recent studies.Table 1reports thebasic technological characteristics and performance of thesystems considered in these studies. In all cases, hydrogen isthe only basic output; net electricity production is generallynegative, i.e. it must be imported to satisfy auxiliary powerconsumption. This follows from three circumstances sharedby many of the natural gas (NG)-based hydrogen plants op-erating today, where (i) electricity is less valuable than

hydrogen; (ii) export of electricity is unattractive becausethe plant is remote from consumption centers and/or thegrid has limited capacity; (iii) export of electricity is con-strained by legislation. Should hydrogen be produced on avery large scale to feed a large fractions of the transporta-tion, industrial and residential sectors, then these circum-stances are likely to vanish because (i) hydrogen will haveto compete for less profitable uses (ii) hydrogen produc-tion will take place in a relatively large number of plantsall across industrial countries, with good interconnections tothe electric grid; (iii) the liberalization of electricity marketsbeing carried out in many countries will facilitate electricityexports.

This is why in this paper we wish to clarify whether theco-production of significant amounts of electricity togetherwith hydrogen can be beneficial in terms of primary energyconsumption and/or reduced CO2 emissions, and whethersuch practice is subject to technological and/or thermody-namic constraints.

Co-production can be accomplished by a number of tech-nologies. In this paper, we limit our attention to establishedcommercial technologies: fired tubular reformers (FTR) andauto-thermal reformers (ATR) for hydrogen synthesis; iron-and copper-based water–gas shift reactors (WGSR) for hy-drogen enrichment; Rankine steam cycles (SC) and com-bined cycles (CC) for power production; pressure swing ab-sorption (PSA) for hydrogen separation; amine absorptionfor CO2 removal. Focusing on commercial technology al-lows us to utilize data and draw from the expertise devel-oped from actual industrial plants. Results can be used as areference to assess the merits of alternative feedstocks likecoal or those of advanced technologies[14–16,42].

S. Consonni, F. Viganò / International Journal of Hydrogen Energy 30 (2005) 701–718 703

Table 1Technologies and performances quoted in recent studies on the production of hydrogen from NG with FTR

Source CO2 capture technology NG input H2 output (LHV) Electric output % of LHV input CO2 captureMWLHV % of input (%)

gross net

Foster Wheeler[14] — 369.7 76.0 1.94 0.0 0.0Foster Wheeler[14] Amines 383.6 73.3 4.05 0.0 85.0Jacobs[15] Amines 361.4 69.2 n.a. −2.32 86.7Jacobs[15] Selexol� 366.6 70.2 n.a. −3.23 85.5Klett et al. [16] — 764.2 75.9 n.a. −0.79 0.0Klett et al. [16] Amines 703.4 82.5 n.a. −2.13 71.0

The values in italics in the last two rows are estimates based on the information given in[16].

2. Technologies, plant configurations and basicassumptions

The variety of options analyzed in the paper comes fromthe combination of the technologies considered for:

(a) hydrogen synthesis: FTR vs. ATR;(b) power production: SC vs. CC;(c) CO2 capture: venting vs. amine absorption.

This originates a total of eight technological options. Forall of them we have assumed the following basic features,meant to be representative of large-scale plants feeding anetwork for H2 distribution and, when applicable, CO2storage.

• With respect to industrial standards, the plant isvery large: NG input is in the range 1200–2000MWLHV , giving an hydrogen output in the range600–1300 MWLHV . This compares with the hydrogenoutput of 350–400 MWLHV of the largest FTRs builtto date and the range up to 2000 MWLHV of hydrogenapparently interesting for ATRs[17]. NG input is de-termined by imposing that the gas turbine (GT) of theplants with CC, CO2 capture and reference steam-to-carbon—1.0 for ATR, or 2.0 for FTR—runs at the designturbine inlet temperature (TIT) while fed only by purgegas; the NG input of the other cases follows as illustratedin Table 2. The selection of two different GTs—GE 7FAfor ATR, GE 6FA for FTR—allows maintaining the NGinput of both reforming technologies within comparableranges.

• The plant is stand alone, i.e. there is no steam or chemicalintegration with an adjoining process.

• Feedstock is “commercial” NG with enough sulfur andparaffins to require de-sulfurization and pre-reforming.De-sulfurization is carried out by catalytic hydro-genation and sulfur absorption over zinc-oxide platesat 380◦C; the hydrogen required for this process(10% by volume of the gas fed to the desulfurization

unit)1 is supplied by recycling a small fraction of thepure hydrogen generated by the PSA unit.

• To maximize electricity production, NG is humidified(and preheated) by a saturator. Plants based on a CCalso include another saturator to humidify and preheatthe purge gas fed to the GT.

• Ahead of the reformer, the mixture of NG and watervapor is heated to 620◦C and passed through an adi-abatic catalytic pre-reformer. By getting rid of higher-order hydrocarbons—ethane, propane, butane, etc.—thepre-reformer strongly reduces the likelihood of soot for-mation downstream of the reformer.

• The gas exiting the pre-reformer is pre-heated to 670◦Cwith combustion gases (FTR) or with reformed syngas(ATR).

• Water–gas shift (WGS) is carried out in two steps: thefirst is promoted by an iron-based catalyst in an adiabaticreactor operating between 320 and 470◦C; the second ispromoted by a copper-based catalyst in a cooled (ATR)or adiabatic (FTR) reactor operating between 200 and230◦C.

• Pure hydrogen separated by PSA is delivered at the plantgate at 60 bar.

• For the plants with CO2 capture, nearly pure liquid CO2is delivered at the plant gate at 150 bar.

For each technological option, the complex heat ex-changer network needed to recover heat from the reformedsyngas and from combustion gases has been arranged tolimit heat transfer irreversibilities (keep low�T ) whilemeeting the assumptions on temperatures, pressures,�T

and�p adopted to represent the state-of-the-art of reform-ers and power plants (see Chapter 3).

1 This value, which is appreciably higher than that typicallyadopted for natural gas (2–5%), is presumably necessary to insureno carbon formation in the pre-heaters that follow desulfurization.High hydrogen concentration in the hydrogenator saturates theunsaturated hydrocarbons more prone to cracking[18], allowinghigher pre-heating temperatures.


Table 2Criteria adopted to size plants with ATR at 950◦C and FTR at 850◦C

Steam cycle Combined cycle

Criterion adopted to NGWLHV , MW Criterion adopted to NGWLHV , MWdetermine NG input determine NG input

CO2 CO2 capture CO2 CO2 captureventing venting

ATR S/C = 1.0 Same NG input of plant 1236.4 1236.4 Purge gas generated by PSA= 1336.7 1236.4with CC and CO2 capture flow needed to fully load GE 7FA

S/C#1.0 a

FTR S/C = 2.0 Purge gas generated by PSA= 1944.7 1195.4flow needed by FTR burners+flowneeded to fully load GE 6FA

S/C> 2.0 Same NG input of plant 1195.4 1195.4 Same NG input of case with 1195.4with CC and CO2 capture S/C = 2.0b

When the same criteria are applied to plants with CC and higher reforming temperature, NG input increases to 2005.2 MW (ATR at1050◦C, S/C=0.80) and 2384.2 MW (FTR at 880◦C, S/C=2.0). The situations corresponding to the shaded cells have not been considered.

aNG input varies from 1236.4 MW, for S/C = 1.0, to 1876.3 MW, for S/C = 2.10.bSince the purge gas needed for the FTR burners increases with S/C, when S/C> 2.0 the purge gas to the GT is not enough, and a

fraction of the NG input must be used directly in the GT combustor. When S/C> 2.75 the purge gas does not even meet the demand ofthe FTR burners; in this case the GT runs 100% on NG and some NG also goes to the FTR burners.

Since thermal power and temperature profile of heat ex-changers depend on mass flow rates, the plant configura-tion does not only depend on the technologies consideredfor reforming, power production and CO2 capture, but alsoon the steam-to-carbon ratio (S/C). The plant schemes inFigs. 1–4give the configurations considered as base casesfor each technology, i.e. FTR and SC with S/C= 3.07, FTRand CC with S/C= 2.0, ATR with S/C= 1.0. Some of theresults for different S/C have been obtained for a differentarrangement of the heat exchanger sequence.

2.1. Fired tubular reformers

FTR are the dominant technology for NG steam-reforming[17–19,37]. They consist of a furnace filled withan array of super-alloy tubes containing a nickel-based cat-alyst which promotes the reforming and WGS reaction ofthe mixture of methane and water vapor flowing inside thetubes:

CH4 + H2O + 206.158 kJ/molCH4 → CO+ 3H2

CO+ H2O → CO2 + H2 + 44.447 kJ/molC

The heat required by the highly endothermic reforming re-action is provided by burners fed with purge gas left afterhydrogen extraction and some NG. Due to the large heatof reaction of steam reforming,2 hydrogen production is

2 The catalyst brings the mixture exiting the tubes close toequilibrium, and the equilibrium constant varies with temperatureaccording to vant’Hoff equation. The large heat of reaction makesthe equilibrium constant a very steep function of temperature.

very sensitive to temperature, so that it is imperative that thereacting mixture inside the tubes reach high temperatures.High pressure is detrimental, because the reaction increasesvolume flow, yet desirable because it reduces the size (andthe cost) of the equipment. In state-of-the-art FTRs the re-formed syngas exits the tubes at 22–25 bar and 820–880◦C.High temperatures, high stresses and high heat fluxes createmost challenging operating conditions for the tube mate-rial, consisting of a nickel-based super-alloy at the leadingedge of metallurgical science. Reforming conditions are ul-timately determined by the requirements on tube creep andlife, which relate to temperature and stress through corre-lations like Larson–Miller’s. Our choice of 25 bar, 850◦Cconforms to the state-of-the-art of large plants for the refin-ing industry.

Figs. 1and2 report the configuration of the plants basedon a SC and a CC, respectively. NG (at 50 bar and ambienttemperature) is heated to the desulfurization temperaturefirst in a regenerative heat exchanger fed with de-sulfurizedgas, then in a heater fed by reformed syngas. After being de-sulfurized and cooled in the regenerator,3 NG is saturated,pre-heated and mixed with steam bled from the power cycleto achieve the specified S/C; then, the mixture to be reformedis heated to 620◦C with syngas exiting the reformer, pre-reformed and finally heated to 670◦C with the exhausts ofthe FTR burners. The gas–gas heat exchanger heating themixture to be reformed to 620◦C ahead of the pre-reformersubstitutes the quench boiler typically placed at the exit ofthe reformer. This appears viable because the pre-reformer

3 The temperature at the outlet of the regenerator is set byimposing a minimum�T of 30◦C.


H8

Hydrogenationand SulfurAbsorption

CO2Absorption

K.O.Drum

PSA

~

~H2 IC

Compr.

H2 Rec.Compr.

StackAir

HPDrum

K.O.Drum

K.O.Drum

H1

H2

H3

H4

H5

H6

H7

H9

H10

H11

H12

H13

H14

H15

H16

H17

H18

NGSaturator

Pre

refo

rmer

HT

WG

SR

LT W

GS

R

Burners

~

BFW H

SteamTurbine

Dea

reat

or

H2 @60bar

NG

FTR

Steam for aminestripping

Process steam

Recycled H2 to Hydrogenation

Fig. 1. Configuration of plant with FTR and SC for S/C = 3.07.

makes soot formation unlikely. The S/C ratio is controlledby adjusting the amount of superheated vapor bled from thesteam turbine (ST) and injected ahead of the gas–gas heatermentioned above.

The combustion gases of the FTR burners exit the furnaceat 1000◦C and enter a series of heat exchangers including,besides the heater of the mixture exiting the pre-reformer,a steam superheater, an HP evaporator, an economizer andan air heater bringing combustion air to 300◦C.4

In the plants with SC, some NG (6–9% of LHV inputto FTR burners) must be fed to the furnace together withPSA purge gas to control the flame. Instead, in the plantswith CC the burners are fed solely by purge gas because theflame is controlled by purge gas bled (at 3.5 bar) from theintercooled compressor feeding the GT.

2.2. Auto-thermal reformers

ATR consist of an adiabatic vessel with a section filledwith nickel catalyst where the heat for the reforming reaction

4 Except in the scheme with SC and CO2 capture, where thecombustion air temperature is limited to 130◦C to leave enoughlow-temperature heat for the economizer. Lowering the combustionair temperature increases the amount of purge gas needed for theFTR furnace and thus the attainable S/C; in turn, lower S/C givelower �H and higher�E.

is provided by the partial oxidation of methane with oxygen

CH4 + 0.5O2 → CO+ 2H2 + 35.670 kJ/molCH4

CO+ H2O → O2 + H2 + 44.447 kJ/molC

CH4 + H2O + 206.158 kJ/molCH4 → CO+ 3H2

These very same reactions could be carried out also usingair as the oxidant. However, unless nitrogen is needed togenerate useful products (like in ammonia production) thecost increase due to the much larger flow rate across thewhole plant makes air unattractive.

Without a heat transfer surface, the operating tempera-ture is not constrained by material properties and it is typ-ically chosen to warrant adequate activity and life of thecatalyst—or the desired syngas composition, a situation typ-ically encountered in plants for urea production. Our basecase at 950◦C is within the range of current industrial ap-plications and allows operating at S/C favorable for CO2capture.5 Despite its detrimental effects on the reforming

5 Higher ATR temperatures greatly increase the amount ofsteam generated in the evaporator placed downstream of the re-former (heat exchanger H7 inFigs. 3 and 4). If this steam flowis too large, superhear/reheat (SH/RH) temperature must decreasebecause the heat recoverable for superheating and reheating is lim-ited. To prevent this reduction of SH/RH temperature—and theconsequent de-rating of the power cycle—one must decrease S/C.


H8H7


CO2Absorption

K.O.Drum

H14

PSA

~Purge ICCompr.

~

~H2 IC

Compr.

H2 Rec.Compr.

StackAir

~

GasTurbine

Air

HRSCHP

DrumSteam

Turbine Deareator

Stack

HPDrum

K.O.Drum

K.O.Drum

H1

H2

H3

H4

H5

H6

H9

H10

H11

H12

H13

H15

H16

H17

H18

H19

PurgeSaturator

NGSaturator

Pre

refo

rmer

HT

WG

SR

LT W

GS

R

Burners

H2 @60bar

Ste

am fo

r am

ine

strip

ping

Process steam

NG

FTR


Fig. 2. Configuration of plants with FTR and CC for S/C = 2.0.

reaction, operating pressure is typically high to reduce thesize and the cost of the reformer. In our case, reformingpressure is set to warrant the extraction of pure hydrogenfrom the PSA at 60 bar.

Figs. 3 and 4 report the configurations considered forSCs and CCs. The path followed by NG through the de-sulfurization unit, the saturator and the pre-reformer is sim-ilar to that of FTRs. The basic differences consist in (i) ASUand oxygen compressor; (ii) extraction of steam for the re-former at high pressure (∼ 70 bar); (iii) no fuel gas from thereformer, so that heat recovery takes place solely from re-formed syngas; (iv) cooled low-temperature (LT) WGS reac-tor, generating low pressure (LP) steam at 12.5–23 bar6 (v)two-pressure-level steam cycle. Extracting heat from the LTWGS reactor increases hydrogen production and the fractionof CO2 which can be captured; these advantages will haveto be weighed against the higher complexity and cost of theLT WGS reactor and of the two-pressure-level steam cycle.

6 For the cooled WGS reactor it is assumed that the evaporationtemperature is 10◦C lower than that of the reacting mixture. Athigh S/C, the pressure of LP steam is increased to warrant atemperature of the reacting mixture at least 10◦C above its dewpoint, to avoid condensate formation in the reactor.

2.3. Pressure swing absorption

High purity (99.99%) H2 is assumed to be removed fromthe syngas at 35◦C using PSA, a proprietary process com-monly used in syngas processing[20–22,41].

Industrial installations based on FTRs operate at about20 bar and reach H2 separation efficiencies in the range85–90%. In this paper, we have assumed 88% for all FTRplants. Maintaining the same H2 separation efficiency at thepressure of∼ 60 bar considered for ATRs requires a morecomplex (and more expensive) system arrangement. The ac-tual design and thus the separation efficiency will be deter-mined by economic considerations. To be consistent withour focus on best available technology and given the infor-mation provided by Jacobs[15] and Allam[23], for ATRswe have assumed a separation efficiency of 85%, a valuesignificantly higher than that assumed by Foster Wheeler[16] and Doctor et al.[24].

On the hydrogen side, the flow incurs a 2% pressuredrop, while the purge gas (remaining H2 along with CO,H2O and CH4) is discharged at 1.3 bar independentlyof the inlet pressure. In ATR plants, the pressure of thereformer is set to give a hydrogen pressure at the PSAexit of 60 bar, a value suited for long-range transport;in FTR plants, hydrogen is discharged from the PSA at


Pro

cess

ste

am


CO2Absorption~

~

ASU

~

Air

HRSCHPDrum

SteamTurbine Deareator

SH

K.O.Drum

NGSaturator

K.O.Drum

K.O.Drum

NG Compressor

H2 Rec. Compr.

PSA

Stack

AirLT

WG

SR

HT

WG

SR

Pre

refo

rme

r

LPDrum

O2 ICCompr.

H18

H6

H3

H8

H9-10

H1

H2

H12 H13

H14

H16

H17

H18

Burners

NGH2 @60bar

ATR

H7

H5

H4

Steam for amine stripping


Vent

Fig. 3. Configuration of plants with ATR and conventional SC for S/C = 1.0.

∼ 21 bar and is compressed to 60 bar by an intercooledcompressor.

2.4. Power plant

Given our interest in exploring the potential for powerproduction, we have considered a steam cycle with op-erating conditions typical of power plants: evaporation at110–130 bar, SH temperature 540–565◦C, RH at 30 bar,540–565◦C. These conditions are more advanced than thosetypically adopted in hydrogen plants, but certainly feasiblein large stand-alone plants like those considered here.

For the plants based on a CC we have considered two GTsconforming to the latest generation of heavy-duty, 60 Hzengines: General Electric 6FA for FTR, 7FA for ATR. Thesmaller power output of the turbine considered for FTRsmatches one of the intrinsic features of this technology,where the fraction of NG input converted to electricity(�E)

is smaller than that converted by ATRs. Considering thesame GT for both technologies would lead to comparisonswhere the NG input of FTRs is much larger than that ofATRs, a situation opposite to the experience dictated byeconomies of scale.

Table 3compares the performances quoted by GE withthose predicted by our model for operation on NG, as well

as with those for syngas-fired operation. When running onsyngas it is assumed that TIT, compression ratio and turbinereduced mass flow are the same of the NG version, wherethe last condition is met by adjusting the air flow (close InletGuide Vanes).7

2.5. CO2 separation and compression

The relatively low partial pressure of CO2 in the syngas tobe decarbonized points to chemical absorption as the tech-nology which is presumably most cost effective. Following

7 Gas turbines typically run in choked conditions, so that anincrease in fuel flow requires either a higher pressure ratio or alower air flow. Higher pressure ratio is preferable because it givesthe largest increase in power output; however, compressor stalllimits its maximum allowable increase to few percent. When airflow is decreased by closing the Inlet Guide Vanes, pressure ratiomight have to be decreased to avoid stall. Our assumption of con-stant pressure ratio and lower air flow is somewhat intermediatebetween what is most desirable—constant air flow, higher pressureratio—and most unfavorable—lower air flow and lower pressureratio. The pressure ratio of the actual gas turbine will be deter-mined by its compressor map, which is kept strictly confidentialby manufacturers.



CO2Absorption~

~

ASU

~

Air

~H2 @60bar

~

HRSCHPDrum

SteamTurbine Deareator

SH

K.O.Drum

PurgeSaturator

GasTurbine

NGSaturator

K.O.Drum

K.O.Drum

Purge ICCompressor

NG Compressor

H2 Rec. Compr.

PSA

Stack

Air

LT W

GS

R

HT

WG

SR

Pre

refo

rmer

LPDrum

ATR

O2 ICCompr.

H18

H7

H6

H5

H4

H3

H8

H9-10

H1

H2

H12 H13

H14

H16

H17

NG

Steam for amine stripping

Pro

cess

ste

am


Vent

Fig. 4. Configuration of plants with ATR and CC for S/C = 1.0.

the performances quoted in[15] and the absorption proper-ties of amines reported in[25], we have assumed that theamine absorber removes 100% of the CO2 in the syngas (ap-parently, CO2 content in treated gas is lower than 100 ppm)and requires 1 MJ of low-temperature heat per kg of CO2captured. Total electricity consumption with a compressorwith four intercoolers is 440 kJel per kg of CO2.

The low-temperature heat needed by the amine stripperis provided by 2.1 bar steam bled from the ST. Alterna-tively, with FTR one could supply this heat by running thereformer at high S/C and condense the steam left in the re-formed syngas, as it is done in the studies carried out byFoster Wheeler[14] and Jacobs[15]. This configuration in-creases�H and the fraction of CO2 captured—due to higherS/C—but gives much smaller, possibly negative�E. Whichconfiguration is preferable will obviously depend on capitalcosts and the values of H2, electricity and CO2.

3. Calculation model and assumptions

Heat and material balances have been estimated by a com-puter code originally developed to assess the performancesof gas/steam cycles for power production with natural gasfuel [26–29]and later extended to handle gasification of coaland biomass[30], unconventional fuels[31], chemical re-

actors[2], fuel cells[32], steam-cooled GT expansion[33]and essentially all processes encountered in advanced powerplants.

The system of interest is defined as an ensemble of com-ponents, each belonging to one of 14 basic types: pump,compressor, turbine, heat exchanger, combustor, steam cy-cle, etc. Basic characteristics and mass/energy balances ofeach component are calculated sequentially and iterativelyuntil the conditions at all interconnections converge towardstable values. After converging, the code can carry out aSecond-Law analysis to trace irreversibilities and exergyflows.

The performances estimated by the model are for de-sign conditions, i.e. it is assumed that all plant componentshave been specifically designed to operate at the conditionsreached upon convergence of the heat/mass balances. Inother words, our results apply to “greenfield” construction.

Tables 4and 5 summarize the assumptions maintainedthroughout all calculations. They are representative of state-of-the-art reforming and power plant technology, but are notnecessarily optimal. Optimization of specific features likepressure drops, heat losses, temperature differences in heatexchangers or enthalpy differences in saturators require thedetailed design of each component, which is beyond thescope of this work.


Table 3Comparison between ISO GT performances quoted by General Electric (www.gepower.com, May 2003) and those predicted by our modelfor operation on NG and purge gas

Turbine GE 6FA GE 7FA

Fuel Natural gas Purge gas Natural gas Purge gas

Source GE Our estimate GE Our estimate

Input assumed to generate our estimatesAir flow (kg/s) 204.0 204.0 201.0 432.0 432.0 420.2Pressure ratio 15.7 15.7 15.7 15.5 15.5 15.5TIT (◦C) n.a. 1,316 1,316 n.a. 1,316 1,316�p compressor inlet (kPa) n.a. 0.0 1.0 n.a. 0.0 1.0�p turbine outlet (kPa) n.a. 0.0 3.0 n.a. 0.0 3.0Fuel flow (kg/s) n.a. 4.43 4.70 n.a. 9.56 11.89

Ar 0.0 0.0 n.a. 0.0 2.14CH4 94.47 48.36 94.47 37.50C2H6 3.12 0.0 3.12 0.0C3H8 0.50 0.0 0.50 0.0

Fuel C4H10(n) 0.50 0.0 0.50 0.0Composition CO n.a. 0.0 2.80 n.a. 0.0 2.77% volume CO2 0.40 0.01 0.40 0.01

H2 0.0 41.25 0.0 46.46H2O 0.0 6.31 0.0 8.98N2 1.01 1.29 1.01 2.16

Fuel LHV (MJ/kg) n.a. 48.43 45.60 n.a. 48.43 38.97OutputCompressor outletT (◦C) n.a. 409 409 n.a. 402 402Exhaust flow (kg/s) n.a. 208.4 205.7 n.a. 441.6 432.1TOT (◦C) 604 604 604.5 602.0 602.5 613.0Power output (MW) 75.9 75.7 75.7 171.7 171.6 169.8LHV efficiency (%) 34.8 34.9 35.3 36.2 36.7 36.6

NG composition is the same assumed for all hydrogen plants. The purge gas composition reported for the 6FA is the one calculated forFTR at 850◦C, S/C = 2.0; the composition reported for the 7FA is the one calculated for ATR at 950◦C, S/C = 1.0. n.a. = notavailable.The detailed input required to calculate each GT has been fine-tuned to best reproduce manufacturer’s data for operation on NG.

The composition of the syngas exiting the reactors is cal-culated as follows:

• ATR and LT WGSR generate mixtures at equilibrium;• the gas exiting the pre-reformer contains no hydrocar-

bon other than methane and meets the conditions for theequilibrium of the WGS reaction. The methane contentis set to the value giving a temperature 10◦C higher thanthat reached at equilibrium under adiabatic conditions.8

• FTR converts 88.5% of the methane that would be con-verted if the outlet gas were at equilibrium;

8 Due to economic considerations, the pre-reformer is typicallydesigned for a methane conversion lower than that reached atequilibrium; on the other hand the WGS reaction, being somewhatfaster, does get very close to equilibrium. The 10◦C differencebetween the actual temperature and the one that could be reachedat equilibrium is often called “chemical approach”.

• HT WGSR converts 97% of the CO that would be con-verted if the outlet gas were at equilibrium.

Given the characteristics of the catalysts[39] and the op-erating practices adopted in industrial plants,9 these as-sumptions reflect the average of well-maintained and well-operated plants.

3.1. Rationale of calculation scheme

Heat and mass balances are calculated through a complexiterative algorithm. One basic independent variable is alwaysthe amount of steammbl bled from the ST and added to thegas to be reformed; such flow is always non-zero because the

9 For example, in a FTR a loss of catalyst activity due tocatalyst poisoning affects mainly the first portion of the tubes, nearto the inlet. Some FTR configurations allow to compensate theseactivity loss by re-positioning the flames, while maintaining thesame plant capacity[18].

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Table 4Assumptions adopted for the reformer, the pre-reformer, WGSR,and the hydrogenation-desulfurization section

Hydrogenation & desulfurization

H2 in feed gas to hydrogenator (% vol) 10T of gas to sulfur absorber (◦C) 380Pressure drop (%) 1

Reformer ATR FTR

Syngas outlet pressure (bar) ∼ 68.5 25Syngas outlet temperature (◦C) 950 850Inlet T of gas to be reformed (◦C) 670 670Inlet T of O2/air (◦C) 110 300a

Reacting gas�p (%) 4 4Overpressure of O2/air (%/bar) 20 0.1Heat loss:�T , ◦C/% of heat transferred 4 1CH4 conversion, % to equilibrium 100 88.5T gas at furnace exit (◦C) — 1000%O2 at furnace exit — 1

Pre-reformer

Inlet T of gas to be pre-reformed (◦C) 620�T due to heat loss (◦C) 2�T chemical approach (◦C) 10

Water–gas shift reactors

Pressure drops (%) 2�T due to heat losses (◦C) Adiabatic LT 1

Adiabatic HT 2Heat loss: % of heat transf.—cooled LT 0.7CO conversion to equilibrium LT/HT, % 100/97

These same assumptions have been maintained also for reformingtemperature 1000◦C (ATR) and 880◦C (FTR).

aExcept FTR/SC plant with S/C = 3.07, where air pre-heattemperature is 130◦C.

saturator never provides enough water to reach the steam-to-carbon ratios considered here.

In FTRs with SC, all the purge gas is used in the furnaceandmbl is adjusted until the temperature of the combustiongases exiting the furnace is 1000◦C (seeTable 4); this alsogives S/C, which cannot be chosen at will.

In ATRs and FTRs with CC,mbl is varied to meet thespecified S/C. Then, in ATRs the specified reformer outlettemperature is met by adjusting the oxygen flow. Instead, inFTRs with CC the fraction of purge gas used in the furnaceburners and the fraction of NG used in the reformer (the restare used in the gas turbine) are adjusted to achieve: (i) tem-perature of combustion gases exiting the furnace 1000◦C;(ii) full GT power.

Table 5Assumptions maintained for all calculations

Air separation unit

ASU power consumption(kJel/kgPURE O2) 958.0

O2 purity (% vol) 95Pressure of O2delivered by ASU (bar) 1.01

Saturators

Pressure drop of gas stream (%) 2Pressure drop of water at nozzles (%) 10Max �T of water at inlet,◦C below boiling point 10Relative humidity of gas at exit (%) 100Minimum �h for mass transfer(kJ/kg(dry gas)) 25

Heat exchangers

Gas-side pressure drop—LP/IP/HP (%) 2/1/0.5Liquid-side pressure drop (%) 2–4Min �T , ◦C Gas–gas at LP/HP 40/30Gas–liquid 15Pinch point�T for evaporators (◦C) 10Heat losses, % of heat transferred 0.7

PSA

Separation efficiency FTR/ATR (%) 88/85Pressure of purge gas (bar) 1.30Pressure drop of permeated H2 (%) 2

Compressors

Polytropic efficiency of O2 compressor (%) 82# of intercoolers set to maintain O2 below 120◦C 6Polytropic efficiency of fuel compressors (%) 77# of intercoolers set to maintain fuels below 150◦COrganic× Electric efficiency of motor drives (%) 92Pressure drop in intercoolers (%) 1Pressure of fuel to GT combustor pressure 1.5

CO2 capture and compression

CO2 concentration in treated gas (ppmv) 100Gas pressure drop across absorber (%) 1.5Heat duty of amine reboiler(kJsteam/kgCO2

) 1000Total electricity consumption(kJel/kgCO2

) 440# of intercoolers in the compression train 4Heat rejected to compression work (%) 150Final delivery pressure (bar) 150

Boiler and steam cycle

SC/CC evaporation pressure HP level (bar) 110/130SC/CC SH and RH temperature (◦C) 540/565RH pressure (bar) 30Max gas temperature at SH/RH inlet (◦C) 1000a

Evaporating pressure LP level (bar) 12.5–23Deaerator pressure 1.40Condensing pressure (bar) 0.05

a1100◦C in ATR with SC and CO2 capture, to make avail- ableenough heat to SH steam up to 540◦C.


Table 6Comparison between the performances of the case study carried byFoster Wheeler[14] and those predicted by our computer code fora plant with the same configuration and the same design parameters

Test case Our data

AssumptionsNG input (MWLHV ) 369.7 369.7T reforming (◦C) 850 850p reforming (bar) 25.4 25.4S/C 3.00 3.00NG input to burners (%) 1.73 1.73H2 in hydrogenation feed (%Vol) 2.00 2.00T hydrogenation (◦C) 380 380T inlet pre-reformer (◦C) 620 620T inlet reformer (◦C) 670 670T inlet WGSR (◦C) ∼ 320 320p inlet H2 compressor (bar) 22.4 22.4PSA separation efficiency (%) 88.0 88.0p steam at turbine inlet (bar) 48 48T steam at turbine inlet (◦C) 380 380Condensation p (bar) 0.1 0.1

ResultsT outlet pre-reformer (◦C) ∼ 520 520CH4 at reformer outlet, (%Vol) 4.24 4.23T outlet WGSR (◦C) 392 388CO at WGSR outlet (%Vol) 1.98 1.98�H (%) 76.0 75.7Gross�E (%) 1.94 1.97Net �E (%) ∼ 0 0.43

3.2. Case study

To verify the agreement between the projections gener-ated by our model and our assumptions with those recentlyappeared in the literature, we have modeled one of the plantsconsidered by Foster Wheeler in a report prepared for theIEA [14]. The plant is significantly different from those de-scribed above: there is a single WGSR operating at mediumtemperature (320–392◦C) and no saturator; the SC featuresno reheat and modest design parameters (48 bar, 380◦C).The comparison shown inTable 6indicates excellent agree-ment between the heat/mass balances developed by FosterWheeler and those generated by our code.

4. Performance estimates

Heat and mass balances of the plants depicted inFigs.1–4 have been calculated for the conditions inTables 2–5,allowing for some variations of reforming temperature andS/C ratio.Table 7gives a summary of overall performances.Fuel inputs and outputs are always evaluated on a LHVbasis. Variations of S/C have been considered only for plants

with CC because:

• ATRs with SC are always inferior to ATRs with CC (seepar. 5.2);

• in order to feed the reformer with just the purge gas,FTRs with SC must operate at only one value of S/C;higher S/C require feeding the reformer with more NGthan the minimum required to control the flame, whilelower S/C make available extra purge gas for export—asituation beyond the scope of this analysis.

For FTR/SC plants, the S/C ratio giving the requiredmatch between purge gas production and reformer heat in-put can be (slightly) varied by varying the air pre-heat tem-perature, the pre-reformer inlet/outlet temperature, the re-former inlet/outlet temperature. For the plant with CO2 cap-ture (third leftmost column inTable 7) the air preheat tem-perature has been adjusted to give the same zero net poweroutput of the plant proposed by Foster Wheeler[14], re-sulting in a air pre-heat temperature of 130◦C (rather than300◦C, seeTable 4) and S/C = 3.07.

As reported inTable 2, when ATR is coupled with CC allthe purge gas goes to the gas turbine; consequently, the sizeof the reformer varies with S/C to match the requirementsof the GE 7FA assumed for our calculations.

For FTR with CC and S/C = 2.0, purge gas productionis enough to meet the requirements of both the reformerand the GE 6FA gas turbine; when S/C> 2.0, the reformerneeds more purge gas (and the purge gas LHV is lower), sothat some NG must be fed to the GT to run it at full power.This is shown inFigure 5. When S/C is larger than 2.75,purge gas production cannot even meet the requirements ofthe reformer; then, the GT is fed solely with NG and thereformer and the CC are coupled only through the steamsection.

To maintain the tightest integration between the FTR andthe CC, when S/C increases one should let NG input riseuntil purge gas production can match the gas turbine heat in-put or, for the same NG input, adopt a smaller GT. This hasnot been considered because very large NG inputs (above2000 MWLHV ) appear impractical and are likely to makeFTRs more expensive than ATRs; on the other hand, smallgas turbines would increase CC specific costs, reducing evenfurther the attractiveness of FTR/CC plants. These conjec-tures need to be verified by an adequate economic analysis.

4.1. Overall performance indicators

From the point of view of the First Law, the quality ofthe thermodynamic system is expressed by:

• �H = fraction of NG input exportable as pure hydrogenat 60 bar;

• �E = fraction of NG input exportable as electricity.

712S.Consonni,F.Viganò/InternationalJournalofHydroge

nEnergy30(2005)701–718

Table 7Overall performances of the configurations investigated in this paper

Power plant fate of CO2 Steam cycle Combined cycle

Vent Capture Vent Capture

ATRTref

◦C 950 — 950 — 950 950 950 950 950 950 950 1000S/C — 1.0 — 1.0 — 1.0 1.0 1.25 1.5 1.75 2.0 2.1 1.0NG input MWLHV 1236 — 1236 — 1337 1236 1385 1532 1675 1820 1876 1787E from GT MW — — — — 190.0 169.8 170.4 170.8 177.3 178.4 178.8 170.9E from ST MW 196.0 — 188.6 — 123.0 101.6 101.4 100.0 96.5 93.7 93.9 111.1�H2

%LHV 53.7 — 53.4 — 53.7 53.4 56.9 59.7 61.8 63.6 64.3 63.1�E %LHV 12.9 — 10.5 — 18.8 16.6 14.1 11.9 10.5 8.9 8.4 10.0�Ex % 64.4 — 64.3 — 70.0 70.2 71.3 72.1 72.9 73.3 73.4 73.6�∗

Ex % 64.4 — 61.9 — 70.0 67.7 68.6 69.3 70.0 70.3 70.4 70.8CO2 capture % 0.0 — 71.7 — 0.0 71.7 76.6 80.4 82.8 85.6 86.5 81.3H/FH % 69.8 — 67.2 — 80.9 79.2 78.6 77.9 77.8 77.1 77.0 78.4�H kg CO2 per 79.9 — 27.4 — 68.9 26.1 20.2 16.1 13.5 11.1 10.3 14.8

GJLHV . . . . . . . . . . . .

FTRTref

◦C 850 850 850 880 850 850 850 850 850 850 — 880S/C — 3.0 3.0 3.07 3.3 2.0 2.0 2.25 2.5 2.75 3.5 — 2.0NG input MWLHV 1195 370 1195 384 1945 1195 1195 1195 1195 1195 — 2384E from GT MW — — — — 80.3 75.7 76.5 75.3 75.1 74.6 — 75.8E from ST MW 59.7 7.2 41.6 15.5 154.1 69.9 67.8 65.7 63.6 54.8 — 122.2�H2

%LHV 75.7 76.0 78.0 73.3 66.0 66.0 66.1 66.3 66.4 66.5 — 72.5�E %LHV 3.2 0.0 0.0 0.0 10.0 8.7 8.7 8.5 8.4 7.8 — 4.8�Ex % 76.6 73.9 78.3 74.2 73.6 74.5 74.6 74.6 74.7 74.2 — 77.3�∗

Ex % 76.6 73.9 75.8 71.3 73.6 72.4 72.5 72.5 72.5 72.0 — 75.0CO2 capture % 0.0 0.0 73.6 85.0 0.0 61.8 62.1 62.3 62.5 62.8 — 66.9H/FH % 80.3 76.0 78.0 73.3 80.3 79.6 79.6 79.6 79.5 78.5 — 80.0�H kg CO2 per GJ 69.4 73.3 18.9 11.4 69.4 30.8 30.5 30.2 30.0 29.8 — 24.7

GJLHV . . . . . . . . . . . .

For comparison, shaded columns report the performances of the case study considered by Foster Wheeler [14]. H/FH is the ‘‘equivalent’’ efficiency of hydrogen production; �H is thespecific emission chargeable to H2 (see par. 5.6). �Ex accounts for the exergy of the pure, pressurized CO2 made available at the plant gate; instead, �∗

Ex disregards the exergy of CO2.The difference between �Ex and �∗

Ex is the CO2 exergy divided by the exergy of the NG input.


Vent CaptureVent Capt0

20

40

60

80

100

% o

f Nat

ural

Gas

Inpu

t

To FTR Burners To Gas TurbineTo ReformerCC

S/C

3.0 .0 2.25 2.52.0 2 .52.75 33.07

SC

Fig. 5. Breakdown of NG input of FTR plants. The fraction toburners shown for plants with SC (∼ 7.5 of LHV input to burners)is needed to control the flame. In the plants with CC, the flameis controlled via a slipstream bled from the purge gas compressor.The fraction to burners shown for the plant with CC and S/C=3.5is needed to complement purge gas production, which cannot meetthe requirements of the reformer.

while the Second-Law perspective is given by:

• �Ex = ratio between exergy out (hydrogen+electricity+CO2) and exergy in (NG);

• �∗Ex=ratio between the exergy of (hydrogen+ electricity)

and the exergy of the NG feedstock.

Note that exergy[34] includes the work generated by re-versibly mixing the output flows into the atmosphere, i.e.by expanding each species to its atmospheric partial pres-sure[38]. Such work�Exmix, which accounts for most ofthe exergy of CO2, cannot be recovered with the technolo-gies currently available. Its inclusion into the exergy flowsis necessary to close the Second-Law balance but overem-phasizes the relevance of output flows—particularly for theCO2 captured.Table 8gives LHV, exergy and�Exmix ofH2, NG and CO2, calculated from JANAF data[35] and,for ipercritical CO2, from [36].

Table 7points out the following:

• CO2 capture reduces�E by 2–3% points but has negli-gible impact on hydrogen production. Yet, in plants withFTR and SC the loss of electric output is partially com-pensated by higher hydrogen output: comparing the 1stand 3rd column from the left in the bottom (FTR) sec-tion of the table, one can see that CO2 capture reduces�E from 3.2% to zero, but�H2

increases from 75.7% to78%.

• Compared to FTRs, ATRs give lower�H but higher�E.When power is provided by a CC, the GT power outputof an ATR plant is a much larger fraction of NG input(and H2 output), thus justifying the different GT sizesadopted for the two reforming technologies.

• ATRs attain higher fractions of CO2 capture, particularlywhen coupled with CC and operated at high S/C; whenS/C�2, CO2 capture is larger than 85%. Instead, FTRscan achieve relatively low CO2 capture. Only by totallygiving up electricity export does one go over 70% with aplant based on a steam cycle (see FTR with S/C=3.07).As explained in par. 4.2, this follows from imposing thatthe burners are fed solely with purge gas, which increases�H but limits CO2 capture.

• The Second-Law efficiency�Ex slightly increases whenS/C and thus the fraction of CO2 captured increase. Evenif the contribution to�Ex of compressed CO2 may bequestionable, because it does not correspond to the workrecoverable with current technology, this trend points outthat CO2 capture does not entail specific thermodynamicdrawbacks. This is highlighted also by the trend of�∗

Ex,the Second-Law efficiency net of CO2 capture.

4.2. Electricity production and CO2 capture vs. H2 output

Fig. 6 reports �E vs �H, as well as the fraction ofCO2 removed. Points in the left part of the diagram (50%< �H < 65%) refer to ATR; the ones in the right part of thediagram (65%< �H < 80%) refer to FTR.

The figure shows that for ATR/CC plants the variation ofS/C introduces a trade-off among electric output—which de-creases when S/C increases—hydrogen production and CO2capture—which both increase when S/C increases. Instead,FTR/CC plants operate in a very narrow range of�E, �Hand CO2 capture independently of the S/C ratio.

ATRs with SC are definitely inferior to ATRs with CCbecause:

• for the same S/C, the plant with SC gives the same�Hand the same CO2 capture of the plant with CC, but�Eis much lower;

• if one lets S/C of the plant with CC vary, it is possibleto reach a condition where�E of the SC and the CC arethe same, but then the CC features higher�H and higherCO2 capture.

For the same S/C, the reforming section of the plants withSC and CC are identical: heat for the saturator and the re-former is taken solely from the reformed gas, which does notvary with the power plant; then, for the same amount (andcomposition) of purge gas and the same heat made availableto the power plant, the CC is obviously more efficient.

With FTR the CC does not show a clear advantage be-cause, despite its higher�E, it exhibits lower�H and lowercarbon capture. In fact, FTR/SC plants where the purge


Table 8Conditions, LHV and exergy of input and output flows

p (bar) T (◦C) LHV (MJ/kg) Ex (MJ/kg) �Exmix (MJ/kg)

H2 60 35 120.0 120.0 0.419NG 50 15 48.44 51.09 1.055CO2 150 35 0.0 0.658 0.433

Reference ambient conditions are 15◦C, 1 atm, 60% relative humidity.�Exmix is the work recoverable by reversibly mixing the combustionproducts into the atmosphere.

S/C=3.5

50

25

20

15

10

5

055 60 65 70 75 80

ηH=hydrogen output as % of NG LHV input

50

60

70

80

90

100

η E=

elec

tric

out

put a

s %

of N

G in

put

CO

2C

aptu

red

[%]

CO2captureyes no

E/NG input (left y-axis)

% CO2captured (right y-axis)

ATR FTR

S/C=2.1

S/C=2.0

CC, S/C=2.0Tref=880˚C

S/C=2.0

S/C=2.1

S/C=1.0

CCS/C=1.0

SCS/C=3.0

CCCCS/C=2.0

S/C=3.5

CC CC, S/C=1.0Tref=1000˚C

SCS/C=3.07

SCS/C=1.0

SC

S/C=1.5

S/C=1.5

Fig. 6.�E and fraction of CO2 removed as function of�H. The figure reports the same data inTable 7. Solid dots—to be read on the righty-axis–are given only for plants with CO2 capture (with venting, CO2 capture is always zero).

gas matches the requirements of the reformer (as we haveassumed here) reach�Ex ranging from 76% to 78% (seeTable 7)—an unsurpassed thermodynamic performance.

Plant performances are very sensitive to the reformingtemperature. For ATR with CC, increasing the reformingtemperature from 950 to 1000◦C generates about the sameeffect of doubling S/C (from 1 to 2) at constantTref=950◦C.

4.3. Second-law analysis

Fig. 7reports the breakdown of�Ex, as well as the exergylosses of four configurations representative of each of thereforming/power generation technologies considered here.

FTRs with SC stand out as the most efficient, whereas theopposite holds for ATRs with SC. The excellent performanceof FTR/SC follows from very low losses in the reformer,while the poor outcome of ATR/SC is mainly due to largeheat transfer losses. In a FTR, the reforming reaction isthermodynamically more efficient because it is fed by a low-

grade fuel like purge gas; instead, in a ATR the heat forsteam reforming is provided by burning (although partially)NG. As for heat transfer, the larger losses shown for ATRswith SC are due to the quench boiler placed downstream ofthe reformer and the boiler fed by the purge gas burners,which operate under large temperature differences.

When coupled to a CC, ATR combustion losses decreasebecause the purge gas burners are replaced by the GT com-bustor. On the other hand, at the high S/C needed to achievehigh CO2 capture, reforming losses are higher because moreNG must be burnt to feed the reaction; at the same time, com-bustion losses are lower because of lower purge gas flow. ForFTRs the CC ends up with lower�Ex because some extra-NG must be burnt into the gas turbine combustor, i.e. thereis a fundamental imbalance between the reforming sectionand the power section.

The sum of the losses due to combustion and chemi-cal reaction is always smaller for FTR because the heatgenerated by the combustion of purge gas is “recycled” to


Fig. 7. Breakdown of the contributions to�Ex (bottom diagram) andof the exergy losses of four relevant configurations (top diagram).

the reformer—a practice more efficient than combusting thepurge gas into the GT or—much worse—into the boiler ofATR plants.

4.4. Potential for CO2 capture

In FTRs with SC, CO2 capture is set by the S/C ratio,which determines the amount of NG converted in the re-former and the amount of CO converted in the shift reactors;unconverted NG and CO end up in the purge gas and gen-erate CO2 emissions from the FTR burners. NG conversioncan be increased by increasing the reforming temperatureand S/C, while CO conversion increases with just S/C. Thus,high Tref and high S/C favor CO2 capture; however, theyrequire large FTR heat input, which may decrease hydro-gen production. In the design considered by Foster Wheeler[14], whereTref = 880◦C and S/C = 3.3, the fraction ofH2 captured in the PSA is kept low to enrich the purge gaswith hydrogen and provide enough combustion power to theFTR burners; as a result, CO2 capture reaches 85% but H2output is only 73.3% (and�E = 0: see fourth column fromthe left in Table 7). Our design, where we have imposedthat the fraction of H2 removed in the PSA is always 88%,gives 73.6% CO2 capture, but H2 output is 78% (see thirdcolumn from the left inTable 7).

In FTRs with CC, CO2 capture is determined either byNG/CO conversion and by the amount of NG burnt in the

Fig. 8. Fate of carbon in NG input for plants with CO2 capture.

gas turbine combustor.Fig. 8 illustrates the fate of carbonin the NG input. Higher S/C decrease the emissions fromunconverted CH4 and CO but, as already shown inFig. 5,also require more NG for the GT. As a result, the fractionof CO2 captured is approximately constant, unless one in-creasesTref. or lets the NG input increase to fully load thegas turbine with just the purge gas.

ATRs feature the largest potential for CO2 capture be-cause, as shown inFig. 8, the higher NG/CO conversionachievable with high S/C is not counterbalanced by adverseeffects. Given that, as already mentioned in par. 5.2, the re-forming section of the plants with SC and CC are identical,the fraction of CO2 captured is independent of the powerplant technology. The maximum S/C achievable with theATR/CC configuration considered here approximately cor-responds to the case with S/C = 2.1 and 86.5% capture inTable 7andFigs. 6–8. Higher S/C cannot be sustained be-cause the increased dew point temperature of the reformedgas prevent the adoption of a LT WGSR; moreover, there isnot enough medium-temperature heat to superheat the steamgenerated in the quench boiler downstream of the ATR.

4.5. Reforming temperature

From the trend shown inTable 7andFig. 6, one wouldexpect that increasing both reforming temperature andsteam/carbon ratio would give high CO2 capture and high�Ex. However, aside from concerns for catalyst and tubelife, increasing the reforming temperature beyond the valuesconsidered here would disrupt the network for heat recoveryand prevent from realizing a good match between the flowsto be cooled (syngas and FTR furnace exhausts) and thoseto be heated (water, steam and mixture to be reformed).This is why we have not considered reforming temperatureshigher than 1000◦C for ATR and 880◦C for FTR.


Table 9Features of reference NG-fired CC considered as reference forelectricity production

CO2 venting CO2 capture

�E,ref (%LHV ) 56 51�E (kg/MWh) 360 40

4.6. Equivalent efficiency and emissions of hydrogenproduction

As it always happens in systems generating more than oneoutput, also for the co-production plants considered here itis difficult to define a single, satisfactory performance indi-cator. In fact,�Ex is helpful but somewhat misleading, be-cause the exergy of H2, electricity and compressed CO2 donot necessarily reflect their economic value, nor the techno-logical efforts required to generate them. In an attempt togive a better sense of how much co-production can actuallyincrease the efficiency of hydrogen production, we have in-troduced the ratio:

H/FH,

where H is hydrogen output and(FH) is the fuel“chargeable” to hydrogen. This ratio could be named“equivalent efficiency of hydrogen production”.FH is cal-culated by assuming that the total NG input is the sumof two terms, one attributable to electricity and the otherattributable to hydrogen:

Total NG input(MWLHV ) = FE + FH = E/�E,ref + FH,

whereFE is the fuel attributable to electricity, i.e. the fuel re-quired by a reference power plant with net efficiency�E,refgenerating the same net electric powerE generated by theco-production plant. In the same way, the CO2 chargeableto hydrogen (�H) is calculated by assuming that total CO2emissions are the sum of two terms, one attributable to elec-tricity and the other attributable to hydrogen:

Total CO2 emissions(kg/s) = �EE/3600+ �HH/1000,

where�E are the specific emissions (kg/MWh) of the refer-ence power plant,E is net electric output,�H are the spe-cific emissions chargeable to hydrogen (kg/GJLHV ) andH

is hydrogen output.To evaluateFH, �H andH/FH it is necessary to define

a reference power plant. We have considered a large NG-fired CC with the efficiency and specific CO2 emissionsreported inTable 9, taken from a study carried out by Chiesaand Consonni[6]. The values for CO2 capture refer to asystem where the fuel gases exiting the HRSG are cooled,treated in a chemical absorption amine system and partiallyrecirculated to the GT compressor.

The “equivalent efficiency of hydrogen production”(H/FH) and the specific emissions chargeable to hydrogen

(�H) are reported in the last two rows of the upper (ATR)and lower (FTR) sections ofTable 7. Adopting a CC ratherthan a SC has negligible effects on the equivalent effi-ciency of FTR—which stays close to 80%, also with CO2capture—while it significantly increases the efficiency ofATR—which increases by about 10% points. ATR+ CCoperated at high S/C reach the lowest�H–10–11 kg of CO2per GJLHV , i.e. less than 20% of natural gas—with H/FHin excess of 77%. This appears an excellent performancewhen considering that the most widespread technologynow used for hydrogen production—FTR+ SC with CO2venting—features an equivalent efficiency in the range of76–80% (see two leftmost columns inTable 7) with specificemissions around 70 kg of CO2 per GJLHV . In other words,a proper selection of existing, commercial technologiescombined with the optimization of the plant configurationcould reduce the CO2 emissions associated to hydrogen useby 6–7 times while maintaining about the same equivalentefficiency. This appealing outcome suggests the importanceof carrying out an economic analysis to verify the condi-tions for the competitiveness of the co-production plantsanalyzed here.

5. Conclusions

Coupling conventional NG reforming technologies withCC, rather than with SC, allows the export of significantamounts of electricity. For FTR, this comes together withlower hydrogen production, while for ATR the CC gives,with respect to steam cycles, higher electricity productionfor the same hydrogen production or, alternatively, higherhydrogen production for the same electricity production. Thesame applies to plants with CO2 capture.

As long as one is willing to pursue an optimum matchbetween the reforming section and the power plant section,the potential for CO2 capture of FTRs is more limited thanthat of ATRs. If the purge gas has to match the requirementsof the reformer (and of the GT, when present), the FTR S/Cratio is fully determined and the fraction of CO2 capturedreaches∼ 62%—with �H ∼ 66% and�E ∼ 8.5%—whenthe FTR is coupled with CC, going up to∼ 74%—with�H ∼ 78% and�E = 0—when the FTR is coupled to a SC.Higher CO2 capture can be reached by FTR/SC plants whenforegoing the optimum reformer/power plant match pursedhere, at the expense of significant reductions of hydrogenproduction.

ATRs can operate over a wide range of S/C ratios whileretaining full integration between the reforming section andthe power plant section. This flexibility allows ATRs to reachCO2 capture above 85%, with�H ∼ 64% and�E almost9%.

From a performance standpoint, the technologies ofchoice for the production of decarbonized hydrogen fromNG appear FTR with Steam Cycles or ATR with CombinedCycles. When operated at high steam-to-carbon ratios, the


latter reach the lowest CO2 emissions chargeable to hy-drogen (10–11 kg of CO2 per GJLHV , i.e. less than 20%of NG) with an equivalent efficiency of hydrogen produc-tion in excess of 77%. This is an excellent performancewhen considering that the most widespread technologynow used for hydrogen production—FTR/SC with CO2venting—features an equivalent efficiency in the range of76–80% with specific emissions around 70 kg of CO2 perGJLHV .

These considerations based only on performance esti-mates and thermodynamic arguments need to be verified bya comprehensive economic analysis.

Acknowledgements

The work described in this paper has been carried outat the Princeton Environmental Institute (PEI) within theframework of theCarbonMitigation Initiative(CMI), a jointproject of Princeton University, British Petroleum and theFord Motor Company. The authors gratefully acknowledgethe support provided through CMI, as well as the construc-tive suggestions and the supporting material provided byR.H. Socolow, R.H. Williams and T. Kreutz at PEI, P. Mid-dleton at BP, F. Saviano at Enitecnologie, D. Simbeck atSFA Pacific.

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